0444890378 fluid catalytic cracking

Studies in Surface Science and Catalysis 76

FLUID CATALYTIC CRACKING: SCIENCE AND TECHNOLOGY

This Page Intentionally Left Blank

Studies in Surface Science and Catalysis

Advisory Editors: B. Delmon and J.T. Yates

Vol. 76

.FLUID CATALYTIC CRACKING: SCIENCE AND TECHNOLOGY

Editors

John S. Magee

Catalytic Science Associates, 12205 Mount Albert Road, Ellicott City, Maryland 2 1042, U.S.A. (Formerly Director of Technology, Katalistiks International, A Unit of UOR Baltimore, Maryland, U.S.A.)

Maurice M. Mitchell, Jr.

Ohio University Southern Campus, 1804 Liberty Avenue, Ironton, Ohio 45638, U.S.A. (Formerly Vice President, Research and Development, Ashland Petroleum Company, Ashland, Kentucky, U.S.A.)

ELSEVIER Amsterdam - London - New Vork -Tokyo 1993

ELSEVIER SCIENCE PUBLISHERS B.V. Sara Burgerhartstraat 25 P.O. Box211, IOOOAEAmsterdam,The Netherlands

ISBN: 0-444-89037-8

t3 1993 Elsevier Science Publishers B.V. All rights reserved.

No part of this publication may be reproduced, stored in a retrieval system, or transmitted, in any form or by any means, electronic, mechanical, photocopying, recording or otherwise, without the prior written permission of the publisher, Elsevier Science Publishers B.V., Copyright & Permis- sions Department, P.O. Box 521,1000AM Amsterdam,The Netherlands.

Special regulations for readers in the U.S.A. - This publication has been registered with the Copyright Clearance Center Inc. (CCC), Salem, Massachusetts. Information can be obtained from the CCC about conditions under which photocopies of parts of this publication may be made in the U.S.A. All other copyright questions, including photocopying outside of the U.S.A., should be referred to the publisher.

No responsibility is assumed by the publisher for any injury and/or damage to persons or pro- pertyas a matter of products liability, negligenceorotherwise, or from any use or operation of any methods, products, instructions or ideas contained in the material herein.

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Printed in The Netherlands

V

Preface According to the table of contents, Fluid Catalytic Cracking: Science &

Technology, is concerned with fifteen different, though related, topics. While this is true, the reader is encouraged to consider the primary focus of the book as a whole to be on performance--performance of the catalyst, of its surface, of the FCC unit, of the feedstocks employed, of the analytical methods used to characterize the catalysts, and of environmentally directed regulations that govern the production of transportation fuels from petroleum.

The authors and the editors have tried to produce a volume that will fill the need for a comprehensive survey of this major field of petroleum processing while maintaining a high level of thoughtful brevity. Thoughtful brevity is one of those things that is difficult to define, but anyone who has sat through a thirty-minute sermon or read a 100 page final report knows what it is.

The subject matter of the book is intended to deal with several important performance issues:

What does the catalyst itself do as a function of its chemical and physical composition? How does molecular structure influence performance? How7 do metal contaminants influence performance? How does the FCC unit itself influence performance? How will environmental legislation influence the way the overall catalyst and cracking unit system must perform?

The emphasis on catalyst performance, particularly commercial performance, essentially dictated that the chapter authors be experienced industrial catalytic chemists and engineers. However, each author approached the task with a clear- cut obligation to connect the roots of the science of FCC catalysis with the technology.

In the case of FCC catalysis, the basic foundation was formed from virtually equal parts of pure science and practical technology, with an enormous sense of urgency caused by the transportation fuel needs of World War 11. Pure research in FCC catalysis was not neglected, and the pioneering work of Eugene Houdry, Paul Emmett, and Paul Weiss, to name but a few, broadened the base on which the industry was built. Evolutionary changes were followed by the revolution of zeolite-containing FCC catalysts. For this we all owe a debt to Charles Plank and Edward Rosinski.

The editors have chosen to document the revolution to date with fifteen chapters of FCC science and technology. As stated before, each author was charged with the task of documenting FCC catalyst performance from the standpoint of both the science and the technology involved in performance. We feel that performance is so important that we have included two chapters and part of a third on catalyst evaluation. Each offers a somewhat different viewpoint on

vi

how catalyst performance should be evaluated in the laboratory. In our opinion all three offer enough value to be considered equally by any serious worker striving to understand how a catalyst can be made to perform in the laboratory in a manner predictive of its commercial operation.

It is also very clear today that the environmental impact of the use of fossil fuels in the transportation sector can have a profound effect on FCC catalysis. Superficially, there seems little connection to the actual science of FCC catalysis, but, as an example, an entire supporting industry has grown based on the catalytic oxidation of CO to C02 and the catalytic elimination of oxides of sulfur and nitrogen from stack gases from the FCC unit regenerator.

Two chapters deal with the relationship of FCC catalysis and the real world of cars, planes, and plastics: Chapter 11 on FCC unit design and operational control, and Chapter 12 on the influences of the structural formula of the hydrocarbon being cracked and cracked-product molecular structure.

The science of FCC catalysis is amply treated in Chapters 2,3,5, and 6. Here the nature of catalytic sites, their influence on catalyst performance, the structure and complexity involved with the zeolite component of the catalyst, and the instrumental techniques involved in surface and structural analysis are described.

As is always the case, and justly so, we would add that the editors are pleased to acknowledge the many people who contributed time, effort, and above all, thought, to this project:

The authors--in reality the book is theirs.

The publishers-their approach to the hundreds of small and large problems in putting these pages between covers was always both professional and understanding.

The critics--our wives, Niki and Marilyn, who encouraged us to complete the project; Dr. G. M. Woltermann, who critically read the text; and Drs. Anonymous, the authors’ peers, who reviewed all of the chapters. These latter catalyst professionals know who they are and rightly deserve the high esteem in which they are held by J.S.M. and M.M.M., Jr.

John S. Magee Ellicott City, Maryland

Maurice M. Mitchell, Jr. Ashland, Kentucky

vii

Contents Preface List of Contributors

Chapter 1. Origin, development and scope of FCC catalysis A.A. Avidan

Chapter 2. The nature of active sites in zeolites: influence on catalyst performance A. Humphries, D.H. Harris, P. O'Connor

Chapter 3. Complexity in zeolite catalysts: aspects of the manipulation, characterization and evaluation of zeolite promoters for FCC D.E. W. Vaughan

Chapter 4. Commercial preparation and characterization of FCC catalysts G.M. Woltermann, 1.S. Magee, S.D. Griffrth

Chapter 5. Correlation between catalyst formulation and catalytic properties 1. Scherzer

Chapter 6. Instrumental methods of FCC catalyst characterization A. W. Peters

Chapter 7. Microactivity evaluation of FCC catalysts in the laboratory: principles, approaches, and applications E.L. Moorehead, 1.B. McLean, W.A. Cronkright

Chapter 8. Realistic assessment of FCC catalyst performance in the laboratory G. W. Young

Chapter 9. Residual feed cracking catalysts M.M. Mitchell, Ir., 1.F. Hoffman, H.F. Moore

Chapter 10. Metals passivation R.H. Nielsen, P.K. Doolin

Chapter 11. Unit design and operational control: impact on product yields and product quality L.L. Upson, C.L. Hemler, D.A. Lomas

V

ix

1

41

83

105

145

183

223

257

293

339

385

... vlll

Chapter 12. The effect of feedstock on yields and product quality W.S. Letzsch, A.G. Ashton

Chapter 13. Shape selectivity in catalytic cracking F.G. Dwyer, T.F. Degnan

441

499

Chapter 14. Additives for the catalytic removal of FCC unit flue gas pollutants 531 A. Bhattacharyya, J.S. Yo0

Chapter 15. Environmental considerations affecting FCC R.E. Evans, G.P. Quinn

563

Subject Index 587

ix

List of Contributors

A.G. ASHTON

A.A. AVIDAN

A. BHATTACHARYYA

W.A. CRONKRIGHT

T.F. DEGNAN

P.K. DOOLIN

F.G. DWYER

R.E. EVANS

S.D. GRIFFITH

D.H. HARRIS

C.L. HEMLER

BP International Limited Sunbury-on-Thames Middlesex, England TW 167LN

Mobil Research and Development Corp. P.O. Box 480 Paulsboro, New Jersey 08066

Amoco Chemical Company P.O. Box 3011 Naperville, Illinois 60566

The M W. Kellogg Company 16200 Park Row Houston, Texas 77084


Ashland Petroleum Company P.O. Box 391 Ashland, Kentucky 41114


Amoco Oil Company 200 E. Randolph Chicago, Illinois 60601

UOP 25 East Algonquin Road Des Plaines, Illinois 60017

Akzo Chemicals Inc. 3250 E. Washington Blvd. Los Angeles, California 90023


X

J.F. HOFFMAN

A. HUMPHRIES

W.S. LETZSCH

D.A. LOMAS

J.S. MAGEE

J.B. MCLEAN

M.M. MITCHELL, JR.

H.F. MOORE

E.L. MOOREHEAD

R.H. NIELSEN

P. O'CONNOR

Ashland Petroleum Company P.O. Box 391 Ashland, Kentucky 41 114

Akzo Chemicals Inc. 3250 E. Washington Blvd. Los Angeles, California 90023

Refining Process Services 4052 Firefly Way Ellicott City, Maryland 21042

UOP 25 East Algonquin Road Des Plaines, Lllinois 60017

Catalytic Science Associates 12205 Mount Albert Road Ellicott City, Maryland 21042

Engelhard Corporation 1800 St. James Place Houston, Texas 77056

Ohio University Southern Campus 1804 Liberty Avenue Ironton, Ohio 45638

Ashland Petroleum Company P.O. Box 391 Ashland, Kentucky 41114

The M.W. Kellogg Company 16200 Park Row Houston, Texas 77084

Ashland Petroleum Company P.O. Box 391 Ashland, Kentucky 41 114

Akzo Chemicals B.V. P.O. Box 975 3800 AZ Amersfoort, The Netherlands

xi

A.W. PETERS

G.P. QUINN

J. SCHERZER

L.L. UPSON

D.E.W. VAUGHAN

G.M. WOLTERMANN

J.S. YO0

G.W. YOUNG

W.R. Grace & Co.- Conn. 7379 Route 32 Columbia, Maryland 21044

Amoco Oil Company 200 E. Randolph Chicago, Illinois 60601

Unocal Science and Technology Division 376 South Valencia Avenue, P.O. Box 76 Brea, California 92621


Exxon Research and Engineering Co. Rt. 22 East Annandale, New Jersey 08801

The PQ Corporation 280 Cedar Grove Road Conshohocken, Pennsylvania 19428

Amoco Chemical Company P.O. Box 3011 Naperville, Illinois 60566

W.R Grace & Co.- Conn. 7379 Route 32 Columbia, Maryland 21044

J.S. Magee and M.M. Mitchell, Jr. Fluid Catalytic Cracking: Science and Technology Studies in Surface Science and Catalysis, Vol. 76 0 1993 Elsevier Science Publishers B.V. All rights reserved.

1

CHAPTER 1

ORIGIN, DEVELOPMENT AND SCOPE OF FCC CATALYSIS

AMOS A. AVIDAN

Mobil Research and Development Corporation Paulsboro Research Laboratory

P. 0. Box 480 Paulsboro, New Jersey 08066-0480 U.S.A.

1. INTRODUCTION

No other petroleum refining process, except for physical separation by distillation, has had a longer history, or more of an impact on the industry than cracking of heavy hydrocarbon molecules to lighter ones.

The increasing use of automobiles in the beginning of the "Petroleum Century" quickly consumed available "natural" gasoline, and to meet the needs, petroleum companies have been finding and producing more crude oil. But complex supply and distribution considerations, coupled with recurring "energy crises", have pushed refiners to upgrade less valuable petroleum products to gasoline. This need spurred William Burton, of crude-poor Standard Oil Company of Indiana, to commercialize the first thermal cracking process in 1913. Two other methods to upgrade heavy-ends to gasoline were developed later: catalytic cracking and hydrocracking.( 1)

Route Pressure First Commercialization Current Status

Thermal: Thermal Low 5 major processes from Coking,visbreaking Cracking 1913 to 1936

Catalytic: Catalytic Low Houdry process, 1936 FCC is a major Cracking (following unsuccessful refinery upgrading

McAfee process, 1915) process

Hydrocracking High Many attempts prior HDC complements and competes with

FCC to modern HDC, 1962

2

A major inefficiency in heavy-ends upgrading processes is the production of low- value coke. The three process routes have dealt differently with this problem. Hydrocracking suppresses coke formation by recirculating hydrogen at high pressure, while Houdry discovered that burning coke restores catalyst activity in catalytic cracking. Three major chemical reaction engineering solutions have been applied to implementing Houdry’s invention: fixed-bed (1936-1941), moving-bed (1941-1960), and fluid-bed (1942-today). The dates in parenthesis represent the heydays of each process.

1.1 What is Catalytic Cracking?

Eugene Houdry discovered in the 1920’s that heavy petroleum fractions crack over a solid catalyst, acid-treated natural clay, to lighter molecules. While clays and aluminas are still important ingredients of cracking catalyst, it was the introduction of zeolites by the Socony-Vacuum Oil Company in 1961 (2) which revolutionized catalytic cracking. The FCC process upgrades a variety of heavy feedstocks to lighter products (Figure 1).

Typical Yields

Figure 1. The FCC Process

The cycle oils can be used as heavy fuels or, upon further upgrading, as distillate fuels. FCC conversion is usually defined as the yield of hydrocarbon products other than cycle oils. FCC naphtha is used as a gasoline blending component upon further treating. The gaseous components are upgraded to gasoline blending components in a variety of light ends upgrading processes such as acid alkylation, etherification, and polymerization. Light olefins, such as propene and ethene, can also be used as petrochemical feedstocks.

3

Today’s FCC catalysts (Figure 2) have grown increasingly complex and they catalyze a variety of desired reactions. The main components (Y zeolite, and active aluminas) catalyze a complex set of cracking reactions, starting with carbenium ion chemistry.

Figure 2. FCC Catalyst Particle

FCC catalyst is usually a porous microsphere (about 50% pore volume) which is spray-dried to a powder with a particle size distribution of 10 to 120 microns, with a particle density of about 1,400 kg/m3. The heart of the cracking catalyst is the Y zeolite (Figure 3) available in many derivatives of varying physical and chemical properties.

4

24.73

24.68

24.65

24.70

24.54 * 24.60 + HY 24.50 +

24.53

24.35

RENH,Y 24.50 - 24.56

WP/, Na,O 13.0

2.5

2.5

0.3

2.5

0.3

2.5

0.3

0.03

0.3

Unit Cell Size (UCS)

Has a Total of 4 Complete, 6 Half and 8 One-Eighth Cages. It Contains 192 SI and At Atoms

Figure 3. Y Zeolite Derivatives

The main catalytic cracking reactions are: - - Isomerization of olefins. - - Hydrogen transfer. - - Alkylation and dealkslation.

and products in catalytic cracking. Weekman (3) has shown that for reactor design purposes, a simple three-lump model contains the salient features of the system:

Cracking of paraffins, naphthenes and side chains of aromatics.

Dehydrogenation of naphthenes and olefins.

Cyclization and condensation of olefins.

There are hundreds of reactions between thousands of components, intermediates

gasoline kl feed -

k3\ Coke + Gas

5

Cracking conversion, X follows an apparent second order:

where k, is related to catalyst activity and @is a catalyst deactivation function, which has usually been assumed to be the same for cracking as for coking.

A more detailed, "lumped" model (Figure 4) was developed by Jacob et a1.(4) This model lumps feed hydrocarbon molecules in broad categories, such as aromatic rings, yet it can accurately predict products distribution for a wide variety of feeds and operating conditions. More recently, new more complex models, tracking thousands of structurally-lumped components are being developed and used.

The cracking catalyst is only one of the many components in a modern FCC catalyst system. Other main ingredients catalyze reactions such as carbon monoxide and sulfur dioxide oxidation in the regenerator, contaminant metals passivation, and further cracking and isomerization over a smaller pore zeolite (Figure 5):

Heavy Naphthenes

Aromatics Substituents

Coke + C4 - lump

Lumped Schematic Scheme for Gas-Oil Cracking

Figure 4. Kinetics

6

Figure 5. FCC Catalyst - The Additive Approach

The control of such a complex mixture of catalytic and non-catalytic agents has become a sophisticated balancing act for the refiner and the catalyst manufacturer.

1.2 The Place of FCC in the Petroleum Refinery

The place and role of the FCC unit in the petroleum refinery has evolved over the past fifty years. Originally, in the 1940’s (Figure 6), the catalytic cracking unit, a Houdry fixed-bed, a moving-bed TCC, or one of the early FCC designs, was meant to complement the thermal cracker. Feed was mostly vaporized light gas oil. When mixed with alkylate and tetraethyl lead, light FCC naphtha could produce aviation gasoline of 100 research octane. After the war, demand for octane slackened, and the units were operated at lower severity for some time.

7

Crude- Oil

Straight Run Gasoline

I b

, Alkylation, Atmospheric - Polymerization - Distillation

Houdry, TCC or

FCC

Cracked Gasoline b

FCC Naphtha b

Figure 6. Catalytic Cracking Revolutionizes the Fuels Refinery (1940’s)

8

I ’ Reformate Pt

Reforming b

Figure 7. Zeolite Catalysts and Riser Cracking Revolutionize FCC (1970’s)

By the 1970’s (Figure 7)’ FCC’s replaced fixed- and moving-bed crackers. The introduction of zeolites (first commercialized in TCC) has had a major effect on FCC design. The dense-bed reactor was replaced by a short contact time riser. Platinum reforming was now well established and some refineries began hydrocracking light cycle oil to increase gasoline production. Environmental regulations, lead-phaseout, and oil supply shocks had a profound effect on the refining industry and on FCC’s. Again the unmatched flexibility of FCC design showed that old units can be modified and upgraded to meet new demands.

9

Crude Oil

- Reformate

Aikyiate

Continuous * Reforming I * --

Alkylation +

Atmospheric Distillation

Etherificatlon

V

FCC

b HDT HDC

Resid Upgrading 4 Butane

MTBE TAME

Light FCC Naphtha

Figure 8. The second clean air act requires gasoline reformulation; tighter emission limits (1990’s)

The 1990’s (Figure 8) show continued evolution. While the full impact of the 1990 Clean Air Act is yet to be determined, FCC is showing robust flexibility and even increased significance. It is a major producer of light olefins in the refinery and those are converted to alkylate and ethers, major reformulated gasoline blending components. Worldwide FCC capacity has been steadily increasing since 1942 (Figure 9). It is still expanding in the 19903, particularly in Pacific Rim countries and in Europe.

FCC naphtha today accounts for about a third of the gasoline pool, and the FCC complex produces over 40% of the gasoline when the olefins produced in FCC are converted to gasoline blending components. This fraction is expected to increase with the increasing demand for ethers and low aromatics gasoline blending components.

10

2,000

1,500

500

0 1940 1950 1960 1970 1980 1990

Figure 9. FCC’s Increasing Capacity shows Steady Increase

2. ORIGIN OF CATALYTIC CRACKING--HISTORICAL PERSPECTIVE

The first commercial trial of catalytic cracking came early in 1915. A. M. McAfee, of the Gulf Refining Companyydiscovered that aluminum chloride (a Friedel Crafts catalyst known since 1877) can catalytically crack heavy oils. While gasoline yield could be increased by 20-3096, the high cost of recovering the catalyst prevented the use of this process.

Thermal cracking was the heavy oil upgrading process of choice at that time, and several processes were developed following the commercialization of the Burton process (Figure 10). These processes, and the impact they had on the industry, are described by Enos.(5)

11

Figure 10: A battery of Burton-Clark cracking stills, each 8'ID by 30' long, processing a batch of 250 bbl.

The invention of catalytic cracking by Houdry, using solid "acid" catalysts, revolutionized petroleum refining (Figures 11,12). Eugene Houdry's long-time interest in racing cars had instilled in him the importance of gasoline quality. Silica-alumina catalysts were identified as effective in cracking gas oil to gasoline and hundreds of catalyst variations were tried at random. Motor performance was determined in Houdry's Bugatti racing car, driving up the same "calibrated" hill. Finally, Houdry settled for an acid-activated clay and established air regeneration to burn coke off the catalyst.

Houdry interested the Vacuum Oil Company (later Mobil), whose representatives visited the laboratory at Beauchamps, France in 1928,(1) and Vacuum's board authorized $100,000 for developing the process, if it could be demonstrated for 15 days. Since results were positive, Vacuum set up a 70 BPD industrial semi-works at the Paulsboro Refinery and in 1931 created with Houdry the Houdry Process Company. In 1933, Sun Oil Company joined the effort to develop catalytic cracking, and progress led to larger scaleup efforts. History was made on April 6,1936, when the first 2,000 BPD commercial cracking unit started up in Socony Vacuum's Paulsboro Refinery. In 1938, a 12,000 BPD unit at Sun's Marcus Hook refinery, equipped with motor-operated

12

valves and an automatic cycle timer was started up. Licensing to other companies got under way too, and by 1940 there were fourteen plants, with a capacity of 140,000 BPD.

Many technical innovations were quickly implemented after the first fixed-bed catalytic cracker. Socony Vacuum invented the molten salt-cooled cracker and together with Houdry, optimized the complex fixed-bed process. The next step was a continuous cracking process using the Thermofor kiln. The first 500 BPD semicommercial bucket elevator Thermofor Catalytic Cracking (TCC) unit was started up at the Paulsboro Refinery in 1941, and the first 10,000 BPD unit was built at Socony Vacuum’s affiliate, Magnolia Oil Company, Beaumont Refinery in 1943. By the end of the war, TCC capacity was nearly 300,000 BPD.

zones, rather than cyclically switching feed and regeneration air in a fixed-bed, was realized in the early bucket elevator TCC, the Houdriflow, the air-lift TCC, and the Fluid Catalytic Cracking unit (Figure 13).

The important idea of moving cracking catalyst between reaction and regeneration

Product - Steam 4

Air fl J F fl I‘

Turbocompressor

Charge Stock

Preheater

Fractionator

Figure 11. Houdry’s Catalytic Cracking Process Flow Diagram

Gas

Aviation Gasoline

Motor Naphtha

Recycle Stock

The continuous activity system (moving- or fluid-bed) solved one of the main shortcomings of the fixed bed-constantly changing catalyst activityj6) The instantaneous gasoline yield can be up to 15% higher than the integrated yield achieved in a fixed-bed reactor (Figure 13).

13

Figure 12. Houdry’s Catalytic Cracking Process Commercial Reactors

14

Regeneratlon

While the early bucket elevator TCC’s could only reach catalystloil rates of about one and one half, the air lift TCC’s licensed by Socony-Vacuum and Houdry in the early 1950’s could reach a catalystloil ratio of 4. Twenty-one Houdriflow units, with a capacity of 280,000 BPD, were licensed by 1956. The first air-lift TCC unit came on stream in the Beaumont, Texas refinery in October 1950, and by 1956 there were 54 Socony Vacuum and licensed TCC units.

C.(.lya 4

Regeneration R.genwstmd

c.t.Iyst

55

w

4s Ga.ollna Vd., x

40

35

30

Integrated Yield

- Amorphous Catalyst

60 70 80 90 100

Conversion, Vol. K

Figure 13. Integrated gasoline yield in a fixed-bed reactor vs instantaneous yield achievable in a constant activity reactor (based on FCC kinetic model and a schematic of a continuous activity system).

The moving-bed TCC air lift process is illustrated in Figures 14 and 15. The reactor vessel is suspended above one or more regenerators (or “kilns”). Regenerated catalyst flows out to a lift where it is conveyed pneumatically to a surge separator. The catalyst then flows by gravity to the reactor. Catalyst distribution and contact with the feed is carefully accomplished in the upper part of the vessel. Many collector pipes separate the cracked product from the catalyst, and the catalyst is again carefully distributed through a stripper. TCC units provide excellent contact for steam stripping, and resulting hydrogen-in-coke levels are as low as 5%, compared with the typical 7% in a fluid-bed stripper. The stripped catalyst is then burned in an efficient kiln.

engineering. It afforded excellent contact between solids and gas in the reactor, stripper and the kiln. Mechanical details in the design prevented excessive attrition of the bead catalyst in the lift system, and the vessels were filled with carefully designed internals to ensure optimal performance. The moving-bed solved the first problem of

Just like the fixed-bed reactor before it, the moving-bed system was a marvel of good

15

cracking elegantly: that of moving the catalyst between efficient contact zones. However, it missed the second chemical reaction engineering principle-the catalyst was still too large.

Alr

Blower

Product to Separation

Recycle

Figure 14. TCC Air-Lift cracking process

16

Liquid Feed

Catalyst

Catalyst Particles ,

Catalyst Bed

Products -1

Purge 1 r and Stripping Steam

I . . * ,..*. - 1 . . a ' * ... .:. ....... . . . . . . I . .

**: ; :.>'. ; : . . . . . . . . . . . . . * * . . . . . . . . . . . . . . . . . . . * : '

. . . 4 . . . . . . 1.. . . . * . . . . . . . . . :. ,* *. '

, . . a e . 1

t . . * . . . . . . . . . . ... .:*;:$.:: ::/;

Catalyst to Regenerator

Figure 15. TCC reactor design

Catalyst

Slide Valve

VaporFeed

- Liquid Feed Cone

Separation Pipes of / Cracked Products

From Catalyst

/ Catalyst Outlet

I - Catalyst Accumulator

The large particle size limited the regenerator temperature to below 650'C, requiring a large regenerator and large catalyst holdup. This is illustrated in Figure 16.

17

s5

6 -

0 I I 1

Yulnul

Effect of Temperature on Coke Burning Rate

Maximal Partlcle Size to Avoid Iptraparticie Temperature Excurslon

Figure 16. Maximal particle size to avoid intraparticle temperature excursion and residence time required to combust a catalyst with 1% coke (with 1% oxygen breakthroughM7)

The fluid-bed catalytic cracking process was originally designed to circumvent the Houdry patents. The use of fine powder was based apparently on a serendipitous discovery by R. K. Stratford in 1934. In 1938, Standard Oil Company of New Jersey (now Exxon), formed a consortium of eight companies: Kellogg, Indiana Standard, Anglo-Iranian, UOP, Texaco, Royal Dutch Shell, and I. G. Farben (which was dropped in 1940) in addition to Jersey, called Catalytic Research Associates, or CRA. CRA's purpose was to develop a catalytic cracking process which would operate outside Houdry's patents.

In 1934, R. K. Stratford of Jersey's Canadian affiliate (Imperial Oil) discovered that fine clay discarded from lube oil treating had catalytic effects in thermal cracking. Four thermal crackers were eventually revamped to "Suspensoid Cracking" by adding 2-10 pounds of powder per barrel of feed. The catalyst was used in a once-through mode, while Jersey decided not to follow this route, development in 1938 switched to a continuous 0.5 BPD pilot plant in which oil vapor conveyed powdered catalysts. Results from the cyclic fixed-bed 100 BPD pilot plant in Baton Rouge showed that best yields were obtained in the early moments of cracking-underscoring the incentive to develop a continuous cracking process. It was felt that conveying a fine powder would be easier than conveying pellets.

pilot unit using pipe coil reactors and a mechanical pump for circulating the catalyst. Later, the pipe coil reactor and regenerator were replaced with upflow beds and the pumps were replaced with standpipes of aerated catalyst and slide valves. A stable and smooth operation was achieved in August 1940, just 19 months before start-up of the commercial unit based on these same concepts and design. The first commercial FCC

By mid-1940, Jersey had demonstrated a powdered catalyst process in the 100 BPD

18

unit, the Model I, was an upflow circulating system (Figures 17 and 18). It was started up in May 1942. By July 1942, its feed rate was already over 17,000 BPD.

Flue Gas Vent

r_ Spent

catalyst

/Hopper Product

Fractlonator - Gasoline

Bottoms

Air Blower

Figure 17. Model I FCC Schematic

19

Figure 18. The three catalytic crackers of the Baton Rouge refinery. The Model I unit is on the left; two Model I1 units are on the right.

The Model I unit had several limitations, and only three were constructed in the 1940’s. However, it also had several features which seem advanced even today. Some of these were dropped in subsequent models, but many were later resurrected. These features are listed in Table 1:

20

Table 1 Evolution of FCC Features

Model I Feature Later Models Current Interest

Catalyst coolers

upnow regenerator and catalyst circulation

Upflow reactor

Quick separation

Flue gas precipitator

Developed heat balanced" concept - all heat produced in regenerator is supplied to cracking.

To lower regenerator pressure drop, bed height was minimized in downflow.

Changed to downflow (Model 11) to lower overall height and reduce cyclone loading.

Large volume in reactor freeboard has increased thermal cracking.

Cyclones only.

To crack resid feeds, catalyst coolers are added to regenerator.

High efficiency regenerators incorporate upflow principles.

All current units have upflow risers; better performance in upflow, with higher velocity.

Closed cyclones eliminate thermal cracking.

Environmental concerns force better dust removal.

High catalyst losses (0.5 Ib/bbl initially, and 0.25 Ib/bbl later) in the Model I favored the "downflow" Model 11. Further evolution and development by Exxon, Shell, Texaco, UOP, Kellogg and others was rapid in the 1940's and early 1950's (Figure 19).

21

Original Downflow (Model I I )

75 1

60

45

Elevation, m

30

15

0

UOP Stacked Design

Kellogg Orthoflow B

S.O.D. Model IV

Figure 19. Evolution of a 20,000-BPD FCCU, 1943-1952

3. FCC PROCESS DEVELOPMENT IN THE PAST 50 YEARS

The FCC process has continuously combined catalyst, hardware, and process technology to produce optimal results. The main elements of FCC process technology include: operating strategy, steady state optimization, reliability, process control, environmental control and integration with other refinery units and refinery energy balance. The FCC process is a major factor in refinery profitability, as it is usually the main upgrading process. FCC has been the most profitable and flexible refining process for nearly half a century because of its ability to meet changing demands.

Major process development highlights in the past 50 years have included: . Gradual improvements in hardware reliability, in such areas as rotating equipment, valves and refractory. It is not uncommon today to have FCC units complete three-to-four-year runs between major turnarounds with nearly 100 percent stream factor. Choosing an optimal operating strategy, and steady state optimization based on advanced process models, has evolved over the years. There are many "Knobs" available to the FCC operator (such as catalyst choice, equilibrium catalyst activity, riser top temperature, preheat, etc.), and even to the most experienced operators it is not always clear which set of conditions, within unit constraints, will yield the optimal results.

Other key process areas are listed in Table 2.

22

Table 2 Development of Key FCC Process and Hardware

FCC Reactor . . . Multiple feed injection, quench. .

More efficient stripping.

Short contact time riser cracking. Feed distribution and atomization, use of dispersion steam.

Quick product quench (closed cyclones).

FCC Regenerator . . Lower NOx emissions.

Lower excess oxygen. Improved air grid designs. .

Improved regeneration efficiency, lower inventory.

Power recovery from flue gas.

Other Areas - Improved, high flux standpipes. . High-efficiency cyclone separators. . Third- and fourth-stage particulate capture systems.

FCC evolution at each operator, or technology licensor, has taken its own path, with many common areas, but also many individual traits.

Shell Oil Company contributed in several ways to FCC development. One of its early Model 111 units was the first to use microspheroidal (MS) catalyst. Use of low- attrition MS catalyst permitted elimination of Cottrell precipitation. This also enabled the elimination of waste-heat boilers, which at the time had maintenance problems. One of Shell’s innovations was the Anacortes unit, completed in 1956, which featured a short-residence time FCC riser reactor followed by a second-stage conventional dense- bed cracker. Other Shell FCC developments include use of expander turbines for power recovery and catalyst fines recycle for particle size distribution control. A review of FCC development at Shell, leading to the most recent Shell Resid FCC designs is shown in Figure 2047)

23

1980's

TO POWER

. ... . . . .. . .;. ;.

a ' :' . .. . . . ..' :.. . .

PRODUCT

I-RESID

1990's

TO POWER RECOVERY u

SWIRL-TUBE SEPARATOR

AIR-

PRODUCT = J +

+STEAM

t - - R E S I D

Figure 20. Shell FCCU Designs

Other CRA members continued developing their own versions of FCC technology following World War 11. After licensing and building many of the early models, UOP introduced its "stacked" unit in 1947 with the 3,000 BPD Aurora unit in Detroit. Thirty stacked UOP units, typically in the 4,000 to 10,000 BPD range, were sold mostly to independent refiners in the early 1950's. UOP later built large side-by-side riser FCC units, and together with Mobil, developed the It high-efficiency" riser regenerator. Together with Ashland, UOP also developed a r a id FCC unit, the RCC (Figure 21).

24

Figure 21. UOP FCCU Designs

J I Catalystend Spent Catalyst Oil Feed to Regenerator

(a) Schematic of UOP FCC Reactor c. 1960-before changeover to all riser cracking.

u

(b) Mobil-UOP High-Efficiency Regenerator FCCU

v (c) Ashland-UOP RCC Unit

25

M. W. Kellogg designed 47 side-by-side FCC units from 1944 to 1955. These units ranged in capacity from 4,000 to 60,000 BPD. In 1951, Kellogg began offering its Orthoflow design (Orthoflow "A"), with the reaction and regeneration zones superposed within a single vessel. Vertical catalyst standpipes employed plug valves to control flow. Six "A" units were built until the introduction of the "Bl' model in 1958. Regenerator and reactor relative positions were switched. The regenerator was placed above the reactor in "B". Fifteen "B" units were built until the introduction of Orthoflow "C" in the late 1960's, following widespread use of zeolite cracking catalysts. Kellogg, like other FCC vendors, switched to riser cracking to take advantage of those high-activity, coke-selective catalysts. Ten "C" units were built until the introduction of "F" units in 1977. The Orthoflow "F" (Figure 22) was the result of evolutionary FCC development efforts at Kellogg. It also includes the "Ultra Cat Regeneration" technology developed by Amoco.(8)

/-- Product to Separation

Upper Cyclones

Riser 90" Turn

Riser Reactor

Two-Stage Regenerator

Cyclones

Feed Injection

Riser Expansion Joint

Disengager

Riser Cyclones

External Plenum

Flue Gas

Stripper

Regenerator

Spent Catalyst Standpipe

Figure 22. Kellogg Orthflow Converter FCC

Kellogg has also been a pioneer in resid FCC, with the first HOC (heavy oil cracker) commercialized in 1961.

26

To convert resid, large amounts of heat are needed to be removed from the regenerator via bed coils and/or a catalyst cooler (Figure 23).

Boiler Feedwater

Steam Makeup

Blowd

Regenerator Bed Coils

Bed Level B Figure 23. Heat Removal from RFCC Regenerator

Today’s catalyst coolers are dense-phase coolers, in contrast with the dilute-phase coolers of the Model I. These, and other key hardware areas of a modern RFCC, are illustrated in Figure 24.

27

\ 1 Short Contact Time Riser

Atomizing Feed Nozzles

Catalyst Pre-Acceleration

Closed Cyclone System

{ 11 t-, Efficient Stripper Design

Regenerator Catalyst Distributor Design

Heat Removal (Bed Coils & External Cooler)

Air Grid Design

Figure 24. Key hardware areas of a modern RFCC unit

More recently, Kellogg and Mobil joined forces in FCC technology licensing, improving the gas oil and heavy oil cracker designs further with improved atomizing feed nozzles and closed cyclones.(9) Closed cyclones cut down the products residence time in the reactor from over 30 seconds to only a few seconds. This eliminates most of post-riser thermal cracking, decreases dry gas make by up to 40%, and increases gasoline distilled yields by about 2.5 vol %. In addition, butadiene yields decreased by up to 50% (Figure 25).

28

1150

1100

1050 360

280

200

I Open Cyclones I Closed Cyclones I

0 ~~

1 2

Time, Hours

10000

Butadiene, PPm in 7500 Cu Stream

5000

12500

0 0 Open Cyclones 0 Closed Cyclones 10000 -

Butadiene, PPmin 7500 - Cu Stream

5000 - I I I

Riser Top Temperature, O F

2500970 980 990 1000 1010

12500 I

- 0 Closed Cyclones 0 0 Open Cyclones

- -

I I I

Riser Top Temperature, O F

2500g!o 980 990 1000 1010

Figure 25. Commercial Experience with Closed Cyclones

Exxon also brought out new FCC designs, moving through the Model 111 units (1945- 1952), to the side-by-side Model IV units. This design was developed to compete with the smaller units, mainly in range of capacities of 5 to 20 TBD being offered by other licensors. The first Model IV unit 15 TBD, went on-stream in Destrahan, Louisiana, in November 1952. The advantages of this design were a reduction in height, an improved catalyst transfer system, a more stable operation, improved process control and less erosion in catalyst lines. The Model IV design proved to be suitable for large as well as small units; Model IV units with capacities as high as 75 TBD are still being operated. Forty Model JY units were build by the early 1970’s when Exxon’s Flexicracker design based on a transfer line reactor became available.

29

Other successful commercial FCC designs, which are available for license, include the resid FCC units licensed by Stone & Webster and IFP.

4. SCOPE OF FCC CATALYSIS

Cracking catalyst systems have also been evolving continuously for over 50 years, and catalysts are still at the heart of the FCC process. Worldwide FCC catalyst production is over 1100 tons per day, and sales are over 500 million dollars per year. The major vendors are listed in Table 3. In recent years, overcapacity has kept prices low, competition stiff, and resulted in several consolidations. Should FCC catalysts be treated as specialty chemicals, as they have been mostly, or as a commodity? The long- term interest of refiners is clearly with the former. Catalyst cost is usually a small fraction of the uplift in FCC, so refiners usually look for the most cost-effective catalyst for their application. Today, refiners have a wide range of choice of quality catalysts.

4.1 Recent Developments

Cracking catalysts have undergone many evolutionary and revolutionary changes. Milestones are listed in Table 4. Today's FCC catalyst system is a complex mixture of functional components (Figure 5). The main component is the FCC catalyst itself, containing Y zeolite, which is providing the primary cracking function. Other components currently include:

Combustion Drornoter-used to control CO emissions in FCCUs without an external CO boiler and for regenerator temperature control. ZSM-5 additive-increases octane and light olefin yields. There are currently over 60 commercial FCC units using ZSM-5; ZSMJ use is expected to increase considerably as the push to produce more light olefins continues. These light olefins are used for producing alkylate and ethers such as MTBE (major components in "reformulated," or "clean" gasolines). A recent development in ZSMJ technology is the use of "high-activity" additives, which have cut makeup by at least a factor of two. Another development is the "high-selectivity" ZSMJ additive, which cracks less gasoline and produces less LPG, to achieve the same increase in octane. The increase in octane is achieved more by improved isomerization-hence, motor octane increases more with the selective than with the high-activity ZSM-5. Desulfurization additives--promote oxidation of SO, to SO, in the regenerator, and adsorption of SO, onto the additive, which is then transferred to the riser. SO, is reduced in the riser and catalyst stripper to H,S, which is later recovered in the gas plant. Sulfur oxide emissions are reduced by up to 70% in complete, and up to 50% in partial CO combustion, depending on sulfur levels.

30

Table 3 FCC Catalyst Manufacture

Estimated Capacity (ton/day)

Davison (Grace) (Lake Charles, Los Angeles, Baltimore, Cincinnati, Canada, Germany)

& (Los Angeles, Texas, Brazil, The Netherlands)

Engel hard (Georgia, The Netherlands)

Katalistiks* (Georgia, The Netherlands)

520

310

225

215

Others (Crosfield, CCIC, etc.) 180

I' Namepla te" Capacity Actual Capacity

- 1750 1450

Total Production

* Katalistiks terminated FCC catalyst production in June 1992.

31

Table 4 Historical Milestones in Cracking Catalyst Developments

1942 1948 1955 1961 1964 1974 1975 1980 1984 1985 1986

Natural Clay, Synthetic Low Alumina Catalyst Microspheroidal Catalyst (Low Alumina) High Alumina Synthetic Catalyst D5 Zeolite TCC Bead Catalyst Spray-Dried Fluid X and Y Zeolites Pt CO Combustion Promoter Ni Passivation Additive Coke Selective Re-H-Y, USY Catalysts ZSM-5 Octane Additive SOX Transfer Additives Y Zeolite Improvements for Low Coke Selectivity, Higher Octane (low non-framework alumina, small and "perfect" crystals, chemical dealumination, etc.)

Table 5 Improvements in FCC Catalysts

1950 1970 1990

Zeolite Content, % Particle Density, g/cc Relative Attrition Index

0 10 up to 40 0.9 1.0 1.4 20 5 1

With properly functioning dust recovery systems (cyclones, third-stage separators,

FCC catalysts are supplied in various grades of particle sizes and attrition resistance. electrostatic precipitators) dust emissions from FCC units can be very low.

At the refiner's choice, post-calcination can reduce loss-on-ignition and improve attrition resistance. The coarse grade does not flow as well as normal fines content material (about 25% less than 40p), and can limit catalyst circulation in poorly- designed standpipes. Separate high-density fines ("fluidization" additives) are available.

The Y zeolite has been the main ingredient in FCC catalysts for 25 years. USY- containing catalyst usage in the U.S.A. has been increasing steadily from the mid-l980s, and currently amounts to about 70% of total use. This trend will probably accelerate abroad as well. The Y zeolite superior selectivity and activity over larger pore

32

amorphous alumina, or silica-alumina has been fine tuned over the years. One big question in FCC catalysts for the future is: Will there be another zeotite, molecular sieve, or ang other microporous structure superior to the Y zeolite? To date, this question has not been answered, but there has been considerable research and speculation. With the increasing importance of resid cracking, larger pore materials have been proposed as cracking catalysts:

Table 6 Current and Future Cracking Catalysts

Amorphous Pillared Y-Zeolite Alumina VPI-5 Clay

Avg. Pore Opening, A SUAl Ratio Range Surface Area, mZ/g Severely Steamed

Relative Cracking Surface Area, mz/g

Activity

7 10-200+ 12 9-15 2.5-10 Wide Range 0 Wide

900 250-400 N.A. 200-400

300 0-150 0-20 0-20

10-1.000 1 .o -0 21

None of the materials with larger pore openings than Y zeolite has so far shown significant cracking activity, or good hydrothermal stability.

The major problems in resid processing are handling high CCR and metal levels in the feed. The CCR contributes to coke yield, hence the need for better coke selectivity. The high metals levels (nickel and vanadium) contribute to catalyst deactivation and to high coke make and require excellent metal tolerance and high makeup rates. In addition, metals passivation is now handled by:

V - Vanadium traps such as magnesium oxide and titanates. Ni - Antimony or Bismuth passivation.

Metals tolerance is likely to continue to be one of the major growth areas in catalyst research and development. However, it is possible that a major breakthrough in catalysis will show that unique structures with pore openings larger than Y zeolite will offer superior metal tolerance and selectivity over Y zeolite.

Other future trends are likely to be a continuation of recent ones: Zeolite content is likely to keep increasing, particularly with emphasis on resid processing. Today’s premium FCC catalysts, with about 40% zeolite, can be extrapolated to about 50% zeolite before significant degradation in zeolite availability and physical properties takes place. Future developments may allow even higher zeolite levels. One additional factor supporting the need for higher zeolite catalysts is the forecasted increase in the use of additives which dilute the concentration of Y zeolite in the FCC catalyst inventory. These additives will include further use of

0

33

SOX transfer agents, better vanadium and nickel traps, and potentially new additives. Z S M J use is likely to increase dramatically if proposed regulations mandating reformulated fuels (such as 2.7% minimal oxygen level in gasoline, maximal aromatics levels, etc.) take place. The FCC complex is currently the source of some of the "dirtiest" fuel components (such as LCO and FCC heavy naphtha) and some of the "cleanest"--light olefins which can be converted to alkylate, MTBE, TAME, synthetic diesel, etc. (Figure 24). Matrix technology is likely to keep improving in metals tolerance, more selective bottoms cracking, and physical properties. Catalyst demetalation technology may become more widespread in commercial use. Equilibrium catalyst metals can be removed by a chemical washing process, such as DEMETTM. A demetalation process can restore activity and selectivity and save on catalyst makeup in a resid FCC.

-

- .

5. THE FCC PROCESS INTO THE 21ST CENTURY

It is tempting to assume that there a re no remaining quantum improvements in catalytic cracking. This was probably the opinion in the decade before zeolites came to the scene. A 1965 plot by Mobil researchers Farber, Payne, and Sailor (Figure 26), based on TCC zeolite catalyst, shows how far we are still from what might be considered ultimate yields in catalytic cracking. Running at low conversion per pass, with intermediate gasoline removal, results in much higher gasoline selectivity than single-pass conversion. The cumulative advantage increases with conversion: 13% more gasoline at 60% conversion, and 24% more a t 80% conversion. The results of multipass runs show much lower coke and light gas yields than single-pass operation. Coke can be produced from condensed and polymerized hydrocarbons formed from reactive intermediates. Coke and light gases can result from cracking of gasoline product. While using zeolite catalysts reduces coke formation due to steric hindrance compared with amorphous catalysts, it does not eliminate it. There remains considerable potential for increasing FCC gasoline yields (Figure 27).

We believe that FCC technology, operating conditions, and apparatus are going to be as similar and as different 50 years from now as today's technology is compared to the Model I. Many basic principles will probably be retained, but the evolutionary process which has been responsible for past progress will continue. Yet, there is room for real breakthroughs.

34

Incremental Q'r and Q-

Free Gasoline Volume % of

Increment Converted

100 80

50

30

20

10

t 4

@

-Runs,

u - u

Varlabk LnSV I - - Single Pasa Runa, 8 LHSV I

I I I I I 0 20 40 60 80 100

Conversion Level at Which Increment is Converted, Volume % Charge

Durabead 5 Wide Cut Mid-Continent Gas Oil

10

llncremental Coke Weight

% ot Increment Converted

Multiple Pam Runs, Varlable LHSV Single Pass Runs, 8 LHSV #

4 #

4 0

0 #

#

I I 1'64/ 0 20 40 60 80 100

Figure 26. "Ultimate" yield in cracking over zeolite satalyst, comparison of once- through to multipass. Data of Farber, Payne and Sailor.

35

100

80 Gasoline + Light Oleflns Yield, wt. % 60

40

FCC Improvements and Potential

-

- - c

I I I

Figure 27. Potential for FCC Yields

Future developments in FCC hardware are likely to be a continuation of recent trends, particularly with increased processing of resid. Hardware such as feed nozzles and catalyst coolers will be improved. Contact times in FCC risers may be reduced even further than today’s practice.

36

Table 7 Residence Time Evolution

Today's Short Dense Bed Contact Time Cracking Riser Cracking

Catalyst Residence Time, s

Vapor Residence Time, s

Temperature, ' C

30-120 5-15

10-60 1.5-5

480-500 520-550

Another area of increased significance in the need to lower regenerator emissions further. Significant advances have been made since the 19603, but further reductions are mandated.

Resid upgrading is becoming more and more attractive particularly in Pacific Rim countries and land-locked refineries. It may be possible that most new FCC units will be RFCC's, and many existing ones will be converted to RFCC's (Figure 28). Integration of RFCC with feed hydrotreating will be mandated to process feeds with CCR levels higher than 7 and metal levels higher than 30 (Figure 29).(7)

FCC's place in fuels reformulation is likely to change FCC's role in the U.S.A. by the mid-1990's and possibly in other locations afterwards. FCC is the key to reformulating fuels in the refinery (Figure 29) as it produces some "dirty" and some of the "cleanest" fuel components.

37

Worldwide Grassroots Resid Crackers

Projected 700 I

Capacity, TBD

525

350

175

0 (1961) 1976 1979 1982 1985 1988 1991 1994

Year

Figure 28. Worldwide Raid FCC Capacity

1 oo+ 100 90 80 70 60 50 45 40 35 30 25 20 15 10 5 0

0 ; 5 10 15 Llght, Pariffinlc Long Residues

Far East e.g. North Sea % Wt. CCR in Feed

20

Figure 29. Raid (370'F C+) Properties in Relation to FCC Processability

38

c2- I r-F( Fuel Gas,

Alkylate MTBE

c, 's c, 's

r - - - i FCC 4 HDT Gas

I I Plant C,'s - Tame 9 Tame

1 9 FCC Naphtha

LCO

MOG - Oleflnlc Gasollne - MOGD - Low Aromatics, High Cetane Clean Diesel -

Figure 30. "Clean" and "Dirty" Fuels from the FCC Complex

I G E

Thermal Crackers

Unsat. C4- Sat. Gas Gas c4- - * Plant Plant

(FCC) E Olefins

E EG,

c4- -L

I ' Raffinate

E - Extraction -

Sat. 1 rs Naphtha t

E Gas Plant

.-)

I? I Refinery Petrochemicals

Ethylene

Propylene

Butylenes

Benzene

Toluene

Xylenes

Figure 31. Current Integration of Refinery and Petrochemicals

39

One current trend which is likely to accelerate, and eventually may be even dominate catalytic cracking is FCC integration with petrochemical production. Some FCC units today are already integrated with petrochemical plants in various ways, illustrated in Figure 31. Ethylene is extracted from FCC fuel gas, propylene is sold as a petrochemical (typically at a much higher price than if it were converted to gasoline), and H, is supplied back to the refinery. The future refinery is likely to be more of a specialty chemical producer than today-centered around the FCC complex. Through innovations in hardware and catalysts, FCC may also replace thermal cracking as the major route to producing light olefins, particularly from heavy feedstocks.

state control scheme is constantly being improved with safety and reliability features, and better steady state optimization. This is currently done off-line in most cases. The future may bring more sophisticated on-line steady state optimization and the predictive advanced control which will anticipate FCC feedstock changes. FCC units which may benefit from such control are those where significant objectives, rates, or feed quality changes take place more than once a week.

Many FCC units are already benefiting from the use of state-of-the-art computer technology--they are likely to keep up with rapid advances in computer-related technologies. Far from being a mature, "low-tech" technology, FCC technology, catalysts, hardware and process will continue to lead petroleum refining in innovation, safety and reliability, environmental impact, and last, but not least, profitability.

Another area of great interest is advanced FCC complex control. The current steady

6. REFERENCES

1

2 3 4

5 6 7

8

9

10

A.A. Avidan, M. Edwards, and H. Owen, Innovative Improvements Highlight FCC's Past and Future, Oil & Gas Journal, January 8,1990. CJ. Plank, and EJ. Rosinski, Chem. Eng. Progr. Symp. Series, 73 (63) (1967). V.W. Weekman, Jr., Ind. Eng. Chem. Process Des. Dev. 8(3) 385-391 (1969). S.M. Jacob, B. Gross, S.E. Voltz and V.W. Weekman, Jr., AIChE J., 22,701-713 (1976). J.L. Enos, Petroleum Progress and Profits, MIT Press, Cambridge, Mass. (1961). A.A. Avidan, and R. Shinnar, Ind. Eng. Chem. Res. 29,931-942 (1990). F.H.H. Khouw, G.V. Tonks, K.W. Szetch, A.C.C. van Els, A. van Hattem, The Shell RFCC Process, Akm Catalyst Symposium, Scheveningen, The Netherlands, June 1991. R.E. Wrench, R.E. Wilson, A.K. Logwinnk, and H.D.S. Kendrick, Fifty Years of Catalytic Cracking, an M. W. Kellogg Publication (1986). A.A. Avidan, F.J. Krambeck, H. Owen, and P.H. Schipper, FCC Closed-Cyclone System Eliminates Post-Riser Cracking, Oil & Gas Journal, March 26,1990. A.A. Avidan, Recent and Future Developments in FCC, Akzo Catalyst Symposium, Scheveningen, The Netherlands, June 1991.


41

CHAPTER 2

THE NATURE OF ACTIVE SITES IN ZEOLITES: INFLUENCE ON CATALYST PERFORMANCE

ADRIAN HUMPHFUES", DAVID H. HARRIS', PAUL O'CONNORb

"Akzo Chemicals Inc., 3250 E. Washington Blvd., Los Angeles, CA 90023

bAkzo Chemicals B.V.,

3800 AZ Amersfoort, The Netherlands P.0. BOX 975,

1. INTRODUCTION

Since their successful introduction some thirty years ago, zeolite catalysts have been the subject of considerable academic and industrial research efforts. Zeolites, or crystalline aluminosilicates, differ from more conventional crystalline materials in that the anhydrous crystal has a large, regular pore structure, making the internal surface available for adsorption or catalysis [l-81. Compared to other types of catalysts, zeolites are extremely active, especially in hydrocarbon conversion reactions and their regular pore dimensions make them selective as to which molecules are adsorbed or converted.

While there are some 40 natural zeolites [9], more than 150 zeolites have been synthesized [10,11], most of which have no known natural counterpart. Currently, only a small fraction of all these zeolites are of commercial interest. By far the major industrial process that utilizes zeolites is the catalytic cracking of petroleum and the replacement of amorphous silica- alumina materials by faujasitic (type Y) zeolites in the 1960's has saved the petroleum industry billions of dollars [12]. Other major commercial zeolite processes include [13,14] hydrocracking (faujasite), hydroisomerization (mordenite), dewaxing (ZSM-5, beta), iso/n-paraffin separation (Ca-A), olefin drying (K-A) and methanol to gasoline (ZSM-5).

In addition to their regular pore structure, zeolites are particularly suited to heterogeneous catalysis since an imbalance in charge between the silicon and aluminum atoms in the framework requires compensating cations be present to maintain electrical neutrality [1,3]. The catalytic activity of

42

zeolites is strongly influenced by the nature of the cation occupying the exchange sites on the structural framework [15,16]. Synthetic zeolites are typically grown in an alkaline environment and type Y, for example, is produced in the sodium form. For most hydrocarbon transformations, however, the basic alkali metal ion Y zeolites show little or no activity [a]. Fortunately, these cations are exchangeable and active acidic catalysts are prepared by replacing varying amounts of Na+ with ammonium ions or rare earth ions such as La3' or Ce3+.

Since it is now generally accepted [6] that the initial event in a cracking reaction is the formation of a positively charged carbon atom, or carbocation, the catalytic applications have been mostly related to those where the zeolite is used as a solid acid. Consequently, in the field of catalytic cracking of petroleum by zeolite containing catalysts , the terms "active site" and "acid site" are synonymous.

The purpose of this chapter is to describe the different types of active sites, their formation, characterization and catalytic reactions associated with hydrocarbon transformations in the general field of fluid catalytic cracking.

2 . ACTIVE SITES

Before the advent of catalytic cracking, the only route to lighter petroleum products was by virtue of thermal cracking, a process well known to proceed via a complex free radical path. A comparison [17] of products obtained by catalytic and thermal cracking is shown in Table 1. Inspection reveals that the shifts in product distributions are such that the two processes must proceed via entirely different mechanisms.

It is interesting to note that silica by itself has no activity for cracking and little acidity [18,19]. Gayer [20], however, in 1933, was the first to note that by introducing small amounts of alumina, both the activity and the acidity of the mixture began to rise. In the following year, Whitmore [21], in his theory of carbenium reactions, proposed acid sites as the active centers. More than a decade later , Hansford [ 22 , 231 and Thomas [ 241 independently proposed the concept of surface acidity and a mechanism to explain catalytic cracking in terms of ionic reactions rather than via a free radical path. The knowledge that a solid surface may be acidic led to the postulate [25] that the primary requirement for catalytic activity is the reaction of such a surface with a hydrocarbon to form a carbenium ion.

Some clarification of the nomenclature is pertinent here [6]. While the general name for a positively charged organic species is carbocation, under the IUPAC system, ions such as CH3+, formerly known as carbonium ions, are now to be called carbenium ions. The term carbonium ion refers to a positively charged species of the ' type CH5+.

43

Table 1 Comparison of Major Features of Thermal and Catalytic Cracking

Hydrocarbon Thermal Cracking Catalytic Cracking Type

n-Paraffins, C, is major product; e. g. , n-C,,H, with much C, and C,,

and C, to C,, a- olefins; little branching

shifts and little skeletal isomerization; H-transfer is minor and non-selective for tertiary olefins; only small amounts of aromatics formed from aliphatics at 932'F

than paraffins

Olef ins Slow double bond

Naphthenes Crack at slower rate

Alkylaromatics Cracked within side chain

C, to C, is major product; few a- olefins above C,; much branching

Rapid double bond shifts, extensive skeletal isomerization; H-transfer is major and selective for tertiary olefins; large amounts of aromatics formed from aliphatics at 9 3 2'F

If structural groups are equivalent, crack at about same rate as paraffins

Cracking next to ring is prominent

'

Reprinted from: P.B. Venuto & E.T. Habib, Jr., "Fluid Catalytic Cracking with Zeolite Catalyststt, M. Dekker, N . Y . , 1979, p.101

Thermal cracking was the first commercial process [26] for obtaining gasoline and other lighter petroleum products. Conversions were low but this method was utilized for more than twenty five years. Then, with the knowledge that an acid catalyst could be used to crack petroleum more efficiently, a variety of homogeneous materials, such as aluminum trichloride, were tried [27]. A long list of commercial problems, however, prompted a move to heterogeneous catalysts and success was finally achieved using natural clays [ 2 8 ] in conjunction with Houdry's continuous regeneration process [29]. Once this viable commercial process was in place, attention was focussed to the improvement of catalysts [30]. Acid washing of natural clays was soon followed by the introduction of synthetic silica- alumina materials. These synthetic materials were amorphous in nature, their structure consisting of a random array of silica and alumina tetrahedra interconnected over three dimensions. When forced into a tetrahedral configuration via oxygen bridges, an aluminum atom develops a negative charge and the resultant compensating cations are the source of the acid sites responsible for catalytic activity in carbocation reactions.

44

After their introduction around 1940, synthetic silica-aluminas dominated the cracking catalyst field for about twenty years. Incremental improvements in yields and selectivity were obtained with such variations as silica-magnesia and silica-zirconia, but the lack of a well defined (ordered) structure led to a wide distribution of active site acidities [31].

Table 2 Comparison of Product Distributions (Gasoline Fraction) in Gas Oil Cracking Catalyzed by Amorphous Silica-Alumina and Y Zeolite

Amorphous Product (~01%) Silica-Alumina Y Zeolite

Naphthenes 10 20

Aromatics 35 45

Olef ins 45 15

Paraffins 10 20

Reprinted from: P.B. Venuto & E.T. Habib, Jr. , "Fluid Catalytic Cracking with Zeolite Catalysts", M. Dekker, N.Y., 1979, p.40

A major breakthrough was achieved in the early 1960,s when Plank and Rosinski pioneered [12] the use of X and Y zeolites in cracking catalysts. The better defined properties of these zeolites, which had only been successfully commercialized a few years earlier [32], provided enormous activity and selectivity benefits to the field of catalytic cracking. Table 2 compares the product distribution obtained from gas oil cracking using crystalline aluminosilicate (zeolite) and amorphous silica- alumina catalysts [17].

Studies have indicated [33] that zeolites are at least l o 4 times as active as amorphous silica-aluminas, most likely due to a greater concentration of active sites. The most significant improvement, however, is the better selectivity of the zeolite cracking catalysts, which is sensitive to the distribution of acid site strengths [34], as well as the zeolitic pore geometry.

In crystalline aluminosilicates, all aluminum and silicon atoms form tetrahedra which are linked by shared oxygen atoms. These tetrahedra join with each other to form secondary building units, which can be interconnected to give numerous distinctive zeolites. Each has a regular and well defined pore structure together with inner cavities. This precise control of pore size is one of the greatest distinctions between zeolites and amorphous silica-aluminas. The structures of the two most important zeolites in the field of catalytic cracking, Y and ZSM-5, are presented for comparison in Figure 1 [35].

45

Figure 1. Structures of Zeolites Y and ZSM-5

Zeolites exist with either one-, two-, or three-dimensional pore structures [ 131. Zeolite Y has channels about 8b in diameter connecting cavities about 13w in diameter in a three-dimensional network which permits diffusion of hydrocarbon molecules into the interior of the crystal. This vast internal void space accounts for the high effective surface area of these materials. In contrast to the large pore faujasite, ZSM-5 is a medium pore zeolite with a unique three-dimensional pore structure consisting of a straight channel and an interconnecting sinusoidal channel which give rise to more shape selective petrochemical reactions.

The tetrahedrally coordinated aluminum atoms in the zeolite framework each carry a negative charge. The compensating cations, however, are not part of the structural framework and are located in several different sites throughout the pores and cavities of the zeolite [15]. Cation occupancy in zeolite Y is confined to the small (sodalite) cage and the hexagonal prism. They are easily exchanged by contacting the solid zeolite with solutions of ammonium and/or rare earth salts and this is the key to changing the acidity of the active sites on the material [ 15,161. An obvious goal for zeolite Y is to substitute protons for sodium ions, a prerequisite for carbocation formation. This cannot be easily achieved directly, however, since acidic solutions effect the removal of aluminum from the zeolite framework through hydrolysis and the structure collapses [36]. A relatively easy route to H-Y, shown in Figure 2, involves treatment with ammonium salts, followed by calcination above 600'F to decompose the NH4+ ion into NH, gas and H + , which maintains the structure [37]. The protons bond with oxygen atoms in the lattice to form -OH groups [38,39]. In this form, HY zeolite has the ability to transfer a proton to an adsorbed hydrocarbonand theactive sites areknownas Bronstedacid sites.

46

Na' Na' 0 0 0 0 0 0 0 \ / \ / \ / \ / \ / \ /

Si Al- Si Si A 1- Si A A A A A A .

0 0 0 0 0 0 0 0 0 0 0 0

NH; NH; 0 0 0 0 0 0 0 \ / \ / \ / \ / \ / \ /

Si Al- Si Si Al- Si A A A A A A

0 0 0 0 0 0 0 0 0 0 0 0

H' H' 0 0 0 0 0 0 0

Si Al- Si Si Al- Si

0 0 0 0 0 0 0 0 0 0 0 0

\ / \ / \ / \ / \ / \ /

A A A A A A

H' 0 0 0

H' I 0 0 0 0 \ / \ / \ / \ / \ / \ /

Si Al- Si Si Al- Si A A A A A A

0 0 0 0 0 0 0 0 0 0 0 0

Figure 2. Formation of Bronsted Acid Sites

47

In addition to direct proton exchange and the calcination of ammonium-exchanged Y zeolites, Bronsted acidity can also be introduced by two other routes. One involves the hydrolysis of ion-exchanged polyvalent cations followed by partial dehydration [40,41]. For example, a Y zeolite in its native sodium form can be treated with a commercial rare earth salt solution (typically a mixture of lanthanum, cerium, neodymium and praseodymium chlorides) to replace most of the sodium ions with polyvalent rare earth ions. The highly charged rare earth ions quickly hydrolyze, creating acid sites as shown below:

~a'+ + H,O --- > La(OH)2+ + H+

La(OH)'+ + H20 --- > La(OH),+ + H+

The rare earth-hydroxy ions occupy sites [ 4 2 , 4 3 ] in the zeolite framework that increase the thermal and hydrothermal stability.

The last, and least used method involves the reduction of metal ions to a lower valency state. For example [ 4 4 ] :

2 CU" + H2 --- > 2 Cu+ i 2 H+

Bronsted acid sites formed by any of these methods can be further dehydroxylated at temperatures in excess of about 750°F to form Lewis acid sites [38,39] as shown in Figure 3 . Lewis acid sites have the ability to accept an electron pair from an adsorbed hydrocarbon to create a carbocation (e.g. via hydride abstraction) .

H' H' 0 0 0 0 0 0

Si Al- Si Al- Si

0 0 0 0 0 0 0 0 0 0

\ / \ / \ / \ / \ /

A A A A A

0 0 0 0 0 \ / \ / \ / \ /

Si Al- Si' A1 Si A A A A A

0 0 0 0 0 0 0 0 0 0

Figure 3 . Formation of Lewis Acid Sites

48

The elimination of water by dehydroxylation should lead to the creation of one Lewis acid site from every two Bronsted sites. This was confirmed by Ward [39], who also observed that the number of Bronsted acid sites converted to Lewis sites increases with temperature until about 1500°F, when only the latter remain (see Figure 4 ) . However, if the calcination temperature is kept below about 110O0F, most of the Lewis acid sites can be rehydrated to restore the Bronsted sites [44].

Y

I

Y P I l2I d-

ACTIVATION TEMPERATURE (t)

Figure 4 . Measured by Pyridine Absorption

Relationship Between Bronsted and Lewis Acid Sites as

The concept of Lewis acid sites readily forming donor-acceptor complexes with, for example, H-, is the basis for Lewis site characterization (see Table 3, later) as well as dehydrogenation reactions with paraffins, which will be discussed in more detail later.

Factors influencing the acid properties of zeolites include the method of preparation, temperature of dehydration and the silica to alumina ratio and distribution of the framework [45-471 atoms. For example, it has been reported by Eberly [48] that the strength of the Lewis sites created via the thermal procedure described above is higher than that of the Bronsted sites.

Correlation between the acidity of a zeolite and its catalytic properties is a difficult task however. Three factors are important here: the total number of acid sites, the ratio of Bronsted to Lewis, and the acid strength distribution (and density) of each type of site. For Y zeolites, a maximum in strong acid sites and cracking activity occurs [49] at silica to

49

alumina ratios (SAR) from about 7 to 15. In contrast, for ZSM-5, hexane cracking ability increases linearly with increasing aluminum content [ 50.1 , leading to the conclusion that the maximum in acidity is a function not only of the zeolite structure but also the surroundings of the individual aluminum atoms in the framework.

To gain further insight into this subject, the concept of "Next Nearest Neighbors11 was proposed in order to provide a reasonable estimate of the acidic behavior of zeolites. The original work by Dempsey [51], Mikovsky and Marshall [52], suggested that the acid strength of a zeolite was related to the distribution of aluminum atoms in the framework. Each framework atom (Si or Al) in the zeolite is in tetrahedral coordination with oxygen and each aluminum atom has four silicon atoms as "Nearest Neighborsll. This is Lowenstein's Rule [53]. The four silicon sites in zeolite Y are connected to nine other framework atom sites and these are the "Next Nearest Neighborsll (Figure 5). According to the original work, the strongest Bronsted acid sites were associated with those framework aluminum atoms which had no "Next Nearest Neighbor" (0-NNN) framework aluminum atoms. Next strongest acidity is associated with 1-NNN sites, with a steady decline in acid strength through the 9-NNN sites.

Thus as aluminum framework atoms are removed from Y zeolite, stronger, more isolated 0-NNN acid sites are generated. As soon as all the acid sites are due to isolated 0-NNN framework aluminum atoms, a maximum in acidity is achieved (approximate SAR of 9-12). The activity will then decrease linearly with the removal of more framework aluminum since all the acid sites associated with the framework aluminum are of similar strength. ZSM-5, because it is of higher SAR than Y (typically 20-loo), has all the framework aluminum atoms as isolated 0-NNN and therefore shows a similar linear relationship of activity with the number of framework aluminum atoms.

the nine Next Nearest Neighbours (NNN) of an aluminium atom, Al. - oxygen bridges.

Figure 5. Next Nearest Neighbors

50

Pine et al. [47] , in their now classic paper, used this "Next Nearest Neighborm1 approach to predict catalytic cracking catalyst behavior with a Y zeolite unit cell model. This subject was recently reviewed by Wachter [54]. A recent paper by Lunsford [55] on the origin of strong acidity in dealuminated Y zeolites also invoked this approach to suggest that the interaction of isolated framework aluminum atoms with extra- framework aluminum species creates strong Bronsted acidity.

Studies using model compounds to look at the reactions of a hydrocarbon with a solid surface have concluded [6,56] that both carbenium and carbonium ion intermediates are involved. Bronsted sites can form both whereas Lewis sites produce only carbenium ions. Therefore, for many reactions, the activity of a catalyst should depend essentially on its Bronsted acidity, since Lewis sites alone do not appear to be active in most hydrocarbon reactions. However, Lewis sites are believed [57] to play a significant role in reactions such as double bond shifts and cisltrans-isomerization of olefins. This anomaly has been rationalized by Barthomeuf [58] through the concept of llsuperacidlt sites, which are believed to arise from a coupling of Lewis and Bronsted sites resulting in a considerable increase in activity of the latter. Others [55,59] claim that superacidity can arise from the interaction of Bronsted sites with extra-framework aluminum generated after ultrastabilization with steam. This concept of superacidity is discussed in more detail in the section on the characterization of acid sites.

Stabilization to a lower unit cell size reduces the number of acid sites on the zeolite since the latter are associated with the framework aluminum atoms. In addition to the increased hydrothermal stability, dealuminated Y zeolites are observed to have stronger acid sites than their higher cell size parents [60]: Depending on the dealumination process, however, various aluminum distributions are possible. Three methods to dealuminate Y zeolites are prominently mentioned in the literature:

(1) Steam calcination of ammonium exchanged Y zeolite (2) Chemical extraction of aluminum ( 3 ) Chemical substitution of aluminum

Although dealumination by steam calcination of ammonium- exchanged Y zeolites [61,62] can be easily achieved, significant quantities of mobile extra-framework aluminum species are generated which migrate to produce a zeolite with an aluminum- enriched surface (AES) which behaves more as an amorphous alumina, resulting in extra coke and gas [63] (see Figure 6). An improvement in selectivity can be made if the material is treated to reduce the amount of extra-framework aluminum.

51

In contrast] chemical extraction using mineral acids and chelating agents such as EDTA [64], or chemical substitution [ 651 using (NH,)*SiF,, affords a zeolite with an aluminum-depleted surface (ADS) resulting in a lower catalytic activity since the outer shell of the zeolite controls gas oil cracking [66-681. Steaming this material to optimize the ratio of framework aluminum to extra-framework aluminum can restore some activity.

Optimization of the zeolite aluminum distribution to produce a more homogeneous aluminum surface (HAS) has recently been described [69].

0 10 20 30 40

Argon etching time (rnin)

Figure 6. Aluminum Distributions

Another crucial aspect of zeolite design is the accessibility of the active sites and the relative contribution of the zeolite surface. The former is strongly influenced by Fluid Catalytic Cracking (FCC) catalyst preparation technology and the morphology of the zeolite, while the relative contribution of the zeolite surface activity will be determined by the zeolite crystallite size and the surface aluminum concentration. The route via which dealuminated Y zeolites are prepared will also have an impact on the pore size distribution of the zeolite

52

particle. Scherzer [62] reported the formation of secondary pores during the framework dealumination process. Lynch et al. [70] claim that the first steaming of NH,Y is mainly responsible for the formation of secondary pores which are, at that stage, completely filled by amorphous material. In this context, a subsequent acid treatment will not create a large quantity of new secondary pores, but eliminates the amorphous material resulting from the previous treatment. Patzelova et al. [71] have measured the secondary pores formed in two USY zeolites and conclude that these pores are not caused by a chaotic collapse of part of the zeolite framework, but by local reordering during the hydrolysis of framework aluminum. They also find an accumulation of non-framework aluminum on the secondary pore walls.

The accessibility of functional sites in a zeolite plays a crucial role in FCC catalyst performance. It has been demonstrated [72] that diffusion in FCC takes place in the non- steady state regime and this explains the failure of earlier attempts to relate laboratory measurements of FCC catalysts to theories on steady state diffusion.

4

- k l k o

''A

I . 1 . 8 . I .

0 1 2 3 4 5

Naphthenic rings

0

16 12 :3 14 16

Carbon number

k = Cracking rate ko = Cracking rate

of hexadecane

Figure 7. Active Site Accessibility of Amorphous and Zeolite Catalysts

Apart from the diffusion aspects, Nace et al. [73] have also found limitations in the accessibility of the zeolite portal surface by comparing the cracking rates of various model compounds with an increasing number of naphthenic rings. In this study they found major differences in the behavior of zeolitic and amorphous catalysts (Figure 7). This illustrates the difficulty of cracking heavy oil with zeolites alone. The understanding that the performance of all functionally active sites in a zeolite will be strongly influenced by their accessibility to the incoming molecules which are to be cracked has led to some new views on catalyst architecture [74].

53

3. CHARACTERIZATION OF ACTIVE SITES

A simple correlation between the number, type and strength of the acid sites in zeolites and their catalytic activity is the ultimate goal of refiners. While such a goal has yet to be fully achieved, considerable advances have been made in the area of quantifying and characterizing the active acid sites in zeolites as well as correlating the data with catalytic performance.

It is now possible to directly determine all aspects of zeolite acidity by a multiplicity of techniques. However, in most cases the conditions under which the measurements are made are far removed from the conditions experienced in a real life situation, such as an FCC unit. Thus only general trends can be realized from the data. Consequently, indirect methods such as model reactions using probe molecules, e.g. n-hexanelheptane, cumene, ethyl benzene or xylenes, under conditions closer to those of a real FCC or isomerization unit, play an important role in relating acidity to activity.

However, even model reactions require caution in interpretation, since a probe molecule such as n-heptane can easily penetrate the pores of Y zeolite. Catalytic cracking can therefore occur at acid sites throughout the zeolite framework. In the real world of FCC, gas oils rarely penetrate much beyond the outer surface of the zeolite. Unless the acid sites on the outer surface are similar in strength, type and density to the inner channel surfaces, the conclusions from the n-heptane probe experiment may bear no relationship to gas oil cracking. This has been shown in a study of gas oil cracking on dealuminated Y zeolites in which it was found that an acid site gradient did exist from the external to internal surfaces and the correlation between acidity measurements using pyridine and n-heptane cracking (both can penetrate the zeolites pores) was consistent, but inconsistent with gas oil cracking [67].

A further problem with model reactions is that no single model compound utilizes all the acidic sites [44]. Thus to fully characterize the acidic nature of a zeolite, a number of different model reactions are necessary. For example, the isomerization of butene will require much milder acid sites than the cracking of butane [17]. Since time constraints do not usually allow for multiple model reactions, a combination of direct measurements of acid character together with a model reaction is generally used.

The realization that zeolites behave like solid acids has allowed the use of characterization techniques similar to those used for aqueous acids, e.g. indicators. some of this work has already been covered in reviews [75,76], in chapters of broader texts [6,8,15,44,77-801 and even in a substantial part of a symposium on catalysis by acids and bases [81] . This section will draw upon this earlier work together with more recently published data. Tables 3 and 4 summarize both direct and

54

indirect methods of determining the acid character of an active site. These tables and in particular Table 4 are not exhaustive in their coverage. The intent here is to give some of the more recently published model reactions as a guide. (Refer to references 6,44,76,80 and 81 for other reactions together with references 8,75 and 79 for more detailed information on direct methods). Since it is not possible in this article to discuss each method in detail, in general only the more frequently used methods will be described.

4. DIRECT METHODS OF ACIDITY DETERMINATION

One of the oldest methods of determining the acidity of solid acids came from a modification of solution chemistry. The Hammett acidity function (H,) and the aryl alcohol indicator (Hp) are analogous to pH and are directly related to the solid acid equilibrium constant (K,) as given in equations (1) - (6) ,

H+ + B = BH+ (1)

H, = pK, + log [B]/[BH+] (2)

A + :B = AB ( 3 )

Ho = PKa + 109 [Bl/[ABI (4)

H+ + ROH = R+ ' + H,O (5)

H, = pK, + log [ROH]/[R+] (6)

where [B] , [BH'] , [AB] , [ROH] and [R'] are the concentrations of the neutral base, its conjugated acid, the addition product of a Lewis acid and a Lewis base, the aryl alcohol concentration and its conjugated acid, respectively.

A series of Hammett indicators has been developed which cover the whole range of acid strengths. These indicators have a different color in the neutral base form than in the conjugate acid form. Thus if the indicator assumes the color of its acid form there must be some acid sites with an H, less than or equal to the pK, value of the indicator. The Hammett indicator (H,) measures the total acidity (i.e. Lewis plus Bronsted) whereas H, indicators give only the Bronsted acidity.

Typically a solid sample is taken and a known amount of butylamine is added in a non-aqueous solution. After equilibration the relevant Hammett indicator is adsorbed, the color indicating either the acidic or basic form. The amount of butylamine necessary to convert the indicator back to its basic form gives the number of acid sites with strength higher than the pK, of the indicator used. The pK, values can then be related to sulphuric acid concentration for easier comparison.

55

Table 3 Direct Methods of Acidity Determination

To Determine the Total Number of Acid Sites Reference

Amine titration Aqueous titration Adsorption-desorption of bases Calorimetry Poisoning of acid sites by bases 'H MAS NMR

To Determine the Acid Site Strength Distr

Adsorption of color indicators W-visible spectroscopy

Calorimetry Shifting of infrared absorption bands 'H MAS NMR Aromatics adsorption

To Determine the Type of Acid Site

Bronsted Sites

Adsorption-desorption of bases a4 , a5

Measurement of H, evolution from reaction with hvdrides

(ii) Aqueous titration/cation exchange (iii) Amine adsorption (iv) Spectroscopy: W-visible

'H MAS NMR

8,75 75 82-87 84,aa, 89 90-92 93-95 , 107

bution

96,97 96-98

a7,92,99,ioo,io5 84, aa, a9

99,101,102

9aIio3-io5 93-95 , 107

75

75 a6,iooIio5,io6

98 I 4 4 93-95,107

FTIR XPS I4N MAS NMR Luminescence

(b) Lewis Sites (i! Reaction with electron donors (ft! Amine Adsorption (111) Spectroscopy: W-visible

ESR/ ENDOR FTIR XPS "N MAS NMR Luminescence

a ~ , ~ 2 , ~ ~ , 1 0 2 , i o 5 , i o 6 , i o a 109 110 111

75 , 101 100,101,105,106

44 112

87 , 101,102 , 106 109 113 111

There are several problems associated with H, and H, indicators. Visual observation of color changes can be extremely difficult. Recent work has shown that a spectrophotometric method is much more reliable [96,97]. This has allowed the authors to establish the relative strengths (to sulphuric acid) of many solid acids including zeolites [97]. As is typical in nearly all solid

56

acid characterization experiments, the data were correlated with a cracking reaction of a model compound, isobutane. The results suggested that acid strength (the intensive factor) dominated over the concentration of acid sites (the extensive factor).

Another problem is that many of the indicators are too large to enter into all the inner sites of many zeolites. The bases used, such as n-butylamine, are also not necessarily specific to acid sites and may adsorb on the zeolite cations, giving artificially high acid concentrations [44]. Finally, as is often the case for direct acidity measurements, the conditions used to determine the H, and H, values are far removed from those of fluid catalytic cracking. Correlations of acidity from these values with catalytic data must therefore be treated carefully.

While Hammett indicators with amines will yield information on both the density and strength of acid sites, aqueous titration will only give the total number or density of acid sites. However, stable end points are rare and the method is no longer significant [75].

The adsorption of gaseous bases, particularly ammonia, followed by monitoring the temperature programmed desorption (TPD), has been used extensively for acidity measurements (see Table 3). In many cases the technique is coupled with other techniques such as calorimetry [ 8 4 ] or infrared spectroscopy [80,99]. Caution is also needed here in the choice of appropriate base. If Lewis and Bronsted acid sites are regarded as being in equilibrium with one another (see Figure 3 ) , then a strong base would drive the equilibrium towards Bronsted acidity. Conversely, a weak base such as hexane would not disturb the equlibrium as much and preferentially interact with Lewis sites [80,88]. Consequently the use of only strong bases may indicate stronger Bronsted acidity than in reality. Ideally several probe molecules covering a broad range of basicity should be used.

To complicate matters further, it has been found that adsorption and TPD of aromatics indicates they preferentially interact with alkali metal cations such as Nat rather than Ht in H-NaZSM-5 [104]. However, keeping the alkali metal cations to a minimum allows the use of aromatic probe molecules for acid strength determination. The protonation of toluene, xylene and trimethylbenzenes on H-Y, H-Mordenite and H-ZSM-5 was followed by W-visible spectroscopy and used to determine the relative acid strength of the zeolites. Dealuminated mordenite and H- ZSM-5 had the strongest acid sites, H-Y had the weakest [98].

Despite the cautions outlined above, the adsorption/desorption of basic molecules is a frequently usec! technique for acidity measurements. The two most popular bases are ammonia and pyridine. Ammonia is very readily and reversibly chemisorbed at temperatures >300°F, making it ideal for TPD studies, which can give an indication of acid site strength as well as the total acidity. The desorption peak maximum is related to the rate at which the sample is heated, so care is needed when comparing

57

data from different sources. Recent claims of greater accuracy using improved TPD techniques have been reported [83]. Pyridine is less amenable to TPD studies as dehydroxylation can occur at temperatures above 390°F [114]. However, a comparison of the adsorption of both bases on dealuminated Y zeolites showed that pyridine gave the best correlation with the theoretically predicted acid sites [85]. The combination of pyridine adsorption and infrared spectroscopy is extremely useful for distinguishing between Bronsted and Lewis acid sites (see Figure 4). A stretching vibration at 1545 cm-' was assigned to the pyridinium ion (Bronsted site), the 1454 cm-' vibration was assigned to coordinated pyridine (Lewis site) and a 1485 cm-' band has been assigned to a combination band [114,115].

Pyridine and 2,6-Dimethylpyridine have successfully been used in the absence of infrared spectroscopy to distinguish Bronsted and Lewis sites in zeolites X, Y, L and Mordenite using a GC adsorption/desorption technique [loo]. 2,6-dimethylpyridine will preferentially adsorb to Bronsted sites. The method cannot be used as readily for smaller pore zeolites such as ZSM-5 since the diffusion rate of 2,6-dimethylpyridine is too slow; however ammonia-TPD on ZSM-5, without complementary infrared spectroscopy, has shown that the highest temperature peak can be assigned to Bronsted acid sites [86].

An area of great interest in the past few years has been the effect of chemical and physical (steam) dealumination on Y and ZSM-5 zeolites. Studies on ZSM-5 in this complex area using adsorbed pyridine with infrared spectroscopy have clearly indicated the Bronsted acid sites are of similar strength and do not change in the absence of steam [105]. Steaming the ZSM-5 resulted in a loss of Brijnsted sites although the activity of hexane cracking increased [115]. This would suggest the presence of "superacidtt sites as has been proposed for steamed Y zeolites [55].

Considerable work using poisoning experiments, MAS NMR, IR and catalytic studies has been carried out on the effect of steaming zeolites and will be discussed in more detail later in this chapter. In the case of HY, a range of dealuminated zeolites of differing SiO,/A1,0, ratios were investigated by gravimetric adsorption of pyridine and it was found that maximum acid strength occurred at about 30 aluminum atoms per unit cell [106]. This would correspond to a silica to alumina ratio of about 11. Many different acidity measurements have given a similar result. In a comparison study of different dealumination techniques on HY zeolite using TPD of ammonia [87], FTIR [92] and microcalorimetry of ammonia adsorption [89], it was found that steaming produced cationic aluminum species which poisoned the strongest framework acid sites. These sites could be recovered by an optimized acid leaching. Dealumination by ammonium hexaflilorosilicate generated more acid sites than steaming but few were Lewis sites.

58

In addition to using amine adsorption to determine the density, strength, and in some cases the type of acid sites, a recent paper [82] has studied the adsorption of amines of increasing chain length in ZSM-5 to determine the distribution of acid sites. The principle is illustrated in Figure 8.

< > Sites closer t h a n this along t h e

channel d i rect ion cannot b o t h be

used by these amines.

Figure 8 . Steric Limitation on the Use of Adjacent Acid Sites

Calorimetry of ZSM-5 and Mordenite has been used in conjunction with ammonia adsorption/desorption [84,116]. In TPD experiments, the temperature peak maxima are influenced by acid site strength, the number of acid sites, the zeolite structure and the heating rate. Consequently a comparison of acid site strengths between different zeolites requires additional information provided by infrared spectroscopy or calorimetry. The disadvantage of both calorimetry and pyridine adsorption- infrared spectroscopy is that they are slow techniques compared to TPD. However, the combination, especially when a model cracking reaction such as n-octane is also studied, allows excellent correlation of acidity with catalytic data [116].

Infrared studies with probe molecules can also yield valuable information on hydroxyl acidity, i.e. Bronsted acid sites. Acidity studies using infrared spectroscopy require that the samples be rigorously dried, however, and all measurements made under anhydrous conditions since the HY Bronsted sites react with water even at room temperature [108]. The area of interest, the hydroxyl region, involves absorption bands above 3500 cm-' and of course water will mask this region. The presence of extra-framework aluminum hydroxyl species after steaming further complicates this region [92,117].

The use of 'H MAS NMR should also be an excellent tool for characterizing acid sites but unfortunately there are many problems associated with the chemical shift values used to determine the acidity of the hydroxyl groups (Bronsted sites) under observation [94]. Useful data are starting to appear, however [93,95,107].

59

Traditionally, Bronsted sites were measured either by the volume of hydrogen evolved upon treatment with LiAlH, or by aqueous titration [75]. These methods have now been replaced by spectroscopic methods (IR, W-vis) as discussed earlier. More recently, x-ray photoelectron spectroscopy of sorbed pyridine on ZSM-5 [lo91 and luminescence of several zeolites containing sorbed pyridine [lll] have been reported as useful techniques to distinguish Brijnsted and Lewis sites. Solid state 14N MAS NMR of tetraalkyl ammonium cations present in zeolites has also been found to be a convenient method for Bronsted acid site determination [110].

Lewis sites can be determined by many of the methods used for Bronsted sites, although recently two methods s ecifically for Lewis acidity have been published. Solid state N NMR of HY and H-ZSM-5 with adsorbed enriched 15N,0 can distinguish Lewis acid sites of varying strengths [113]. ESR and ENDOR spectra of qui- nones adsorbed onto thermally activated HY can identify radical cations due to donor-acceptor complexes with Lewis acids [112]. Radical cations have also been observed at low temperatures in the interaction of olefins with zeolites [118]. These radical cations may play a role in the catalytic reactions of olefins.

This review has concentrated on the Bronsted and Lewis definitions of acidity, but other forms of acidity have been proposed [119,120], one of which, Usanovitch acidity, invokes the concept of one electron transfer to form radicals and radical ions.

I?

5. INDIRECT METHODS OF ACIDITY MEASUREMENT

The use of probe molecules in model catalytic reactions has been found to be an excellent method to characterize the acidity of catalysts and some are presented in Table 4 . Guisnet, in an excellent review of the method [121], has stated that many reactions involving olefins, aromatics and paraffins can be used. The reactions involved need to be simple, free of side reactions and preferably catalyzed by only one type of active site [121]. If these conditions are met, then the use of several different model reactions on the same catalyst should allow full acid site characterization, especially if a direct method of acidity determination is also included.

Several good examples of the use of model reactions such as isomerization and cracking have been carried out on a catalyst containing HY zeolite [56]. From extensive earlier work it is knownthat different reactions require different acid strengths. The cracking of paraffins requires the highest acid strength, followed by alkylaromatic cracking, aromatic alkylation, olefin isomerization and alcohol dehydration needing the weakest acidity [144]. With this knowledge, the impact of pyridine desorption temperature on the various reactions gives the minimum strength that acid sites must have to catalyze the reaction [56]. Significant differences between reactions was

60

observed, with the slowest reactions requiring the strongest acid sites for catalysis. For example, isomerization of 3,3- dimethyl-1-butene requires weak acid sites and is about one hundred times faster than n-hexane cracking, which requires strong acid sites.

Table 4 Indirect Methods of Acidity Determination Using Probe Molecules in Catalytic Reactions.

1. Cracking Reactions Reference

(a) n-Hexane (alpha test) (b) n-Butane (c) n-Heptane (d) Cumene (e) Isobutane ( f) Neopentane (9) 2,3-dimethylbutane/3-methylpentane

2. Isomerization Reactions

(a) Xylene (b) Cumene

3. Disproportionation Reactions

(a) Toluene (b) Ethylbenzene (c) Cumene

4. Alcohol Dehydration Reactions

(a) Methanol to hydrocarbons (b) 1- and 2-Butanol to hydrocarbons (c) Isobutanol to hydrocarbons (d) t-Butanol to hydrocarbons (e) Cyclohexanol to hydrocarbons

5. Miscellaneous Reactions

(a) H, - D, Exchange (b) Cyclohexene hydride transfer

122 , 123 124 67,125,126 127 , 128 97 , 129 90,129 130

128,131,132 127,128

133,134 135,136 127 , 128

128 , 137 138 139 140 141

75,142 143

The disadvantage of these model reactions is that they are very time consuming if a detailed acid characterization is desired, although the use of a microreactor within a GC instrument can speed up the process [145].

61

ACTNITY "ALPHA"

10

10 100 1000 PPM AL I I I I 1 SiO,

100,000 10,000 1.000 100 10 - A ' A

Figure 9. Catalytic Activity of ZSM-5

The use of model compounds to carry out simple catalytic activity comparisons of several zeolites without trying to obtain detailed acid strength characterization is a fairly simple and rapid process. Along these lines, Mobil has developed an alpha test, using n-hexane as the model feedstock, to evaluate their many ZSM zeolites [122,146]. Figure 9 illustrates the correlation between the framework aluminum content of unsteamed ZSM-5 and its catalytic activity. Union Carbide has developed a similar test using n-butane and has found (Figure 10) that stronger acid sites are associated with higher silica content zeolites, as reflected by faster cracking rates [124].

Of the other probe molecules used for model reactions (Table 4), cumene is of particular interest since it can undergo many different types of reactions, providing information on the relative concentrations of Bronsted and Lewis acid sites [127].

Although considerable advances have been made in the characterization methods for the total number, type and relative strength of acid sites, there are still many inconsistencies between acidity and catalytic activity.

Y zeolites are known to increase in catalytic activity as aluminum is removed from the framework, passing through a maximum at about 30 A1 atoms per unit cell (e.g. , see Figure 11). This can be explained by invoking the concept of isolated (0 Next Nearest Neighbors) framework aluminum atoms having the strongest acidity, as mentioned earlier [ 5 4 ] . However, only a fraction of these strong acid sites appear to be catalytically active.

62

SdAI

looo ~ 32: 15, 7 110 1 l7 3 7 4 G 3," 312 2; 2

c / \

P ._ t / I " 10

NH.V (61"

NH,ERIONITE"'

\ - " NH,ERIONITE"

0.1 1 I I I I I I I I I 0 .03 .06 .07 .12 .15 .18 .21 .24 .27 .30

Al

Al + Si -

Figure 10. Cracking of n-butane: the pseudo-first-order rate constant K, versus aluminum content

This has been shown by the impact of poisoning experiments with NH, [ g o ] and Na+ [91]. Much smaller quantities of poison are needed to destroy the catalytic activity than expected based on the total framework aluminum content. It has also been found that extra-framework aluminum species, generated during most methods of dealumination, are necessary for maximum catalytic cracking activity [41]. The amount and type of extra framework alumina necessary is dependent on the reaction under study, and is the subject of much debate in the literature [55,147].

63

However, recent work has proposed that cations such as [A1(OH),A1I4+ occupy exchange sites in the beta cages and may be bound to framework oxygens. Nearby acid sites become I1superacidt1 sites through the powerful inductive effects of these cations [55]. These "superacid" sites are thought to be the source of the high catalytic activity and are easily poisoned. The introduction of rare earth cations to the zeolite framework is thought to proceed in a similar fashion, although the absolute activity is lower than for aluminum cations [41].

There is some debate about whether the aluminum cations contribute to super Bronsted acidity or not. Studies of aluminum exchanged-Y zeolite, USY and SiC1,-dealuminated Y, indicated that cationic aluminum gave strong Lewis acid sites, but the Bronsted acid strength decreased with increasing extra- framework aluminum [148].

; E

0

c <

0.5

0.4

0.3

0.2

0.1

0 10 20 30 4 0 5 0 Framework A l / Unit C a l l

Figure 11. Dependence of catalytic activity for cumene conversion at 550'F on framework aluminum content

64

ZSM-5 has also been found to have increased catalytic activity after mild steaming. Severe steaming or dealumination in the absence of steam gave the expected correlation of activity with framework aluminum content (see Figure 9) [149,150]. Poisoning experiments with NH, and Na+ gave the same results as for zeolite Y [90,91]. Since ZSM-5 is a high silica zeolite, it already has isolated aluminum framework atoms, thus the increased activity has to be due to non-framework aluminum species. It has been proposed that mild steaming generates aluminum species which interact with nearby Bronsted acid sites to generate "super- acidityt1, much the same as for Y zeolite, and requires pairs of aluminum framework atoms [ 149,1511 . Subsequent MAS NMR and catalytic studies have shown that there is no enhanced Bronsted acidity nor a significant number of paired aluminum framework atoms. The enhancement in catalytic activity is due to interaction of the reactant molecule with a framework hydroxyl (Bronsted site) and an extra-framework aluminum species which is bound in someway to the framework. Attempts to remove the extra framework species by acid were unsuccessful. It was concluded the species were small, non-charged, non-hydroxylated and therefore Lewis acids [152]. It is clear that the geometry around the Bronsted acid site will play an important role in allowing the extra-framework aluminum species to bind to the framework.

Related work by Loeffler, Kazansky and co-workers used diffuse reflectance infrared spectroscopy to study steamed H-ZSM-5. They found up to five different types of acid sites were generated depending on the steaming conditions used [117]. Most of the non-framework aluminum species were hydroxylated and weaker in acidity than framework bridging hydroxyls. Some, however, exhibited Lewis acidity and were found capable of polarizing small paraffin molecules [153]. These species were thought responsible for the enhanced catalytic acitivity in n- hexane cracking. It was proposed that the Lewis sites initially dehydrogenate the hexane to 1-hexene which then rapidly cracks at the framework Bronsted sites. This was supported by the fact that 1-hexene addition to mildly steamed H-ZSM-5 inhibited cracking activity, but enhanced activity with unsteamed H-ZSM-5 containing no Lewis acid sites [154]. In general Kazansky and co-workers feel the many different types of Lewis acid sites, both within and outside the zeolitic framework, play an important role in initiating the many different reactions in catalytic cracking.

The concept of enhanced activity from a synergism between an extra-framework Lewis acid and a framework hydroxyl invokes the role of local or site geometry. So far only the numbers, strength, distribution and type of acid sites have been discussed. The geometry around the acid sites may also be important. An ab initio molecular orbital calculation to monitor the effect of increasing the Si-0-A1 angle (TOT angle) on acidic properties has been performed. It was proposed and verified experimentally that increasing the TOT angle led to a decrease in acidity of the bridging hydroxyl groups [155].

65

The accurate measurement of site geometries is extremely difficult. Techniques used for pore geometry determination such as adsorption [ 8 ] , decane cracking [156] and 129Xe MAS NMR [157] are not sensitive enough for site geometry. A recent paper investigating Iz9Xe MAS NMR line splitting, however , has indicated the potential for obtaining microscopic information within zeolite crystals [158]. Advances in instrumentation such as MAS NMR, synchrotron powder X-ray diffraction, scanning electron microscopy and combinations of these techniques, have created powerful analytical tools for detailed structure determination [159]. Recently, for example, combined use of such techniques was able to show that the supposedly different zeolites, KZ-2, Theta-1, ZSM-22 and NU-10 had the same crystal structures but different morphologies [160]. The introduction of two-dimensional (2D) 29Si MAS NMR with x-ray diffraction, allows the determination of three-dimensional silicon-aluminum framework connectivities, either on *'Si enriched zeolites [161] or zeolites containing 29Si in natural abundance [162]. The recent development of double resonance (DOR) 27Al MAS NMR has dramatically improved the resolution of 27Al MAS NMR spectra [163]. Prior to this technique, 27Al MAS NMR had played a minor role to 29Si MAS NMR in structure determination, due to the problems of quadrupolar interactions. Future use of combinations of these techniques can be expected to yield detailed information about structures, bondin and hence local geometries. Already the use of 'HI 27Al and 'ki MAS NMR on HY and H-ZSM-5 has given information on the site geometry of isolated Bronsted acid sites [164].

Earlier it was stated that one of the major problems with zeolite acidity characterization is that many of the methods used are at conditions far removed from those of the industrial processes for which the zeolite based catalyst is intended. To compensate for this, acidity measurements under ambient conditions will often be combined with cracking reactions of model compounds at conditions similar to commercial operations. However, the ideal would be to make all measurements at typical catalytic cracking conditions.

A new technique, Temporal Analysis of Products (TAP) can determine products under catalytic reaction conditions. Typically products are identified by a real time quadrupole mass spectrometer [165]. This technique has not yet been applied to zeolitic cracking. In situ cross polarization magic angle spinning "C NMR (CP/MAS NMR) of samples sealed in a variable temperature CAVERN (cryogenic adsorption vessel enabling rotor nestling) apparatus, has been used to monitor the reaction of propene on HY from very low [166] to high [167] temperatures. It was found that enriched 13C propene was highly mobile even at temperatures well below ambient. Oligomerization occured at ambient temperature and cracking at high temperatures. The only carbocations observed were those of alkyl substituted cyclopentenyl cations which were not thought to be catalytically involved in oligomerization. However, long-lived alkoxy species

66

Table 5 Comparison of Thermal and Catalytic Cracking

Thermal Cata lvt ic

Catalyst - SiOJ A1,0,

C, composition:

ic, (wt%) nC, (wt%) c,= (wt%)

0.9 2.8 2.3

iC,/nC, 0.3 (Degree of Branching)

c6 composition: (~01% on C, paraffins)

n-hexane 63 3-methylpentane 18 2-methylpentane 16 2,3-dimethylbutane 3 2,2-dimethylbutane 0

8 . 4 2.5 6.0

3.4

9 48 27 13 3

iC,/nC, (DOB) 0.6 10.1

Gasoline composition: (~01% on gasoline)

paraffins naphthenes olef ins aromatics

53 14 30

3

56 19 9 16

Degree of Cyclization 0.6 3.9

Gasoline Octanes

RON MON

70 63

91.5 80.5

presumeably formed from covalent bonding of incipient carbocations with framework oxygen atoms, were observed. Although this technique is extremely useful for observing reactions in situ, a disadvantage is that the product selectivity- will be affected by the inherent long residence times. Thus olefins and aromatics are observed in flow reactor cracking of propene but not in the CAVERN apparatus.

It is anticipated that as these new techniques become more widely available, the correlation between cracking reactions and zeolite acidity will improve dramatically.

67

6. CATALYTIC CRACKING WITH ZEOLITES

The positive impact of carbenium-ion cracking on FCC gasoline yields, composition and octanes was the major benefit in the switch from a thermal to a catalytic cracking process in the 1940,s. Oblad et al. [168] showed that catalytic cracking of vacuum gas oil produces a much more desirable gasoline product than that from thermal cracking. While the latter is quite olefinic, catalytically cracked gasoline contains a large amount of aromatics and branched paraffinic compounds.

Table 5 shows a comparison of gasoline compositions for thermal and catalytic cracking at about equal gasoline yield. Both the C, and C, paraffin compositions clearly indicate a drastic shift towards a higher quantity of branched compounds with the catalytic process and the aromatics yield is strongly enhanced.

Similar observations were reported by Greensfelder et al. [169], who evaluated several catalysts by cracking pure hydrocarbons. Their results with n-hexadecane also show a strong increase in the formation of branched paraffins with acidic cracking catalysts.

The reactions that occur when a hydrocarbon molecule reacts on the solid surface of a catalyst all involve positively charged organic species, usually carbenium ions [16]. The initial reaction to form a charged species can occur via a number of paths, the most common being hydride abstraction [170-1731. Figure 12 shows examples of hydride abstraction from. (1) a paraffin at a Lewis site (L) and (2) via direct attack of a proton from a Bronsted site (H').

(1) CH,-CH2-CH2R + L ----- > CH3-C+H-CH2R + LH-

(2) CH3-CH,-CH2R + H+ ----- > CH3-C+H-CH,R + H,

Figure 12. Formation of a Carbenium Ion via Hydride Abstraction

Other paths to carbenium ion formation include the addition of a cation to an unsaturated molecule and heterolytic Eission. The stability of carbenium ions increases as follows [16,174]:

Primary < Secondary < Tertiary

Tertiary species are thus greatly favored, accounting for the high degree of branching associated with catalytic cracking [16,172]. Only an acidic type of ionic cracking involving a tertiary carbenium ion will produce branched compounds such as isobutane (iC,) . On the other hand, thermal cracking, by virtue

68

of its free radical path, will tend to produce non-branched compounds and fuel gas (C;) [6]. The ratio C,-/iC, then seems to be an appropriate indicator to monitor the l1quality1' of catalytic cracking [175]. The following data (Table 6 ) from Greensfelder et al. [169] illustrate that strong selectivity advantages are achieved with catalytic cracking.

Table 6 Selectivity Advantages of Catalytic Cracking

iC,/nC, C,-/iC, Fuel Gas (mol/mol cracked)

THERMAL 0 6 2

THERMAL Actv. Carbon 0.06 27

CATALYTIC Alumina 0.2 14 SiO, /A1203 3.8 0.6

0.54

2.72 0.41

Once formed in the initiation reaction, carbenium ions can pursue a number of different reactions which are determined by the nature and the strength of the acid sites involved. The three dominant reactions of carbenium ions are [6,16,17]:

(I) The cracking of a carbon-carbon bond (2) Isomerization ( 3 ) Hydrogen transfer

Other reactions such as alkylation, cyclization and condensation also occur, along with some reverse reactions such as polymerization, dealkylation and dehydrogenation [172].

The cracking reaction, otherwise known as beta-scission, is a key feature of ionic cracking and is responsible for the majority of motor transportation fuels produced from crude oil. In this process, shown in Figure 13, the C-C bond located in the beta position to the carbon atom with the positive charge is broken. The energy required to split this bond is lower than that needed to break adjacent C-C bonds since a higher activation energy is necessary for methane and ethane formation. The products of this cracking reaction are an olefin (which is promptly desorbed) and a new carbenium ion, allowing the sequence to continue. Beta-scission is a unimolecular, endothermic reaction and is therefore favored by high temperatures but is not equilibrium limited.

69

CH,-C+H-CH,-CH,-CH,-CH,R ---- > CH,-CH=CH, + C+H,-CH,-CH,R

Figure 13. Beta-Scission

The isomerization of carbenium ions, often an endothermic reaction, can occur via a charge or a skeletal process. An example of charge isomerization, a 1-2 hydrogen shift in the propyl cation, is illustrated in Figure 14(a).

(a) Charse Isomerization

CH,-C+H-CH, = CH~-CH,-C+H,

(b) Skeletal Isomerization

CH,-CH,-C+H-CH,-CH~R = CH,-F+-CH,-CH,R = C+H,-FH-CH,-CH,R

Figure 14. Isomerization of Carbenium Ions

CH3 CH3

Figure 14(b) shows an example of skeletal isomerization consisting of a 1-2 hydrogen shift and a 1-2 alkyl shift. Although hydrogen shifts of this type are much faster than alkyl shifts, both types of migration are relatively easy for carbenium ions, leading to product configurations with high ratios of branched to normal products. Isomerization also allows primary carbenium ions to rearrange to more stable species prior to beta-scission [176].

Hydrogen transfer, or more correctly, hydride transfer, is the dominant reaction of Y zeolite and is the reason for the observed selectivity differences between Y zeolite and amorphous silica-alumina catalysts. Hydride transfer reactions involving paraffins and carbenium ions, as shown in Figure 15, are critical in the catalytic cracking of hydrocarbons since they are responsible for the chain propagation that occurs after the formation of the first carbenium ion on the catalyst surface [177].

CH,-C+H-CH2-CH,R + RlH ---- > CH,-CH,-CH,-CH,R + R,'

Figure 15. Hydride Transfer

70

The hydride transfer benefits associated with zeolitic cracking lead to several significant advantages over amorphous silica- aluminas. Firstly, the activity of the Y-zeolite catalysts is much higher [12,33,178,179], affording substantially larger gasoline yields. Secondly, due to the dominance of hydride transfer, the composition of the gasoline produced over Y zeolites is much less olefinic and more aromatic [180]. Hydride transfer promotes the interaction of olefins with naphthenes to form more refractory paraffins and aromatics (Figure 16). By reducing the concentration of highly reactive olefins, the number of secondary reactions is cut substantially and the gasoline is ttstabilizedtt.

Olefins + Naphthenes ---- > Paraffins + Aromatics

Figure 16. Stabilization of Gasoline via Hydrogen Transfer

Thirdly, more C, to C,, products and less C, and lighter products are formed, once again attributed to the increasing importance of hydride transfer. This observation, by Nace [177] and Thomas [181] is due to the rate of hydride transfer being much greater than the rate of beta scission, causing the cracking to stop at a higher molecular weight via hydride transfer to the carbenium ion.

The hydrogen redistribution to stabilize gasoline is a consequence of the high concentration of hydrocarbons and the abundance of acid sites on the zeolite surface. Since the catalytic selectivity of a zeolite is a function of both acid site density and acid strength distribution [171], the degree of hydrogen transfer can be influenced by stabilization to a lower unit cell size. As *framework aluminum atoms are progressively removed, the number of acid sites on the Y zeolite surface continues to decrease while the acid strength of those sites remaining continues to increase [60] reaching a maximum at about 30 aluminum atoms per unit cell (Next Nearest Neighbor Theory). Stronger acid sites favor tighter binding of a carbocation to the zeolite surface [171] and decreasing site density implies larger distances between adjacent sites. Tighter carbocation binding will lead to a decrease in hydride transfer, as will decreasing site density, since hydride transfer is a bimolecular reaction.

This strategy was successfully employed in the 1970,s to meet the challenge of lead phasedown from gasoline [69]. Zero (or low) rare earth content Y zeolites were ysed in commercial FCC catalysts to minimize hydride transfer, thereby boosting light olefin yields, thus compensating for the loss of octane due to the removal of tetraethyl lead.

71

In contrast to zeolite Y, pentasil zeolites such as ZSM-5 are synthesized with only trace amounts of aluminum. Consequently the acid sites on this zeolite are few in number but very high in strength. These, together with the elliptical straight channels and the near circular sinusoidal channels, explain the unique catalytic properties of this material which is used in FCC operations to increase the yields of high octane C,-C, olefins at the expense of lower octane normal- and monomethyl- substituted paraffins in the gasoline fraction [182].

Several model reactions have been used to help elucidate reaction mechanisms. Abbot and Wojciechowski have used a range of paraffin feedstocks with HY zeolite to determine the mechanisms involved in catalytic cracking [183]. They have proposed two mechanisms, both involving Bronsted acid sites. One mechanism, thought to occur predominantly with linear paraffins, involves initiation via pentacoordinated carbonium ions followed by either cleavage to a smaller paraffin and a carbenium ion, which can adsorb onto the catalyst, later desorbing as an olefin, or rearrangement of the carbonium ion followed by cracking to give branched paraffins. Hydride transfer can also occur between the adsorbed carbenium ion and a feed molecule, which leads to cracking of the feed molecule via beta-scission of the resulting carbenium ion. This chain type mechanism was proposed to be slower than protolysis to the carbonium ion for linear paraffins. For the branched paraffins the opposite situation is proposed with hydride abstraction to carbenium ions which then undergo beta-scission cracking as the favored mechanism.

A similar mechanistic study of just n-hexane cracking has been reported by Wielers and co-workers [123], in which they observed the same two possible cracking mechanisms. They invoked a cracking mechanism ratio (CMR) of the protolytic carbonium ion mechanism to the classic beta-scission carbenium ion mechanism and found that framework aluminum content and zeolite structur'e type influence the mechanism of cracking. The lower the framework aluminum content, the higher the reaction temperature and the smaller the zeolite pore size, the higher the CMR value. Thus high SiO,/Al,O, ratio ZSM-5 will tend to crack via carbonium ions with preferred olefinic and branched products, compared to low ratio Y zeolite which will prefer the traditional carbenium ion beta-scission. Even high ratio Y zeolite has been shown to favor the protolytic mechanism [ 1841. It was concluded that the classical beta-scission mechanism is favored by two adjacent acid sites and low temperatures.

The cracking of n-hexane over steamed ZSM-5 appears to be in sharp contrast to the above results. Kazansky and co-workers have proposed that Lewis sites on the non-framework aluminum species initiate the reaction by dehydrogenation to form more reactive 1-hexene, which then cracks at the framework Bronsted sites [154].

72

Finally a model reaction with cyclohexene has been shown to be extremely useful in determining high acid site density and high acid strength in zeolites. Cyclohexene cracking affords cyclohexane (via bimolecular hydride transfer), methylcyclopentene (via unimolecular isomerization) and methylcyclopentane (via a combination of both). Comparison of the relative concentrations of products can therefore be used to measure the hydride transfer properties of a catalyst [143].

7. CONCLUSION

A complex series of reactions occurs when a hydrocarbon molecule encounters a zeolite. These proceed via carbocation intermediates and are dependent upon the nature and strength of the acid sites present on the zeolite surface as well as steric factors. Modifications to acid site strength and density lead to changes in catalyst activity, stability and selectivity.

8. ACKNOWLEDGEMENTS

The authors would like to acknowledge the help of Jan Nieman, Jan Roelofsen, John Pearce, Leo Moscou and Dennis Stamires for helpful suggestions on the manuscript.

9.

1

2

3

4

5

6

7

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83

CHAPTER 3

COMPLEXITY IN ZEOLITE CATALYSTS: ASPECTS OF THE MANIPULATION, CHARACTERIZATION AND EVALUATION OF ZEOLITE PROMOTERS FOR FCC

D.E.W.VAUGHAN

Exxon Research and Engineering Company, Rt. 22 East, Annandale, N.J., USA, 08801.

1. INTRODUCTION

FCC catalysts are the most complex catalysts available in the market place today, amounting to ceramic micro-spheres (40 to 120p) comprising aggregates of 0.1p to 2p crystals of one or more zeolites, 1 to 5p kaolin crystals, sub-micron magnesia or alumina phases, all cemented together with low molecular weight polymers of silica, alumina, or both. The catalyst has been repeatedly modified to incorporate functionalities additional to the primary one of cracking high molecular weight crude oil molecules. In addition to the primary faujasite promoters, the current catalysts may include matrices of specific chemistry and pore size distribution to pre-crack very large molecules into the size range able to diffuse into the primary promoter; secondary promoters able to crack the low octane linear paraffins exiting the primary promoter into lower molecular weight olefins; nickel and vanadium traps or passivators; SOX transfer agents and CO promoters. All of these components are optimizable for a particular catalyst aimed at converting a specific feedstock into a targeted product slate in a defined FCC unit. This chapter deals with some of the controllable characteristics of the zeolite components, possible zeolite problems, and the methods for understanding them.

All primary cracking catalyst promoters need to have large pores (> 7A). The present family exclusively comprises one or other forms of exchanged or modified faujasite, X, Y , US-Y, LZY-210, or other variously dealuminated and exchanged forms, several members of which are shown in Figure 1. Secondary promoters are added to further change the converted product distributions, and these may be larger or smaller pore materials depending on the feed and desired product distribution. Larger pore promoters, such as pillared clays or other large pore pillared layered or amorphous structures [l] may be used as a primary promoter or added as a co-promoter to pre-crack more of the heavy fraction of the feed into faujasite convertible molecular weigh range reactants [2,3]. This group of materials is now very large, having controlled pores in the range from about log, to 25A with pillars of Si@, ZrO2, Ti@, Al2O3, and various cation modified variants, together with a diversity of mineral and synthetic sheet structures. Although very large pore zeolites such as VPI-5 [4] and "cloverite" [5] are now known, they do not have the thermal or hydrothermal stability to be usefully used in FCC

84

catalysts at this time. However, they do indicate possibilities for the future and that previous perceptions of crystalline pore sizes limits were erroneous. (Extensive computational studies of theoretically possible zeolite structure is presently being persued by several groups, giving strong impetus to synthesis programs to make very large pore materials.) Possible alternatives to faujasite types include beta [6] and PSH-3[7] /MCM-22 [8], both larger pore zeolites having the requisite hydrothermal stability and catalytic activity for use in FCC.

Smaller pore zeolites may be used to manipulate and control the post-faujasite cracked products by cracking the linear and single branched gasoline components into Q and lower hydrocarbons having a high olefin content. The first wave of interest in secondaq promoters followed the 1973 oil crisis, when it was thought that because of the increased crude oil prices more of the FCC products would be directed to higher valued petrochemicals. Inflation soon lowered the effective oil price and interest in FCC as a petrochemical source waned until the late 1980's when a decreasing volume, but higher octane, gasoline market coincided with a rising petrochemicals market, particularly for those components useful for making high octane additives such as alkylate and MTBEETBERAME.

This paper describes some of the zeolite characteristics which influence activity, selectivity and variability. Whereas some zeolites, such as X, Y and A, are readily made with high levels of purity and reproducibility in large quantities using relatively simple manufacturing methods, others are difficult to make in a like manner. Ostensible minor manufacturing changes may produce significant differences in a zeolite product which may only be recognized when tested as a catalyst. Subtle structural and compositional differences may be difficult to characterize, often resulting in debates over what is real and what is imaginary. Most catalyst manufacturers are periodically confronted with seemingly identical products which "do" and "do not" function as designed. Hopefully this review will provide some awareness of the complexities of zeolite catalysts, particularly as interest expands beyond zeolites of the faujasite variety to the rapidly expanding opportunities of the rest of the "molecular sieve" world.

2. CATALYST CONFIGURATION

As presently configured fluid catalytic cracking is controlled by a multi-functional catalyst system which, in addition to cracking gas oil molecules to lower molecular weight fractions, controls the release of sulfur as H2S, passivates, captures or otherwise controls metals activity, promotes CO combustion to C02, and possibly cracks gasoline range low octane linear paraffins to lower molecular weight LPG and olefinic molecules. A single composite catalyst has the problems of component compatibility in fabrication, physical properties, activity maintenance and differential deactivation. The alternative approach of using multiple mono-functional catalysts in the FCC unit is presently favoured, as reflected in commercial practice of adding separately, and controlling the activity independently, the CO promoter, Ni passivator and ZSM-5 cracking co-promoter. In this way the selected activity function may be turned on and off by degrees by simply adding more or less of a particular component as required for a particular feedstock. It is also a simpler problem to match matrix and one active component than it is to optimize many functional components in a single matrix. In the former case, compatibility in density and particle size between monofunctional FCC microspheres is seen as desirable to maintain catalyst homogeneity in the unit. However, it may also be possible to use differences in particle physical properties to further refine the system to accommodate differential catalyst deactivation rates. One such approach proposes incorporating a high coke resistant co-promoter, such as ZSM-5, into a high density matrix (titania, zirconia) to increase its residence time in the reaction zone and reduce it in the regenerator, or 'vice versa' [9]. The reverse should be possible with a high activity but fast coking zeolite such as mordenite. Attrited "fines" are a problem with soft catalysts, but the addition of a once through

85

rapidly deactivated small particle component is also a method of manipulating the unit [lo]. Clearly both compatibility and incompatibility can be exploited and utilized in FCC catalysts and unit designs and optimizations.

2.1 Primary Promoters

Conventional Y type faujasite, the dominant primary promoter in FCC, is easy to make reproducibly over a range of compositions, is readily exchanged without difficulty, and has high thermal and hydrothermal stability. Its three dimensional, large pore, 12-ring channel system is not readily blocked and in many regards it may be viewed as an ideal catalyst. One notes however, that faujasite, as used in FCC, is a family of materials having considerable structural and compositional variability, and includes many materials subjected to major post- synthesis chemical re-synthesis, often referred to as secondary synthesis [ 111. The diversity of these promoters has been expanded by the as-synthesized pure and intergrown faujasite variants. These are listed in Table 1 and illustrated in Figures 1 and 2. Faujasite can be visualized as being built from the sheet of connected sodalite cages, shown on the left of Figure 1, by rotating each successive added sheet by 60' before connecting it to the previous sheet - this is described as analogous to cubic stacking of spheres in which the layers can be said to have an ABC stacking sequence. An alternate mode of stacking is to connect the sheet to its mirror image - resulting in a hexagonal stacking mode having an AB sequence, popularized as "Breck Structure-6'' or BSS. In contrast to the faujasite structure, which comprises 139 diameter cages connected through 12-ring 89 windows, BSS can be viewed as a 12-ring channel structure of diameter varying between 8 to 159 with connected (12-ring window, 89 diameter) side cavities 109 in diameter. Depending on specific crystallization conditions it is clear that these stacking modes can switch layer to layer, and recent research has shown the existence of not just the faujasite (X, Y , ECR-4/32) forms but a whole family of possible structurally differentiable materials, related to one another as a function of the amounts of faujasite (cubic stacking-cp) and BSS (hexagonal stacking-hp) units. The silicoaluminophosphate faujasite analog, SAPO-37 [12], is analogous to zeolite X, and could also be included in this list as it too is of interest in FCC [13]. However, SAPOs are complex in that their activity depends on the specific nature of variable silica distributions in the crystals [14]. At very low silica content the Si randomly replaces A1 in framework positions imparting low excess charge (acidity) to the frame-work, but with increasing Si content it begins to form silica "islands", or micro-domains, within the SAP0 crystals, which have lower acidity.

Table I Variants of faujasite and related primary FCC promoters.

X Y

ECR-4 ECR-32 c s z - 3 c s z - 1 ZSM-2* ZSM-3* ZSM-20 ECR-30 ECR-35

.-

3.0 to lo 1.5 to 3.5 1.5 to 3.5 1.6 to 2.1 2.8 to 4.0 3.5 to infinity 3.0 to 15 2.0 to 12

CP CP CP CP

CP W. 4-10% hp distorted cp;<lO%hp similar to ZSM-3

40cp/60hp+ cp CP W. -20% hp

100%hp to 70hp/30cp block cphp intergrowth

__-_______--____________________________--_ *Very thin crystals and difficult to fully characterize by any method at present.

86

layer iype 1

layer type 2

ECR-30 (BSS)

Figure 1: Depending on how sheets of double 6-ring connected sodalite cages are stacked, two readily identifiable "end member" structures can be built - faujasite and Brecks Structure-6 (ECR-30) having distinctly different pore structures. Mixing of the stacking order results in a family of intergrowth structures having different activity and stability properties.

Hydrothermal instability is another problem with ALPOs and SAPOs, as both Al and P readily form hydroxylated species at high temperatures in the presence of steam [26]. In all probability extended research will demonstrate that the SAPO-37 compositions also have the structural intergrowth complexity of their aluminosilicate counterparts. With the exception of X, all these materials may be dealuminated by well known secondary synthesis methods, described in [l l] , using steam, steam plus leaching of detrital aluminum ions, high temperature treatments with Sic14 in the vapor phase or at low temperatures in non-aqueous solvents [27], and low temperature framework "exchanges" with (NH,)SiF6 [28].

5

4

sk, 3

2

1

ECR-32

i

ECR-30 ECR-4 ZSM-2/3

ZSM-20 csz- 1 csz-3

Y

X

FAU BSS

Figure 2: Relationships between the FAUBSS promoters as functions of composition (Si/Al ratio) and relative intergrowths of the two structural forms.

The catalytic complexity of these different dealuminated compositions has been extensively discussed by many authors [29,30,31,32 and Table 1 refs.]. An important property of the steam deactivated forms is the secondary pore structure which evolves within the crystals

87

as a result of selective steam destruction of the crystal lattice [33]. This meso-strumre within a micro-structure may be an important contributor to the conversion and selectivity towards large hydrocarbon molecules, representing a unique form of secondary promoter. As detrital alumina fragments (Lewis acidity) are thought to promote polymerization of multi-ring structures to "coke" in these meso-pores, treatments to remove the alumina detritus (by reaction with acids, EDTA, etc..) are a recent development. The structural issues of intergrowth and mesopore influence on activity and selectivity are still largely unexplored in a systematic way. Many of the intergrowth materials (ZSM-2,3,20; CSZ-1Q) have a distinctive platelet form, and the morphologies which gives short diffusion path-lengths may be more important than pore size and shape. Several morphologies of faujasite/BSS type materials are shown in Figure 3.

2.2. Secondary Promoters

In comparison to faujasite the zeolites used as secondary promoters in FCC catalysts are far more complex and suffer from problems ranging from poor structural perfection and reproducibility, to pore blockage by coke or template degradation products deposited in relatively restricted pore systems. Their synthesis procedures, with the exception of ZSM-5, have not been thoroughly established, and in some cases the zeolites are difficult to make in the laboratory. For some zeolites the best, though unsatisfactory materials from the viewpoint of phase and elemental purity, may be the mineral forms. Important facets of the characterization and evaluation of these secondary promoter zeolites is of particular interest, with the focus on the physical, chemical and structural diversity and manipulability.

Zeolites of interest as secondary promoters, in addition to the ZSM-5 currently used commercially, include several small, medium and large pore materials listed in Table 2, where the priority references are indicated, and morphologically illustrated in Figure 4. Numerous additional patents are issued in which variations focus on specific preferred zeolite chemical compositions, levels of co-catalyst and associated matrix and additional zeolites. A generalization of catalytic functionality in the small and intermediate pore zeolites is that they crack the low octane n-paraffins into shorter chain paraffins and olefins. This mechanism for increasing octane values has been practiced commercially in reforming for many years under the Mobil "Selectoforming" label [341. The smaller the pore, the lower the carbon number of

beta(BET) L (LTL)

mordenite(M0R) offre tite( OFF) gmelinite(GME)

femeri te(FER) erioni te(ER1)

ZSM- 5 (MFI)

PSH-3/MCM-22

12-ring, channels 12-ring ,channels 12-ring, channels 12-ring, channels 12-ring, channels 10-ring, channels 10-ring, channels &ring, cavities

unknown

79 1351

6.59~79 [371 6.59 [381 79 [391

3.69x5.18 [391

79 [361

5.39x5.69 [401 4.29x5.49 [411

88 r421

88

Figure 3: The typical Na-Y (from which all present FAU FCC promoters are derived) morphology comprises mixed crystals of equant octahedra in numerous twinned, distorted and fragmented forms - intergrown, flattened, squashed, truncated - as shown in (a). When grown in a mixed K,Na-gel the Y crystals form flattened octahedra as shown in (b). At high levels of Cs added to a Na-gel very thin platelets only tens of unit cells thick crystallize (c) (CsZl), but at low levels of Cs, thick platy crystals (d) form (CSZ-3). CsZ-1 is similar morphologically to ZSM-3 (e). High seeding promotes very small crystals (f). Mark= lp.

89

Figure 4: Examples of secondary promoters include beta (a); gmelinite (b); fenierite (c); mordenite (d); offretite (e) and mazzite (0. With the exception of beta, a multi-channel structure, all these materials have their principal diffusion path along the length of the crystals. Marks= 1p.

90

the products. 1Zring zeolites (offretite, mordenite) and 10-ring zeolites ( ZSM-5, femerite) tend to make iso- and normal butanes and butenes, and 8-ring zeolites make C2 and CJ. (In Selectoforming with hydrogen present the 8-ring zeolite erionite produces LPG. However, the use of erionite may be restricted in the future because of reported asbestos like health problems [43], unless a platelet form can be synthesized.) The issue of creating higher octane molecules in the FCC unit, with the help of secondary promoters which aromatize LPG (as in the BP/UOP Cyclar process [MI) or isomerize straight chains, is beginning to be revisited [45].

in gasoline demand resulting from higher prices and more efficient automobiles, indicated a need to focus more on the possibilities for higher valued petrochemical feedstocks off the FCC unit. Several zeolite catalysts were demonstrated to be effective in selectively converting gasoline range hydrocarbons to ethylene, propylene and butylenes. As inflation lowered the effective price of oil, the petrcchemical interest waned and the interest in Octane increased, beginning a period dominated by the introduction of US-Y Octane FCC catalysts. The continued octane demand into the 1980s stimulated the introduction of FCC ZSM-5, which convert low Octane gasoline range n-paraffins to isobutylenes, an important alkylation feedstock, and the resurrection of multi-zeolite FCC catalysts. The present situation is one where gasoline demand is decreasing and petrochemical demand is increasing. It is further complicated by the increasing demand for petrochemical fuel components such as alcohols, ethers and alkylation feedstocks as components in new, more environmentally friendly, reformulated gasolines and alternate fuels. An earlier review [46] discussed the different specific selectivity effects in FCC of several zeolite combinations, but more recent research, and the discovery of numerous new zeolites, has greatly expanded the interest in multi- promoter catalysts, including the evaluation of mple zeolite promoted FCC catalysts [47] which may include isomerisation and olefin oligomerization functionalities in addition to selective cracking properties. Refinements include catalyst particle size and density control to manipulate residence times in the reactor.

In the 1970s, the increasing price of crude oil, and the anticipated attendant decrease

Hitherto most interest has been on enhanced octane and reformulated gasoline components, but the need to process increasing levels of heavier feedstocks stimulates the search for secondary promoters to increase the conversion of heavy cycle oils. Although much effort has aimed at improved matrix activity for this purpose [48,49], improvements in the stability of pillared clays [50] and the gathering momentum in larger pore zeolite synthesis will increase the use of controlled larger pore materials having increasingly diverse chemistries [5 11 as co-promoters [52]. Numerous pillared layered structures [ 11 have possible applications in this area. Particularly interesting are the silicic acids extensively investigated by Mobil, some of which can be pillared [53], and others which seem to be tunnel structures of random meso-pore sizes [54] similar to the aluminosilicate mineral imogolite [55]. Optimized FCC catalysts may well be combinations of meso-pore materials to pre-crack the large multi-ring "resid' molecules, a faujasite to crack the main body of the feed to gasoline range products and a medium pore material, such as ZSMJ, to crack the straight chain molecules into low molecular weight olefins.

3. SYNTHESIS CONTROLLABLE VARIABLES

3.1 Zeolite Morphology Control: Crystal Size, Shape, Dispersion.

In a process such as FCC, in which the feed contains a wide range of molecular sizes and types, some reactions will take place on the surface and others inside the zeolite crystal. Promoter morphology may then be an important variable in controlling activity and selectivity in addition to the physical properties of a catalyst. Significant attention has been given to increasing the surface to volume ratio of faujasite by reducing crystal sizes by manipulation of

91

Figure 5: Few zeolites have the demonstrated morphological manipulability of ZSM-5, (a-c), or LTL, (d-f). (A function possibly related to the intensity of research efforts.) These scanning electron microscope micrographs show some possible variations in crystal size and shape as functions of gel compositions and reaction conditions. Marks= lp, except (c) which is 25p.

92

synthesis conditions, and to directing the shape to forms which reduce diffusion path-lengths and improve the packing efficiency of crystals within the catalyst agglomerates. Mixtures of crystal sizes and shapes improve the packing properties of the catalyst components and therefore enhance the physical properties [56]. Controlling the morphology of a co-promoter - not always a controllable variable in zeolite systems- is an additional feature that may contributes to catalyst optimization.

ZSM-5 is the most intensively investigated zeolite in the co-promoter group, with controllable crystal sizes from ~ 0 . 1 ~ to >loop, as shown in Figure 5 ,which illustrates several shapes and sizes of ZSM-5 and Linde-L (LTL). Agglomerates are common when crystal sizes are below 0 . 1 ~ and an important issue is whether such agglomerates comprise multiple single crystals or are crystalline mosaics. Examples of these are shown in Figure 6. Depending on the dimensionality of the pore system, the former may give much greater accessibility to reactants than the latter, particularly when the individual crystals are well dispersed in a matrix; however, the latter may have higher stability.

Figure 6: Small crystals, having the appearance of individual crystals in SEM micrographs, in the higher resolution electron microscope (HREM) may be shown to be agglomerates of nano- crystals (a) or crystal mosaics (b).

The issues of size and shape of zeolite crystals are not important ones so long as the pores are large and the reacting molecules are small. Furthermore, at the high temperatures characteristic of many catalytic reactions, the molecular kinetic energy is high and the zeolite pore vibrational modes expansive, stimulating the diffusion of larger molecules (in flexible zeolite structures these "breathing" modes may promote diffusion). For the cracking of light gas oils in faujasite promoted cracking catalysts this is clearly the case, but with heavier feedstocks the inability of increasing fractions of the feed to enter the pores shows as decreasing conversion to lighter products. A much favoured approach in the industry is to add an active larger pore matrix to "pre-crack" the large molecules into the faujasite pore size range. Pillared clays are an example of a larger pore co-promoter achieving this result [57]. An alternative approach is to increase the surface to volume ratio of the faujasite and therefore increase the external area of the crystallites, increasing the reactive surface for large molecules [%I. This can be achieved either by mechanical means, or by developing methods to directly synthesize faujasites with such properties. Successful direct synthesis methods include growing small crystals by high seeding techniques [59]; by poisoning the surfaces of growing crystals at a small size by adsorbing large anions [58]; or by synthesizing platelet forms which have short channel lengths and therefore good pore blockage resistance [60]. Extreme cases of the latter for faujasite types are CSZ-1 [23] and CSZ-3 [22], which may comprise platelets up

93

to one micron diameter but only twenty unit cells thick (2008). An early observation on faujasite having crystal sizes of about 500A was that it had abnormally rapid deactivation rates in the presence of steam, probably reflecting the preferential dealumination of the surface, or destruction of the crystal lattice by high rates of surface hydrolysis - a process hardly noticeable in large crystals [61]. Subsequent studies indicate that in high Si/A1 Y such degradation is less problematical [62], and that there are distinctive improvements in the selectivities for gasoline, light cycle oil, coke and light gases. Shorter diffusion paths allow gasoline and lower olefins to escape from the zeolite before hydrogen transfer can occur, improving gasoline octane and alkylate feed quality.

them are not three dimensional pore systems, and they tend to have high aspect ratios (ie. length to diameter ratio). The single channel 12-ring zeolites, mordenite, mazzite, offretite and L all often have lath, needle or long prismatic shapes with the long axis parallel to the channel. Catalytically such pore systems are “shut-down‘’ by minor coke deposits and the challenge is to synthesize these materials in a form having a short channel length. This has been best demonstrated for zeolite L, as shown in Figure 5, and for mordenite. The 10-ring zeolite ferrierite is difficult to prepare as crystals having short channel lengths, but ZSM-5/11, the subject of numerous synthesis studies, can be readily made in sizes from 400A to 100 microns (Figure 5 ) with different aspect ratios. Mechanical methods of crystal diminution are not particularly effective for crystals having a lath or needle morphology, although some crystals of this kind can be reduced by “cryo-chopping“ in a liquid nitrogen cooled “coffee bean“ type of mill [631 (ie. make the long crystals brittle before chopping them).

3.2 Chemical Non-homogeneity.

The use of other, non-faujasite, zeolites presents different problems because most of

The variability of composition across a crystal, and specifically between surface and bulk, may be an important parameter in interpreting catalytic results. Only in the case of large crystals has this parameter been fully investigated. Various geochemical studies have established that compositional zoning is a common occurrence in minerals and studies on large crystals of ZSM-5 have confirmed that it readily occurs in relatively rapidly grown large synthetic zeolite crystals, illustrated in Fig. 7 [64]. In smaller crystals, more characteristic of those used in FCC and other catalysts, the evidence is sparse, but available analyses indicate homogeneous compositions. However, the development of more refined analytical tools is progressively improving the resolution of micro-probe analyses to allow meaningful

unit cell

Figure 7: Example of the varying SVAl ratio contours across a section of a large ZSMJ crystal (96 Si+Al atoms/UC) [59]. Alternate models of compositional variation are shown in Figure 8.

94

characterization of the chemical variability of less than one micron crystals. An example of this is the chemical differentiation of the FAU and "Breck-6" components of ZSM-20, which show the latter sections to have marginally higher Si/Al ratios than the former [65]. Analyses of this kind would be particularly desirable for analyzing the compositional profiles of FCC promoters made by various secondary synthesis methods - materials which often show variable catalytic properties as a function of specific process conditions (batch to batch variation) for supposedly identical catalysts. An important question is the extent to which chemical profiles can be controlled, either in the synthesis process or by secondary synthesis chemical reactions. This is important if the relative acidities of bulk and surface influence product selectivities. It has been demonstrated for ZSMJ that the Si/Al content of successive growth layers can be manipulated during crystal growth by the addition or subtraction of aluminum compounds in the synthesis liquor [66,67]. For other zeolites, this kind of "in process" compositional manipulation may only co-crystallize separate crystals of lower Si/A1 ratio, or possibly initiate the crystallization of impurity phases. The secondary synthesis route may be more generally applicable, as de- aluminating complexes may be tailored to be larger than the zeolite pores (an ion sieving effect), facilitating A1 extraction only from surface sites.

Depending on the function of the zeolite, the desired surface and bulk reactivities may be quite different. For a shape selective co-promoter such as ZSM-5 a high silica, low acidity, exterior shell would minimize non-selective surface reactions and higher Si/Al interior compositions would maximize selective reactions in the interior pore system. This surface passivation can also be done by reacting the surface acid sites with silanes, as demonstrated for xylene isomerisation catalysts based on ZSM-48 [68], a process which reduces transalkylation and coking reactions on the non-selective crystal exteriors. If however, pre-cracking of large molecules is important, a higher activity crystal surface will be desirable. Detailed evaluation of the hexane cracking patterns as a function of Si/Al ratio on ZSM-5 shows distinctive trends as functions of A1 content and cracking temperature [69], indicating specific reaction fine-tuning opportunities for co-catalysts of this type.

Figure 8: Several chemical and structural modifications shown may have specific catalytic advantages, and have been extensively modelled, particularly in the auto-exhaust catalyst area.

3.3 Structural Homogeneity

One of the issues here is whether a zeolite is single phase or contains levels of impurities below the level of detection by routiie analytical methods (usually 5 to 10%). Another is whether it is possible to control the epitaxial overgrowth of one structure on another (not to be confused with intergrown materials), and to obtain reactivity advantages by doing

95

so. Such a composite crystal could impart important additional ingress-egress selectivity on a catalyst, and possibly be an effective method for protecting the inner catalyst from poisons such as coke or metals precursors. These "egg" or "cherry" models (Figure 8) have been extensively modelled by reaction engineers [70], particularly for the design of auto exhaust catalysts. The first observation of this phenomenon in zeolites was the overgrowth of mazzite on a crystal of ECR-1, shown in Figure 9 [71].

Subsequently, in the same system of structures, mordenite overgrowths were reported to have been observed on crystals of mazzite [72]. In both these composites the overgrowths do not modify the pore systems as the second phase overgrows parallel to the pore direction. However, in the offretite-erionite group, and in structures related to mordenite [73] and to ferrierite [74] they do modify the pore structures, resulting in variable catalytic and sorption properties for apparently identical zeolites obtained from different sources. Evaluating the catalytic properties of these kinds of zeolites therefore requires an informed selection of samples and necessarily testing more than a single catalyst. Molecular selectivity can also be changed by various sample pre-treatment methods. Controlled chemical vapor deposition (CVD) techniques are well established for depositing overlayers of one material on another and have been used in several zeolite systems [75]. More extensively used is the adsorptive

Figure 9: A crystal of ECR-1 showing an overgrowth of mazzite on the right hand side crystal face and an intergrowth layer (center). The main 12-ring channels are clearly seen. The structure model is shown above. See Figure 10 for models of alternate layer permutations.

96

silanation by reactive molecules such as silanes. Pore mouth blocking has been particularly successful using variations of this method [76,77], but it may be difficult to control, particularly when the catalyst is subjected to a reactive (deactivating) environment. These methods may either narrow the pore mouths of the zeolite, increasing its selectivity, or passivate the external surface to reduce the activity for non-selective surface reactions. In FCC a composite catalyst could contain different modified faujasites to impart selectivities more characteristic of say ZSM-5, but having the cost and compatibility with zeolite Y.

3.4 Intergrowth Structures

The widening use of high resolution electron microscopy is revealing the complexity of zeolite crystal structures in a way that could only have been imagined a decade ago[78]. Many zeolites having catalytic and sorption properties incommensurate with their idealized structures are now better understood in terms of intergrowths of multiple structures [79]. The faujasite - "Breck-6" series of materials [80,81] and zeolite beta [6] are examples of long standing problems rationalized in this way. In some cases new modelling approaches allow quantification of such intergrowths on the basis of X-ray diffraction patterns [82], a simpler and lower cost technique than HREM. MASNMR in some cases may also be a useful rapid analysis method. In the case of offretite, 13C-MASNMR can be used to track the tetramethyl- ammonium (TMA) cation trapping sites, the presence of additional 13C peaks characteristic of different TMA sites indicating intergrowths that may not be seen in a standard powder X-ray diffraction pattern [83]. Although this method can be readily used for simple templates like TMA and TEA (tetraethylammonium), with less stable or more complex molecules, which may degrade or convert to other species during the synthesis process, the spectra become too complex for analysis. An example of TMA-offretite with and without intergrowths is shown in Figure 10; the former sample has a high n-hexane sorption capacity and the latter a low capacity reflecting blockage of the 12-ring channels.

6 C %+%k%k%*

Figure 10: 13C-MASNMR spectra for un-blocked (bottom) and blocked (top) offretites. In the former the peak at about 58.5 ppm is characteristic of TMA in the gmelinite side cage and at 57.2 ppm TMA in the 12-ring channel. The upper spectrum shows the first peak plus a series of peaks between these two, indicating a series of intermediate sized cages atypical of the offretite structure and indicative of intergrowths.

The nature and extent of intergrowths may vary within a given material but general types can be recognized. The following have been observed in materials of interest in FCC:

97

1. RANDOM: Occasional and incidental - as seen in CSZ-3, mordenite , offretite, gmelinite, femerite and ZSM-5. 2. ORDERED: Systematic and seemingly reproducible - observed in zeolites which seem to be specific mixtures of two forms; examples of which are beta (50/50) and

3. LARGE NUMBER OF SPECIFIC COMBINATIONS: New structures may be defined as ordered permutations of two or more structural units or orientations of the same unit seen in the pentasils [84] (ZSMJ/ 11) and mazmorites [85](ECR-l, mazzite, mordenite).

The first rarely show up in X-ray diffraction patterns but are seen in HREM micrographs and as streaking in electron diffraction patterns. In some cases with larger crystals (eg. gmelinites) the blocks of intergrowths may be large enough to be seen in low magnification microscope [86]. The single channel materials may have low sorption capacities for hydrocarbons, frequently varying from preparation to preparation. Figure 1 la illustrates this phenomenon for CSZ-1. The second type, because they give reproducible levels of intergrowths in a similar mode, and therefore reproducible X-ray diffraction patterns, often are mistaken for novel pure materials. HREM shows the true mixed structures of beta in Figure 1 lb and ZSM-20 in 1 Ic. The third type may only be initially recognizable in the HREM mode. Once defined they can be readily modelled and the X-ray diffraction pattern simulated either by adding the end member patterns in the required relative amounts [79], or using more sophisticated computational methods which add contributions layer by layer [81]. These mixed materials are best evaluated by mapping the complete range of theoretical X-ray diffraction patterns for the possible structures and comparison matching the experimental against the theoretical patterns. Examples of the mazmorites are given in Figure 12, together with the schematic of ECR-1 shown in real life in Figure 9. Failure to adequately define these materials and their synthesis parameters will result in a high level of variability, batch to batch, in the catalytic properties of apparently identical materials. In a demanding process like FCC the expected selectivity changes may not materialize because of structural degradation (some intergrown materials are structurally strained because of minor lattice incompatibilities at the interfaces [87] resulting in inferior steam stability) or unexpected poor coke selectivity.

ZSM-20 (40/60).

4. STERIC HINDERANCE - A CHANGING VARIABLE.

Pores within zeolites, being of molecular sizes, can be readily blocked in a wide variety of ways. The smaller the pore the more readily the pore diffusion characteristics can be modified. In FAU the pore is large enough to obscure partial blockage effects, but in the large number of smaller pore potential secondary promoters the effects may be dramatic. Probably the best example is the conversion of 4A (Na-LTA; or the higher silica ZK-4) into 5A (Ca- LTA) as a function of the exchange of sodium ions by calcium ions. The former does not sorb n-paraffins whereas the latter does as the number of blocking cations are reduced (one Ca2+ replacing two Na+). Similarly, pores are modified by small amounts of polar molecules, such as H20, NH3, CO, etc., solvating the cations and effectively increasing their size [88]. In some cases the solvated or complexed cation may move position within the channel or cage, on solvation, from a non-blocking site to a blocking site. In catalytic reactions "coke" is a particularly effective pore block, but so too may be other reaction products such as water and olefins, which may move cations into pore blocking positions in the structure or actually fill the pores, as seen in the methanol to gasoline conversion on ZSM-5. Some preferred ZSMJ xylene isomerization catalysts require small controlled amounts of coke to be deposited in the pores to optimize selectivity[89]. Other optimized ZSM-5 xylene isomerization catalysts contain compounds of phosphorous, zinc, magnesium or boron partly blocking the pores or titrating acid sites at pore mouths on the crystal surfaces [90,91]. In some cases salts, derived from impurities in the feed or from catalyst regeneration processes, may be deposited in the pores and cause the catalyst to be non-regenerable. (Deposition of these "obstructors" selectively on

98

the outside of the crystals will reduce surface reactivity.) All of these blockages may be effective at low concentrations, being more pronounced for zeolites having one or two channel pore systems. Detection may be feasible only in sorption or catalytic experiments.

Figure 11: Occasional and incidental intergrowths of BSS in FAU are shown in a crystal of CSZ-1 (a), but in almost equal amounts in ZSM-20 (b). Beta (c) comprises equal amounts of two enantiomorphic zeolite components.

99

Figure 1 2 The third kind of intergrowth structure mixes a small number of layers in specific repeatable permutations. Simple examples for ECR-1 and the mazmorites are schematically shown.

100

A special case of variable pore blockage arises in zeolites made with organic templates which must be burned off before the zeolite can be used as a catalyst promoter. Inadequate "bum-off'' procedures leave residual material in the pores, and excessively high temperature "bum-off'' may cause partial structural degradation, often promoted by small amounts of steam produced by the template oxidation process. This is most pronounced for zeolites in which the template is trapped in small cages bounded by six or eight ring cages, such as offretite and mazzite [92]. Lower Si/Al ratio materials are more susceptible to partial structure collapse than higher Si/Al zeolites; gallium substituted zeolites are even less stable.

5. CHARACTERIZATION APPROACHES.

Detailed zeolite characterization is essential in order to define key catalyst characteristics and to understand variabilities in catalyst performance. Useful protocols for the characterization of zeolites have been describedin detail elsewhere [93]. Particularly powerful are some of the old methods refined with new instruments. These may be as simple as low pressure sorption or as complex as synchrotron diffraction methods [94]. Silicon, aluminum and proton MASNMR have become particularly useful tools in the characterization of de-aluminated faujasitic promoters[94]. Catalytic results are a valid tool in the characterization of catalysts [95], but they are derivative and supportive and not primary or direct. They do not give specific structural information and therefore do not define the source of structure related problems, although variable catalytic results may be an excellent indicator that problems do exist in a particular material. Too often extensive catalytic studies are initiated on the basis of a container label name, and time and effort is mis-spent investigating "non-characteristic" samples of a zeolite. The pitfalls are numerous and are exemplified by the analyses of commercial H+- mordenite (MOR) and H+-ferrierite (FER) from different vendors, shown in Table 3. The X- ray diffraction patterns were characteristic of the type materials and the small crystal sizes indicate no expected diffusion limitations. Much more extensive analyses are therefore needed to explain the differences in sorption properties, which are probably related in some way to structural defects. To effectively evaluate the properties of many zeolite catalysts used as secondary promoters, several samples and variables need to be included in the study [67]. Careful preliminary analyses of the catalysts are rewarded with reproducible and comprehensible catalytic results.

Table 3 Variations in the properties of mordenite and ferrierite from different catalyst vendors.

ZEOLITE VENDOR Si/Al CRYSTAL SORPTION CAPACITIES SIZE SHAPE H2O# n-C~H12'

I.L wt.% wt.%

MOR A 7.9 <0.1 euhedra 18.2 4.7 MOR B 10.0 0.1x0.5 prisms 15.2 7 .O MOR C 10.5 0.2x0.6 prisms 17.0 8.2 FER D 6.2 <0.1 euhedra 10.0 2.1 FER E 8.4 <0.1x0.5 plates 9.0 1.3 FER F 8.4 <0.1x0.6 plates 6.6 0.6

#35% relative humidity; 25OC. *45 torr; 25OC.

purity, crystallinity (or yield), and to define crystalline contaminants; scanning electron Necessary and minimum analyses are always X-ray diffraction to establish phase

101

microscopy to establish crystal size, shape, composition (with microprobe) and impurities at low concentrations (4%); bulk analysis to establish zeolite stoichiometry and flag possible contamination; and sorption measurements with molecules relevant to probing the pore size of the zeolite (n-hexane is a good standard for 8-, 10- and 12-ring structures; neopentane or cyclohexane for 12-ring and larger pores). If crystal sizes are so small (< 20 to 30 unit cells, shown by electron microscopy) as to show appreciable peak broadening or even amorphous character, IR is a superior method for establishing sample crystallinity [96]. Other desirable analyses, depending on the nature of the problems, include electron diffraction, structure imaging, chemical profiling, IR, and MASNMR.

The main issue is that, as catalysts become more complex, their performance is increasingly related to structure -composition relationships. There are presently four or five distinctively different FCC faujasite promoters which are relatively simple compared to many actual and potential secondary promoters. With increasing use of the latter (recent patents showing great interest in ZSM-5, beta, MCM-22, Linde-L), routine catalyst characterization will have to become much more sophisticated to attain optimum unit performance.

6. OUTLOOK

FCC is viewed as a mature process, but its place at the core of refinery fuel conversion processes has ensured major research and development activity and spectacuIar evolutionary developments over the years in hardware [97] and catalysts [98]. The future will clearly continue the progress but with more focus on heavy oil conversion and petro-chemical synergies, rather than total focus on gasoline and octane barrels (the latter having decreasing importance as high octane MTBE is increasingly used as a gasoline oxygenate source in the USA and an economic octane booster elsewhere). The trend to increasingly complex multi- component catalysts and composite catalyst mixtures will clearly continue, providing increased feed flexibility and molecular product control to the refinery of the future.

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Ohlmann, H. Pfeifer and R. Fricke, Stu.Surf.Sci.Catal. 65, 171 (1991). 49.J.E. Otterstedt, Y.-M. Zhu and J. Sterte, Appl.Catal., 23, 140 (1988). 50.J.R. McCauley, U.S. Patent 4,952,544 (1990). 51.D.E.W. Vaughan, Proc.8th Intl. Zeolite Conf., Ed. P.A. Jacobs and R.A. van Santen,

Stud.Surf.Sci.Catal.494, Elsevier (Amsterdam), 95 (1989). 52.J. Sterte and J.E. Otterstedt, Appl.Catalysis, 3, 131 (1988). 53. I.D. Johnson, P. Chu and C.D. Kresge, U.S. Patent 5,008,481(1991). 54. J.S. Beck, C.T Chu, I.D. Johnson, C.T. Kresge,M.E. Leonowicz, W.J. Roth and J.C. Vartuli, US Patent 5,102,6431 5,110,572 (1992); Int.Pat.App1. (PCT) WO 91/11390 (1991). 55. V.C. Farmer, M.J. Adams. A.R. Fraser and F. Palmeri. Clay Minerals, El, 459 (1983). 56.D.J. Cumberland and R.J. Crawford, "The packing of particles.", Handbook of Powder

Technology, Elsevier Press (Amsterdam), 4, (1987). 57.F. Figueras, Catal.Rev Sci.Engrg., 3 , 4 5 7 (1988).

Patent 4,740,292 (1988)

(199 1).

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58.M. Farcasiu and T.F. Degnan, Ind.Eng.Chem.Res.,Z, 45 (1988). 59.G.C. Edwards, E.W. Albers and D.E.W. Vaughan, U.S. Patent 3,755,538 (1973 ) 60.G.C.Edwards, D.E.W.Vaughan and E.W.Albers, U.S. Patent 4,175,059 (1979). 61.K. Rajogopalan, A.W. Peters and G.C. Edwards, Appl.Catal., a, 69 (1986). 62.M.A. Camblor, A. Corma, A. Martinez, F.A. Mocholi and J. Perez-Pariente, Appl.Catal.,

63.Tekmar Inc., Cincinnati, Ohio. 64.R. von Ballmoos and W.M. Meier, Nature, 289,782 (1981). 65.D.E.W. Vaughan, M.M.J. Treacy and J.M. Newsam, NATO AS1 Ser.B 221, "Guide-lines

for mastering the properties of molecular sieves.", Ed. D. Barthomeuf, E.G. Derouane and W. Holderich, Plenum Press (London), 99 (1990).

66.L.D. Rollmann, U.S. Patent 4,088,605 (1978). 67.N.Y. Chen, N. Miale and W.J. Reagan, U.S. Patent 4,112,056 (1978). 68. P. Ratnasamy and S.K. Pokhriyal, Appl.Catal., 55, 265 (1989). 69. A.F.H. Wielers, M. Vaarkamp and M.F.M. Post, J.Catal., 127, 51 (1991). (For a theoretical view see: W.J. Mortier, Proc. 6th Intl. Zeolite Conf., Ed. D.H. Olson and A. Bisio, Butterworthsl 1ntl.Zeolite Assoc. (London), 734 (1984).) 70.A.V. Niemark, L.I. Kheifez and B.V. Fenelonov, 1nd.Eng.Chem. Prod.Res.Dev., a, 71.M.E. Leonowicz and D.E.W. Vaughan, Nature, 329,819 (1987). 72.F. Fajula, F. Figueras, C. Gueguen and R. Dutartre, U.S. Patent 4,946,580 (1990). 73.J.D. Sherman and J.M. Bennett, Am.Chem.Soc. Adv.Chem.Ser., m , 5 2 (1973). 74.R. Gramlich-Meier, W.M. Meier and B.K. Smith, Zeit.Kristallogr., 169, 201 (1984). 75.M. Niwa and Y. Murakami, J.Phys.Chem. Solids, 50,487 (1989). 76.R.M. Barrer and J.-C. Trombe, J.Chem.Soc. Faraday l , B , 2786 (1978); ibid 2798

55, 65 (1989).

439 (1981).

(1978). (See also ref. 61). 77.Y. Yan,'J. Verbiest, J. De Hulsters and E.F. Vansant, J.Chem.Soc. Faraday 1,& 3087

(1989): ibid 3095 (1989). 78.C.N.R.Rao and J.M. Thomas, Accts.Chem.Res., 18, 113 (1985). 79.G.R. Millward, S. Ramdas and J.M. Thomas, Proc.Roy.Soc., m, 57 (1985). 80.D.E.W. Vaughan, M.M.J. Treacy, J.M. Newsam, K.G. Strohmaier and W.J. Mortier,

Amer.Chem.Soc. Symp.Ser., 398, 544 (1989). 81.J.M. Newsam, M.M.J. Treacy, D.E.W. Vaughan, K.G. Strohmaier and M.T. Melchior, in

"Molecular Sieves v.l.", Ed. M.L. Occelli and H.E. Robson, Van Nostrand Reinhold (New York), 454-472 (1 992).

82.M.M.J. Treacy, J.M. Newsam and M.W. Deem, Proc.Roy.Soc., A433,499-520 (1991). 83.M.T. Melchior, D.E.W. Vaughan and A.J. Jacobson, 60th Colloid and Interface Symp.,

Atlanta, (1986) (unpublished). 84.G.T. Kokotailo and W.M. Meier, Chem.Soc. Spec.Publ.3, "Properties and Applications

of Zeolites.", Ed. R.P. Townsend, 133 (1980). 85.M.E. Leonowicz, D.E.W. Vaughan and K.G. Strohmaier, U.S. Patent 4,892,721 (1990). 85.G.T. Kokotailo and S.L. Lawton, Nature, 203, 621 (1964). 87.M.M.J. Treacy, J.M. Newsam, D.E.W. Vaughan, R.A. Beyerlein and S.B. Rice, "Analytical Electron Microscopy", Ed. D.C. Joy, San Francisco Press, 161 (1987). 88.R.M. Barrer, "Zeolites and clay minerals as sorbents and molecular sieves.", Academic

89.D.H. Olson and W.O. Haag, Am.Chem.Soc. Symp.Ser. 248,275 (1984). 90.W.W. Keading, J.Catal., B, 512 (1985). 91.J.C. Vedrine,A. Aroux, P. Dejaifve, V. Ducarme, H. Hoser and S. Zhou, J.Catal., z, 92.D.E.W. Vaughan and K.G. Strohmaier, in "Molecular Sieves v.l.", Ed. M.L. Occelli and

93.J.M. Thomas and D.E.W. Vaughan, J.Phys.Chem.Solids, 54,449 (1989).

Press (London), Ch. 7-2, (1978).

147 (1982).

H.E. Robson, Van Nostrand Reinhold (New York), 92-104 (1992).

104

94.D.E.W. Vaughan, M.M.J. Treacy and J.M. Newsam, NATO AS1 Ser.B 221, "Guide-lines for Mastering the Properties of Molecular Sieves.", Ed. D. Barthomeuf, E.G. Derouane and W. Holderich, Plenum Press (London), 99 (1990).

95.J.A. Martens, M. Tielen, P.A. Jacobs and J. Weitkamp, Zeolites, 4,98 (1984). 96.P.A. Jacobs, E.G. Derouane and J. Weitkamp, J.Chem.Soc.Chem.Commun., 591 (198 1). 97.A.A. Avidan, M. Edwards and H. Owen, Oil and Gas J., m, 33 (1991). 98.1.E. Maxwell and J.E. Naber, Catal. Letters, 12, 105 (1992).


CHAPTER 4

COMMERCIAL PREPARATION AND CHARACTERIZATION OF FCC CATALYSTS

GERALD M. WOLTERMANN: JOHN S. MAGEE: and STEPHEN D. GRIFFITIT

a The PQ Corporation, Conshohocken, Pennsylvania Catalytic Science Associates, Ellicott City, Maryland

' UOP, Des Plaines, Illinois

I. INTRODUCTION

When f i s t introduced commercially in the early 1960's, zeolite cracking catalysts contained either type X or type Y zeolite incorporated in several different matrix compositions. These matrices had evolved directly from the amorphous catalysts used in the 1940's and 1950's. The main purpose of a matrix in the early zeolite catalysts was to moderate the high intrinsic activity of the zeolite to prevent its rapid deactivation as a result of coke formation. Throughout the 1960's and well into the mid 1970'9, zeolite containing FCC catalysts were modified to take advantage of the increased conversion and liquid yield and low coke make afforded by catalytic cracking over zeolites. Demands on catalyst performance were modified to include increased gasoline octane, increased activity, and improved metals tolerance. Today new demands are requiring FCC catalysts to increase the yields of light olefins and branched isomers in reformulated gasolines. Such demands will naturally require significantly different catalyst compositions.

To meet these ever-changing yield requirements, which are often further defined in each specific refinery, requires a large number of catalyst formulations. At last count, more than 150 catalyst grades were said to be available to the refining industry (1). An equal or even greater number of potentially useful cataIytic materials exist in industrial and university laboratories. Many of these will no doubt make their way to commercial FCC units.

During the 1970's and 1980's, the "additive" era of the FCC catalysts evolved. Separate co-catalyst systems designed to control carbon monoxide and SO, emissions are now commonly used. In addition, the first small pore zeolite used in FCC catalysis was introduced. This material, a pentad, is effective in converting straight-chain gasoline-range hydrocarbons into C, and C, olefins while isomerizing gasoline-range olefins to isoolefins. Nickel and vanadium passivators are also

105

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commonly used in high metals residual feed operations either as additives or as an integral part of the catalyst particle. These additive systems are described in detail in Chapters 10, 13, and 14.

From the vantage point of the chemist and engineer preparing catalyst, entire new fields of preparative science opened in the 1980’s. Treated clays and gels no longer made up entire FCC catalyst systems. Rather the FCC preparative scientist could realistically be called an inorganic architect of new molecular specific structures with controlled active site strength and location. Procedures for the specific and reproducible placement of active sites have become available in the published and unpublished work of many catalyst preparation workers. These procedures enable the selectivity of the carbenium ion reactions catalyzed by new generations of FCC catalysts to be controlled. Further, new co-catalyst additives have broadened the scope of FCC catalyst chemists much beyond the chemistry of silica, alumina, and rare earth.

Regardless of the method used to produce FCC catalysts, the methods involved in statistical process control (SPC) are increasingly being used to regulate and define the properties of finished catalysts. The preparation of FCC catalyst depends on the accurate and homogeneous mixing of a variety of components. Although each of these components serves a specific function in the catalyst particle, they often interact synergistically. Using SPC ensures that the mixture and ratios of these various components are optimal and that tight statistical control is maintained in all phases of the manufacturing scheme.

This chapter describes the basic components of all present day FCC catalysts. Methods of synthesizing and analyzing the resulting catalyst and its components are also described as well as several materials not currently used that may be commercialized in the near future. It is the purpose of the present chapter to review the chemical changes that have occurred in FCC catalyst preparation since the introduction of zeolites into FCC catalysts in 1962 and to describe the basic chemical methods used to prepare effective catalysts. The catalytic functions of various active ingredients are additionally described in detail in Chapters 2, 3, and 5.

11. THE ZEOLITE COMPONENT OF THE FCC CATALYST

Without doubt, the most important component of the modem FCC catalyst is the zeolite. Although zeolites have been known and characterized since the 19th century, their commercial use did not become practical until the pioneering synthetic work by Barrer and Union Carbide researchers in the 1940’s and 1950’s (2,3). The first commercial use of zeolites was as adsorbents utilizing the small pore type A zeolite first synthesized by Union Carbide (3). Carbide developed and commercialized the synthesis of zeolite X and zeolite Y (4). The use of these zeolites in FCC catalysts only became feasible after Plank and Rosinski discovered that the zeolite had to be dispersed in a relatively inactive matrix to prevent overcracking and deactivation (5).

Zeolite Y is sti l l the major active ingredient in FCC catalysts, but it has undergone many modifications since its introduction to meet the changing needs of

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the refining industry. This section describes the commercial synthesis of zeolite Y as practiced by the major producers of FCC catalysts and the subsequent modifications of the zeolite that are used to produce high-quality FCC catalysts for today's market.

Milton, who originally patented the synthesis of zeolite Y, used precipitated silica, sodium aluminate, and caustic soda as the starting material (4). An induction or aging step of 24 hours was required at lower temperatures before actual crystallization began. Later work at the Davison Division of W. R. Grace and Co. extended the synthesis to include the use of sodium silicate and seeds to eliminate the aging step (6,7). Hayden et al. at Engelhard developed methods that allowed the in situ synthesis of zeolite Y from thermally treated kaolin and work at Filtrol developed the synthesis from acid-leached metakaolin (8-10).

A. Commercial NaY Svnthesis

Three major synthetic techniques are used in the commercial preparation of zeolite Y for FCC. These techniques are sodium silicate based, precipitated silica based, and clay based.

1. Synthesis Using Sodium Silicate

The synthesis of NaY from sodium silicate involves the use of aluminum sulfate, sodium aluminate, and initiator (seeds). Initiator is prepared by slowly mixing sodium aluminate solution into a solution of caustic and sodium silicate under conditions of high agitation and controlled temperature (7). The resulting clear solution must be aged for 12 to 16 hours at temperatures below 120°F prior to use. The aging must be done in the absence of any agitation. The synthesis of the zeolite Y using the initiator and sodium silicate proceeds as outlined in Figure 1. Because zeolite Y is metastable, the crystallization temperature must be kept between 200" and 218°F to prevent the formation of zeolite P. Also, after complete crystallization, the slurry must be quickly quenched to temperatures below 150°F to prevent degradation of the product to zeolite P. The crystallization is generally done in the absence of agitation, because high shear results in low crystallinity.

Zeolite crystallization is autocatalytic. Crystal formation usually follows an S-shaped curve (Figure 2). The rate of crystallization and the silica-alumina ratio of the final product are highly dependent on the silica-soda-alumina ratio of the initial gel. The relative ratio of these components and water must be within defined limits or phase field for complete crystallization to occur. Stoichiometries outside of the phase field result in impurities such as zeolite P or the formation of amorphous solids.

After crystallization, the zeolite is usually separated from the mother liquor by filtration. The mother liquor, or filtrate, contains sodium silicate, sodium sulfate, and a small amount of sodium aluminate. The silicate is in the form of a disilicate (2 SiO,.Na,O). The mother liquor is often recycled and used as a component in the batch makeup of the zeolite synthesis (1 1). Hence, the disilicate solution is neutralized with aluminum sulfate to form a silica-alumina gel, which is fitered and

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Figure 1 : Seeded NaY Preparation Schematic

aSodiu;;;minate I , J. , Age 12-1 6 HE.

I Sodium Silicate

Sodium Hydroxide Water

Alum-Slow Addition Rate

Sodium Silicate Mother Liquor to Recycle b Product Cake Recycled M.L.

Figure 2: Zeolite Y Growth Curve

7/1.8/1 .O SiO,/Na,O/Al,O, Mole Ratio 100

90

80

70 Crystallinity 60

Wt%

40

30

20

10 0

Time (HE.)

109

washed. The silica-alumina gel is then used in the synthesis to replace some of the silica and alumina obtained from other raw materials. Recycle not only avoids disposal of waste streams but also reduces raw material costs.

2. Synthesis Using Precipitated Silica

The synthesis of zeolite Y using fumed or precipitated silica as the raw material source proceeds in a manner similar to systems based on sodium silicate but with some important differences. Because sodium silicate is not used in the process, the use of silicate-based seeds and mother liquor recycle is not an option (5). Because seeds generally are not used, the initial gel is low-temperature (100°F) aged prior to crystallization. Also, the use of silica instead of sodium silicate removes the necessity of using aluminum sulfate. Sulfate is not present during crystallization as is the case for the sodium silicate based synthesis.

3. Synthesis Using Clay

Two commercial methods of synthesizing zeolite Y have been developed in which kaolin clay is the starting material. One case involves total conversion of the clay into zeolite, whereas the other requires only partial conversion.

a. Leached Metakaolin

The synthesis of zeolite Y can also be accomplished by starting with the acid treatment of metakaolin to remove alumina (10). When sufficient alumina is removed, the kaolin is then mixed with caustic, aged, and crystallized to produce pure zeolite Y crystals. The removal of aluminum is required to bring the overall soda-silica-alumina ratio to the correct ratio for zeolite Y growth. Calcination of the kaolin to metakaolin is required to activate the alumina present in the clay. The clay calcination temperature is critical as is the quality and consistency of the clay source. This method is not presently practiced commercially.

b. Clay Calcined at High Temperature

Use of high-temperature-calcined kaolin in zeolite synthesis was described in several patents by Hayden, et al. (8,9). In this technique the complete conversion of the material to zeolite Y is neither obtained nor desired. Rather, the clay is calcined above the high-temperature kaolin exotherm to an incipient mullite phase. Only part of the alumina present in the clay is active for zeolite synthesis.

In the original patent, the synthesis was done on extrudates containing both high-temperature-calcined-kaolin and metakaolin (8). Crystallization was generally carried out in the presence of caustic in a mineral oil medium. This procedure produced catalyst for moving bed reactors.

Later patents extended the technology to spray-dried microspheres and aqueous slurries. In the latter, the kaolin is spray dried to form microspheres suitable for FCC. The microspheres are then calcined and treated with caustic solution to form approximately 25 to 30 wt-% zeolite Y dispersed throughout the

110

microsphere. Aging of caustic-microsphere slurry at temperatures around 100°F for 6 to 12 hours is required before crystallization at 180°F. Filtration of the slurry again results in formation of a disilicate mother liquor. Brown shows how to retain part of the mother liquor in the microsphere by removing the washing step after filtration (12). The material is subsequently flash dried before further treatment to limit the matrix surface area and decrease gas and coke make.

More recent patents using similar technology require both metakaolin and high-temperature-calcined kaolin along with seeds to obtain crystalhities near 60% in the microsphere (13). Crystallization time and variation of the clay-type ratio presumably can be controlled to vary the degree of crystal formation. In all cases, the part of the microsphere not converted to zeolite Y provides a stable matrix, which binds the zeolite and contributes toward the cracking of molecules too large to penetrate the zeolite supercage.

4. Comparison of Zeolite Y Synthetic Techniques

Of the three methods used in the synthesis of zeolite Y for FCC, the Fist two are similar enough so that use depends mainly on the local economics of the producer, such as raw material cost and waste disposal cost. The in situ technique developed by Hayden results in significantly different properties (9). On one hand, this method of synthesis appears to present some advantages in raw material cost and improved zeolite dispersion. Furthermore, it is self-binding and hence avoids the problem associated with various binders as outlined in the section on binders below. On the other hand, in situ synthesis is less flexible than methods using ex situ zeolite synthesis. Formulation changes involving additives, different zeolite types, and matrix variations are considerably easier to accomplish in systems not constrained by in situ zeolitization. Additives could retard or prevent zeolite crystallization, and different zeolites, for example mordenite and ZSM-5, are difficult to include in the catalyst composition.

B. Svnt hesis of Aluminum Deficient Zeolite Y

Until recently, the direct synthesis of zeolite Y was restricted to product silica-alumina ratio of less than 6.0. Because of the increased chance for low crystallinity or impurity formation when operating under conditions designed to yield silica-alumina values higher than 5.6, commercial production usually is limited to silica-alumina ratios below 5.6. However, in the early use of zeolite Y in FCC catalysis the hydrothermal stability of zeolite Y was discovered to be highly dependent on zeolite silica-alumina ratio. This conclusion led to methods of secondary synthesis designed to increase the silica-alumina ratio of as synthesized sodium Y. The discovery by Maher and McDaniel of ultrastable Y (USY), the first of the aluminum deficient Y zeolites, led to its inclusion in FCC catalysts as early as 1964 in XZ-15 produced by Davison (14,15). However, the use of USY aluminum-deficient zeolite in FCC catalyst quickly disappeared because of its substantially lower activity compared to fully or partially rare earth exchanged Y zeolite (CREY and REHY). The increased olefinicity and gasoline octane obtained by using USY did not become important until regulations mandating the decreased use of lead in gasoline came on the scene in the mid-1970’s.

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The Octacat series of catalysts containing USY in concentrations greater than 25 wt-% appeared on the market in 1976, but the relaxation of the regulations restricting lead in gasoline temporarily decreased interest in zeolite octane catalysts. With the mandated elimination of leaded gasoline in the 1980’s, octane-enhancing FCC catalysts again became extremely important to the oil industry. Catalysts containing USY again appeared on the market along with new technolog for producing dealuminated (high silica-alumina ratio) Y zeolite.

Techniques have been developed for the chemical dealumination of the zeolite. Such dealumination would minimize the presence of nonframework alumina (NFA) generated by the hydrothermal techniques used to produce USY. Union Carbide patented and commercialized dealumination of zeolite by ammonium fluorosilicate (AFS) for use in FCC catalysts (16). Commercial techniques involving the hydrothermal treatment of the zeolite to form a conventional USY containing NFA formed during dealurnination were now modified by chemical washing to eliminate the NFA. Because of the superior stability, octane-producing characteristics, and coke selectivity of these various aluminum deficient zeolites, they quickly controlled the major portion of the United States and European catalyst markets.

1. Hydrothermal Dealurnination

Maher and McDaniel prepared two types of USY by combining ammonium exchange and calcination to reduce sodium content and unit cell sue (UCS) of the zeolite (14,17). USY-A relied on calcination of a partially exchanged ammonium Y zeolite to produce zeolite Y with UCS of 24.55 A. Additional exchange and final calcination temperatures above 1300°F resulted in USY-B, which exhibited a UCS of less than 24.35 A and sodium levels of less than 0.1 wt-%. The technique used wet zeolite and deep bed calcination so that the material was self-steamed. The presence of steam promoted hydrolytic attack on the framework alumina resulting in zeolite dealumination. Care as to the amount of sodium left on the zeolite and the calcination temperature was required to prevent catastrophic alumina loss and zeolite crystal destruction.

In an alternate technique described by Kerr and others, the desired dealumination occurred with shallow bed calcination of ammonium Y in flowing steam (18-20). The flowing steam allowed better control of the steam partial pressure and rate of dealumination, and reduced calcination temperature.

Current commercial methods of producing USY by hydrothermal techniques involve modification of the two previously described procedures. The starting sodium zeolite is ammonium exchanged to approximately 3 wt-% sodium on zeolite (21,22). The zeolite is then calcined to yield product with a UCS of approximately 24.56 A and a crystallinity of 80 to 90 wt-%. The calcination temperature is dependent on the degree of dealumination desired and the steam partial pressure. Steam can either be injected into the calciner or evolved from zeolite. The drying temperature prior to calcination can be varied to increase the water content of the ammonium Y zeolite. After calcination, zeolite can be subjected to further ammonium exchange to reduce sodium content to less than 0.5 wt-% or the calcined zeolite can be included in the FCC catalyst and final exchange done in situ.

112

2. In Situ Dealurnination

Another commercially practiced method of producing USY is the so-called in situ technique. In this case, sodium Y is included in a FCC microsphere, and the calcination and exchange steps are done on the catalyst as opposed to the pure zeolite (23). This method is required when the in situ method of zeolite synthesis developed by Hayden is used (9). It also can be practiced by those using other manufacturing techniques. One advantage of this method is the spray dried microspheres are easier to exchange, filter and handle than are the zeolites with relatively small particle sue. Also conversion to USY occurs usually in the presence of an alumina containing matrix, which stabilizes the zeolite and increases activity when the alumina migrates into zeolite sites (24). Matrix alumina may also serve as a soda sink.

Hydrothermal dealumination of zeolite Y to produce USY causes aluminum ions to be removed from the zeolite framework. This removal necessarily leaves holes or defects in the crystal and causes some loss of crystallinity. The presence of hydroxyl nests in the area of the removed alumina has been observed by a number of investigators (25-27). Presumably,the hydroxyl nests are destroyed at elevated temperatures, and silica from the noncrystalline phase that is present after USY formation can migrate to these sites and be inserted into the structure (28). Some, but not all, of the defects are thought to be removed by this migration. What is left is residual amorphous alumina or silica-alumina that is present in or around the zeolite. This alumina is generally octahedral in nature and is referred to as nonframework alumina.

Researchers using nuclear magnetic resonance (NMR) techniques have also observed the presence of nonframework alumina, which is known to be present by bulk analysis and acidity studies, but is not accounted for by NMR (29). The coordination state of this alumina is not known. Other workers have also suggested that some of the alumina removed from the framework exchanges into available cation sites on the zeolite (30). Such exchange is said to improve the overall stability and activity of the USY. Again, such exchanged alumina are octahedrally coordinated.

3. Chemical Dealumination

Chemical dealurnination, as the name implies, involves the removal of alumina from zeolite by the action of various chemicals. These reactions are carried out at temperatures below the boiling point of water.

a. Dealumination by Chelation

In addition to hydrothermal dealurnination, the chemical dealumination of zeolites is also an important method of producing aluminum-deficient Y zeolite. It has been known since the early 1960’s, chelates, such as acetylacetone and EDTA, could be used to extract aluminum from the zeolite framework (31,32). The chelates generally formed a soluble aluminum complex in the solvent, and, as a result the extraction process required great care because rapid removal of aluminum can result

113

in crystal destruction. Also, removal of more than 50% of the framework aluminum usually resulted in significant loss of zeolite crystallinity (31). Chelating techniques were never used commercially, probably because of the difficulty of control and the large amount of organic waste streams formed in the process.

b. Silicon Halogenate Dealwnination

Another method of chemical dealumination involved the reaction of silicon halogenates with partially exchanged NH,Y (33,34). Specifically, silicon tetrachloride was used to remove aluminum from the zeolite framework and insert silica into the framework vacancy. Because silicon tetrachloride reacts rapidly with water vapor, the reaction must be done in a nonaqueous solvent or by using vapor-phase silicon tetrachloride. Again, the rate of aluminum removal is important to prevent crystal degradation. This technique, which involves silica reinsertion has the advantage of producing a defect free surface after the dealumination process. The overall problem with commercially implementing this technique is the corrosive nature of silicon tetrachloride and the anhydrous conditions required for its use.

c. Ammonium Fluorosilicate Dealwnination

Breck and Skeels described a combination dealumination and silicon reinsertion process using ammonium fluorosilicate (AFS) and partially exchanged ammonium Y (16). The overall reaction scheme for this process is given in Figure 3 (35). The AFS reacts with framework aluminum under carefully controlled conditions of temperature and addition rate. The aluminum forms ammonium hexafluoro-aluminate (AFA), which can slowly hydrolyze to form aluminum hydroxy- fluorides and HF. AFA reacts with alum [A12(S04),] to form ALFZ+ by the following reaction:

(NH4)@6 + AlZ(SO4), -> 3AW' +3SOP + 3NH4+ + 3F

The zeolite is separated from the alumina and fluoride by filtration. The resulting dealuminated Y was shown to retain all of its initial crystallinity and to be defect free. Because the reaction is stoichiometric, the amount of dealumination can be controlled by the amount of AFS added. Theoretically, the technique can be used to produce zeolite at high silica-alumina ratios. However, the greater the degree of dealumination desired, the slower the addition rate of AFS and the mere difficult the process is to control to prevent crystal degradation. Also careful control of the addition rate is required to prevent silica-alumina concentration gradients from the surface to the zeolite interior. When the reaction is properly done, little or no gradation occurs (37).

Subsequent to the Breck-Skeel's patent, several patents have been issued that substitute heteroatoms such as cobalt and iron in the zeolite Y framework by using ammoniumfluoro salts of the desired metal (37,38). This method of producing zeolites with various transition metals or other heteroatoms in the framework is called secondary synthesis. The technique is applicable to a variety of hetero ions and a variety of zeolites, for example cobalt, iron, chromium, and zeolite Y and various pentads.

114

Figure 3: Mechanism for Chemical Stabilization Using AFS*

1 1) [AIO;], NH,'+ H,O <=> [AIO;], H+ + NH,OH

2)

3)

[AIO;], H' + 3 H,O <=> [(OH) J + AI(OH),

NH,OH + H+ => NH,+ + H,O

4) (NHJ,SiF, + 4 H,O <=> Si(OH), + 6 F- + 4 H+ + 2 NH,'

5)

6)

7)

8)

AI(OH), + 6 F <=> AIF," + 3 OH

3 OH- + 3 H+ => 3 H,O

[(OH) J + Si(OH),=> [SiO,] + 4 H,O

[AIO;], NH,' + (NHJ,SiF, => [SiOJ + (NHJPIF,,

Where: [ 1 Indicates an Element of the Lattice Framework

<=> Indicates a Reversible Chemical Equilibrium

=> Indicates an Irreversible Reaction 'After Hinchey; 1987

4. Combined Hydrothermal and Chemical Dealumination

Combinations of hydrothermal and chemical treatment in the manufacture of zeolite Y are also practiced commercially (39-41). These techniques usually involve initial zeolite dealumination by standard hydrothermal methods described previously. These standard methods can involve either in situ or ex situ dealumination.

After ion exchange, the zeolite (or catalyst) is subjected to one or more low pH exchanges or acid washes to remove NFA. Use of organic chelating agents to effect NFA removal has also been described (42). In the case of acid wash, the solution consists of ammonium nitrate or sulfate acidified with nitric or sulfuric acid to 2.5 to 3.0 pH. The post calcination acid wash removes the majority of the NFA but causes some further dealumination of the zeolite. The resultant product is highly crystalline, and the concentration of NFA is low. However, structural defects are not removed and defect-free crystals generated from silica insertion techniques are not obtained.

5. Direct Synthesis of Aluminum Deficient Y

Recently, workers from Exxon have described direct synthesis of zeolite Y with silica-alumina ratios as high as 12 (43,44). In this method of synthesis, tetraakylamine hydroxides are used as templating agents, and the crystallization occurs under autoclave conditions at elevated temperature and pressure. The product zeolite is highly crystalline and defect free. Also, according to the authors,

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this technique results in random distribution of the alumina throughout the crystal lattice (45). As mentioned above, alumina gradients can occur when using silica insertion techniques improperly. The expense of the templating agents and equipment for this process make its widespread commercial use unlikely. The method does serve as a useful laboratory technique for comparison with other methods of producing dealuminated Y zeolite.

6. Comparison of Dealurnination Techniques

Comparison of the value of various dealurnination techniques is highly dependent on the care during synthesis. The quality of the product, nc matter what the technique, is critically affected by careful control of reactant addition rates, temperature, and so forth. Hence, poor or improperly performed synthetic techniques can easily produce inferior product from an otherwise entirely acceptable process. Intercomparison of preparative techniques described in the literature requires experience and a statistically significant number of repeat runs before it can be assured that the synthetic technique is proper.

The silica-insertion techniques described by Breck and Skeels represent a seemingly simple inorganic metathetical reaction, which should, in its simplicity, produce an essentially defect-free and highly crystalline zeolite compared with the other commercial methods (16). Slow dealumination followed by rapid insertion of silica into the vacancy created in the zeolite framework prevents local crystal collapse or defect formation. By their very nature, hydrothermal techniques would not be expected to exhibit this degree of control. Also, because the Breck-Skeels technique involves rendering the alumina removed from the framework soluble, separation of the resulting NFA from the zeolite is easily accomplished by filtration. The product zeolite would thus not contain NFA if rigorous preparative techniques were followed.

The advantage or disadvantage to catalyst activity and selectivity presented by NFA is beyond the scope of this chapter but is described in Chapter 2. The consensus is that some removal of NFA is clearly desirable. All FCC catalyst manufacturers now produce some grades of catalyst in which the amount of NFA has been minimized. As mentioned in the section of "Ammonium Fluorosilicate Dealurnination," zeolite dealumination using AFS is a reasonably straightforward stoichiometric reaction. The zeolite silica-alumina can be controlled exactly by a variation of the AFS/NH,Y ratio. Although techniques have been developed using time, temperature, and steam partial pressure to achieve various degrees of dealumination by hydrothermal methods, they are by nature more difficult to control. Furthermore, a wide range of silica-alumina ratio Y zeolite can be prepared using AFS without losing crystallinity or generating crystal defects if AFS addition rate is carefully controlled (16).

Wang, et al., have suggested that the use of AFS can produce zeolite with large gradients in silica-alumina ratio between the surface and interior of the zeolite particle (46). Such gradients are not found when hydrothermal methods of dealumination are used. The gradient results in silica enriched surfaces, which can affect zeolite activity toward cracking of heavy molecules. Presumably, the

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alumina-depleted surface would possess fewer active sites on which to crack larger molecules. These larger molecules may not readily diffuse into the zeolite particle. Whether this gradient would remain after steaming of the catalyst is unclear. The existence of such gradients is highly dependent on the AFS addition rate (47). With careful addition, such gradations can be eliminated.

111. THE FCC CATALYST MATRIX

Early attempts to use pure zeolite in the catalytic cracking of gas oil met with failure. Because of the high surface area, site density of the zeolite, and generally long catalyst-oil contact times involved in bed cracking in early FCC units, the product stream was high in undesirable secondary reaction products such as coke and light gas. Clearly, overcracking occurred, and coke rapidly deactivated the pure zeolite catalyst. Plank and Rosinski demonstrated that a simple dilution of the zeolite in a matrix reduced overcracking and coke poisoning and produced a zeolite containing FCC catalyst that was far superior to existing amorphous silica-alumina or acidified clay based catalysts. Amorphous FCC catalysts were the dominant materials in use for 20 years prior to the invention of zeolite containing FCC catalysts (see also Chapter 1.)

Early zeolite FCC catalysts contained rare earth exchanged zeolite (type X or Y), kaolin clay, and a binder. The binder was synthetic silica-alumina or alumina, usually peptized pseudoboehmite as was used in many prezeolite FCC catalyst systems (48-50). The in situ technology of Hayden was binderless in the sense that particle hardness was accomplished by the thermal treatment of the spray dried clay microsphere prior to zeolitization (9). No specific component was added to the spray dryer feed to actually bind the clay particles together. All three of these binders were active themselves, although with different selectivities in the catalytic cracking of oil molecules. The reaction over these matrices was considerably less selective than over zeolite (more coke and gas and less gasoline). Because in the 1970's gasoline selectivity was paramount and FCC feed was rather paraffinic, the cracking activity of the matrix was considered a disadvantage. Changes in unit design to take increasing advantage of zeolitic cracking and the steadily increasing catalyst zeolite content, increased demand upon catalyst attrition resistance. This increased demand led to the introduction in the mid-1970's of binders based on silica sol. The binder was low in surface area and essentially inactive in the cracking of gas oil. It produced a significantly harder catalyst than did the previous ex situ technology.

Increasing use of heavier feeds and the need for increased metals tolerance created the necessity for reintroducing active matrix components into the FCC catalyst. This need has steadily increased over the last several years. New FCC catalysts are required to be extremely hard but also sufficiently porous to crack heavy, large oil molecules into the fuel oil range or to molecular sizes capable of being cracked in the small pores of the zeolite component. Moreover, matrices must also be capable of reducing the effect of contaminant metals (usually vanadium and nickel) on gas and coke make and serving as a soda sink. They should also interact synergistically with the zeolite to maximize yields of transportation fuel and

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me silica Sol

Aluminum Chlorhydrol (ACH)

limit the formation of coke and light gas. Each of these matrix components will now be described in detail in the following sections.

Surface Area Activity

20 m2/g Very low

60-80 m2/g Moderate

A. Binders

Peptized Alumina

Self-Binding (In situ)

The oil industry has gradually increased demands for attrition resistant FCC catalyst. These demands have arisen from the need to reduce particulate air pollution from catalyst fines and to prevent the loss of expensive catalyst. Hence, the binder has become an increasingly critical component of the catalyst matrix. Commercial FCC catalyst producers have devoted substantial time and effort to improving catalyst binding. At present, four binder systems are used in the commercial production of FCC catalyst, and an extraordinary number have been examined in laboratory preparations. Table 1 shows the four binder types and their main characteristics.

300 m2/g High

- High

Table 1 Commercial FCC Binders

1. Silica Sol

The acidified silica sol binder was introduced commercially in the mid-1970's. The binder was attractive for its low activity, ease of manufacture, and dispersibility. The sol is produced by an intimate mixture of sodium silicate and acidified aluminum sulfate solution (51). A clear, low-viscosity silica sol results. Because silicate gels rapidly in the pH range from 11.0 to 3.5, the silicate and acid alum must be mixed rapidly to prevent local pH excursions and micro gelation of the sol. Thus extremely effective local mixing is required.

Furthermore, gel time is also strongly affected by reactant concentrations and solution temperature. Consequently, the sols are made at relatively low silica solids (10 wt-%) and cooling of the components is required. Mediocre mixing, elevated temperatures (> 100"F), or high silicate concentrations cause the formation of turbid sol or gel particles with poor binding characteristics. Once formed, the sol is in a metastable state (52). Particle growth is dependent on temperature, pH, and silica concentration. Even sol formed under conditions known to impart good binding can form microgel within one hour of synthesis and form a clear silica gel within several hours at a pH equal to 2.5 to 2.7. Although producing the sol from the interaction

i i a

of silicate with sulfuric acid is possible, the sol product is unstable with respect to changes in pH and rapidly gels when mixed with other FCC catalyst components. The presence of aluminum (111) offers some buffering capacity to the sol and prevents rapid gelation of the spray dryer feed slurry. When the sol is mixed with the other components, it rapidly disperses throughout the slurry. #en spray dried, the sol rapidly gels, binding the components together.

2. Aluminum Chlorhydrol

The aluminum chlorhydrol (ACH) binder was introduced commercially in the late 1970’s (53). The binder is a hydroxylated aluminum based sol containing chloride as the counter ion. Aluminum chlorhydrol can be produced by reacting aluminum metal with hydrochloric acid to get a clear solution at pH 3.0 to 4.0 (54). The aluminum-chloride ratio can be varied and determines the amount of various molecular weight alumina oligomers in the sol. Likewise, the material can be produced by the slow hydrolysis of AlCI, solution with base. Addition rate of the base to the AlCI, solution is critical as is mixing to prevent the formation of Al(OH), gel.

Although the equilibria of alumina species in the sol are complex, it is better understood than that of silica sol in the same pH range. The predominant species seems to be [Al,304(OH)24(H20),2]7+. This species has been characterized and exhibits a Keggin-type structure as shown in Figure 4 (55). The central aluminum ion is tetrahedrally surrounded by four groups of octahedrally coordinated aluminum hydroxy species. Each group contains three aluminum atoms. Compared to silica sol with average particle sizes of 5 to 10 nm, the size of the aluminum oligomers is relatively small. Thus dispersion throughout the FCC matrix would thus be expected to be superior.

Unlike silica sol, the solution structure of ACH is not particularly sensitive to time or temperature but is sensitive to pH. Because of its high chloride content (12%), ACH is also considerably more corrosive, especially at elevated temperature, to stainless steel catalyst preparation equipment than is silica sol. The ACH remains stable indefinitely at constant pH and can be refluxed without gelling and without any other noticeable change in sol properties.

In practice, ACH is mixed with other FCC components, and the resulting slurry is spray dried. The spray dried microspheres, unlike the spheres bound by silica sol, show little green strength. They must be directly calcined at temperatures high enough to cause dehydroxylation of the alumina in order to be slurried without damage to the microsphere or to gain attrition resistance. The spray dried microspheres cannot be slurried in water and maintain physical integrity prior to firing, as is true with acidified silica sol bound microspheres. Because the calcination of ACH-bound catalysts produces dehydroxylated alumina species, the binder does possess some surface area and cracking activity. Whether the active sites are Lewis or Bronsted acids is not clear.

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Figure 4:

The drawing shows how the 12 AIO, octahedra are joined together by common edges. The tetrahedron of oxygen atoms in the center of the structure contains one four-coordinate aluminum atom. (Johansson, 1960)

3. Peptized Alumina

Another commonly used alumina based binder is prepared by peptizing pseudoboehmite alumina (PBA), usually with formic acid (56). The action of the acid on the reactive alumina produces an alumina sol with particle size in the range of several hundred nanometers. The opaque, milky-white color of the sol testifies to its significantly larger particle size compared with silica sol or ACH. Other monobasic acids such as HCI and HNO, can also be used to peptize the alumina, but formic seems to be the commercial choice, because of fewer problems in terms of stack emissions or corrosion. Pseudoboehmite alumina cannot be peptized with di- or tri-basic acids, such as sulfuric acid. After acidification, the peptized alumina is mixed with the other FCC components and spray dried. Like ACH, the spray dried microspheres have little green strength and cannot be slurried in water without significant damage. The sol itself is somewhat unstable with respect to pH, time, and temperature. Depending on alumina concentration and pH, the sol will gel within ten minutes to several hours. Unlike silica sol, the gel particles seem to be only weakly bound, and whether gelation affects the binding properties of the sol is unclear.

The spray dryer feed slurry prepared with peptized alumina is noticeably more viscous than corresponding slurries using silica sol or ACH as binders. In practice a small amount (3 to 5%) of stabilized polysilicic acid (PSA) is added to the spray dryer feed to improve overall binding (57). Stabilized PSA can be prepared by passing dilute sodium silicate solution through an acid ion exchange column to remove most of the sodium ions. Subsequent stabilization is affected with minimal

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amounts of base , usually ammonia (58). Pseudoboehmite possesses surface areas of more than 300 m2/g and undergoes a phase change to gamma alumina at FCC unit regenerator temperatures (59). The acidity of gamma alumina is well documented (60,61). Hence, this binder is active for gas oil cracking and serves a dual purpose as an active bottoms cracking component and binder. Bottoms are the large high-molecular weight components in the FCC feed. Usually their size makes them too large for diffusion into the zeolite pores (diameter=7.4 A)

4. In Situ Binding

The in situ type of FCC catalyst preparation introduced by Engelhard contains no added binder (9). The first preparative step calls for forming kaolin microspheres. However, clay-to-clay particle binding after spray drying is quite weak, and the particles cannot be slurried in water without complete breakdown. After high-temperature calcination (approximately 1800 O F ) the particles develop significant attrition resistance and maintain their microspherical shape during the second step of the in situ crystallization process in caustic solution at 180°F.

The crystallization step imparts even more hardness to the microsphere and forms intrastructural Nay. Hydroxyls on the surface of the clay platelets react with one another at high temperature to eliminate water and form oxygen-bridged bonds between the particles.

During crystallization, significant silica is removed from the particle to leave macropores which are filled by the intrastructural zeolite particle growth. Significant porosity in the 30 to 100 A range results. The remaining matrix is higher in alumina than the starting kaolin because of the silica removal and possesses significant acidity (62). The combination of acidity and high surface area make this material an active matrix for bottoms cracking. The large matrix surface area and exposure of contaminant metal present in the clay can lead to high coke and gas make. Retaining some of the silica (as described by Brown) reduces some of this undesirable matrix activity by reducing overall surface area (12). Recent technology using combinations of metakaolin and seed technology seems to reduce the porosity and surface area to an even greater extent (13,63).

5. Influence of Particle Size Reduction on Binding

The average diameter of an FCC catalyst particle is 70 p. Particles greater than 8 p in diameter can not be effectively bound into a 70 p particle. Hence, particle sue of the components making up the microsphere plays a significant role in overall catalyst attrition. The surface chemistry of these components is also significant in determining overall catalyst attrition.

With in situ synthesis, the clay particles in the spray dryer feed are already quite small (50% point = 2 p; 90% point < 1 p ) and require no further reduction in particle size. In the mixed component methods of FCC catalyst synthesis previously described, the zeolite and bottoms components frequently have particle sizes over 5 microns and 90% points above 10 microns. Some form of particle size reduction is required. Power milling of these particles to the 2 micron APS range is practiced by

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all the commercial FCC producers not using in situ crystallization. The type and degree of milling varies considerably depending on the amount and initial size of the matrix components being processed and the type of binder.

Dry milling techniques are used in some instances. Wet milling techniques, such as sand or media milling, are also quite effective. Milling not only changes particle size but can also increase slurry viscosity and temperature. These negative effects often need to be compensated for by using surfactants and slurry cooling. Figure 5 shows typical effects of particle size on relative attrition using constant catalyst formulation. Although the results are for a silica sol based binder, the effect is qualitatively the same in all binder systems.

Figure 5: Relative FCC Attrition vs D90 Particle Size 100 W

90

80

70

-

-

- Relative Attritinn 60 % a - ... I..."..

30

4 5 6 7 D90 Particle Size (Microns)*

90 Wt% of Sluny Particles Less Than

B. THE ACTIVE PORTION OF THE MATFUX

As described in section 111 A, modern FCC catalysts require active stable matrices capable of converting heavy oil molecules to lighter products and minimizing the effects of contaminant metals. Three such matrix components have been used in commercial FCC catalysts, whereas others have shown large promise in laboratory testing.

1. Alumina

By far the most common material added to commercial FCC catalysts to accomplish heavy oil conversion is alumina in various forms and amounts. Alumina has long been used as an active support in reforming and hydrotreating catalysts

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(64-66). In FCC catalysts, some silica is included with the alumina either in the form of added silica (polysilicic acid), silica binder or is present naturally in the clay based in situ process. Workers in the alumina area have shown that under hydrothermal conditions, silica stabilizes the alumina surface area (67,68). Undoubtedly, silica also forms tetrahedrally coordinated silica-alumina species on the alumina surface to produce Bronsted acidity and a more active cracking surface. Other known technologies such as surface area stabilization by lanthanum impregnation are also possible (69).

Commercial manufacturers use PBA, amorphous alumina, and aluminum chlorhydrol sols to add alumina to the FCC catalysts. The amount and type of alumina added to the catalyst formulation depends on the FCC unit severity and the feedstock being processed. In the case of PBA, the material is generally peptized prior to addition. Because of the small particle size achieved by peptizing the PBA, no additional mechanical particle size reduction is required. Amorphous aluminas are necessarily subjected to some form of mechanical particle reduction prior to inclusion in the catalyst matrix. Usually the particle size must be reduced to an APS near 2 microns to produce the extremely attrition resistant catalysts required to meet particulate emission standards of today. The particle size reduction is achieved by mechanical milling techniques, and the alumina is added to the spray dryer feed as an aqueous slurry. Some care must be taken to prevent reagglomeration of the particles in slurry by adjusting the pH below the isoelectric point of the alumina.

In the case of in situ technology, free, active alumina is formed by the nature of the crystallization procedure itself. Further alumina addition is not necessary. In all of these cases, most of the added alumina converts to a high surface area gamma or chi alumina on steaming or in commercial use.

The surface area, pore structure, and hydrothermal stability of the alumina are critical in FCC operations. In general, the function of the active alumina is to crack large, heavy molecules boiling typically above 650°F into a material with a lower molecular weight. These heavy molecules are converted to molecules small enough to be further recracked by the zeolite into the gasoline or LPG range.

A second surface function served by alumina is to absorb condensed aromatics, which are not crackable. The absorption prevents the molecules from blocking access to the zeolite and decreasing catalyst activity.

A third function of the alumina is to react with the vanadium metal contained in porphyrin structures, which are present to a greater or lesser degree in all feedstocks, and prevent its migration to the zeolite. Migration to the zeolite will result in crystal destruction by vanadic acid formed from V,O, and H,O in the regenerator. Unfortunately, alumina can also serve as a good support for both nickel and vanadium and increase their coke and gas making tendencies. Alumina by itself is not a particularly good metals passivator. Finally, the alumina also imparts stability to the catalyst by mitigating the destructive effects of hydrothermal dealumination on the crystal structure of the zeolite. Hence, overall catalyst stability and activity is improved.

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2. Silica-Alumina

Amorphous silica-alumina, in which aluminum is tetrahedrally coordinated to silicon through oxygen bridges, was the active cracking component in many FCC catalysts before the discovery of zeolites as an active FCC catalyst component (70,71). Even though the active sites thus formed were similar to those in zeolite X and Y, the amorphous silica-aluminas were much less gasoline and coke selective than the zeolite containing catalysts and were much less active. In the early years after the introduction of zeolite cracking catalysts, amorphous silica-alumina matrices were used in conjunction with the zeolite. However, because of their relatively non-selective cracking characteristics and the fact that they were poor binders, their use was abandoned in favor of better binders. However, with the pressing need for active bottoms cracking components, amorphous silica-aluminas are again being examined (72).

3. Pillared Interlayered Clays

Both Vaughan and Shabtai described the use of pillared interlayer clays (PILCS) in FCC (73,74). These materials are prepared by intercalating large, inorganic based cations such as ACH into the interlaminar region of swellable clays, mainly smectites. After calcination of the intercalated clay, the ACH is dehyroxylated to form alumina based pillars which then prevent the clay layers from collapsing to their original interlayer spacing. Materials with pores of approximately 9 to 10 A are formed using ACH. The internal pore size can theoretically be varied by the size and amount of the intercalating agent employed. The pores are well defined and zeolitic in nature, and pore diameters significantly larger than the approximately 7.4 A pore mouth found for faujasite are possible. Such materials appear to have an application in the cracking of feeds containing a large portion of heavy cycle oil (HCO) or to increase light cycle oil (LCO) production from the selective cracking of HCO into the LCO range.

Since the original Vaughan, et al., patents a large amount of work has been done on PILCS (75-80). However, most of the materials described lack the necessary hydrothermal stability required of FCC components. Recently, a patent by McCauley described a Cerium-ACH pillared clay that seems to possess the required hydrothermal stability (81). These materials also possess pores in the 17 to 18 A range. Table 2 compares the cracking of gas oil over two examples of bentonite exchanged with cerium-ACH with a commercial FCC catalyst at constant reaction conditions. The LCO selectivity is excellent compared with the zeolite-based control, but the amount of coke make over the PILC is extreme.

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Table 2 Product Yields in the Conversion

of Gas Oil over PILC

* After Steaming at 1500°F for 5 hours

Others have also described hydrothermally stable PILC’s (82). The increased activity in developing these materials for FCC use and their potential advantages suggest they may become commercially available.

4. Acid-Treated Metakaolins

Lussier has recently described another interesting bottoms cracking component (83,84). Kaolin clay calcined to metakaolin and then treated with HCI at reflw conditions to remove alumina can generate high surface areas in the leached metakaolin even when only small amounts of alumina are removed. The treated metakaolin possesses both stable surface area and substantial Bronsted acidity. The surface area, stability, and activity of the treated metakaolin and average pore size depend upon clay calcination temperature and the amount and temperature of the HC1 treatment. Alumina solubility and surface area are greatly reduced if the clay calcination temperature varies outside a preferred range. If the amount of HC1 used to treat the clay is excessive, materials with high surface area will result, but the treated metakaolin will not be particularly active because of excessive removal of aluminum.

Treatment temperature also affects average pore sue of the final clay. The active component in the treated clay is believed to be a surface silica-alumina species, that is tetrahedrally coordinated. Because some polymerized alumina species are likely tc be produced in solution under mild leaching conditions, the treated clay slurry could also provide some of the binder (ACH) for the final catalyst. Clay treated by this procedure is believed to be present in some commercially available catalysts.

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C. TheC lav Portion of the Matrix

Clay is added to most FCC catalysts as an inert densifier, that is, to improve the apparent bulk density (ABD) of the catalysts without having an effect on catalyst activity and selectivity. In all known FCC catalysts, the clay of choice is kaolin. Its ability to form high solids pumpable slurries, low fresh surface area, and ease of packing as a result of its platelet structure make kaolin particularly suitable. Most manufacturers purchase the kaolin as a 60 to 65 wt-% slurry (usually dispersed with sodium silicate or tetrasodiumpyrophosphate) and use it without further processing.

Some parameters of the kaolin are critical for FCC manufacturing. Speci- fically, the kaolin particle sue must be small to ensure that the resulting FCC catalyst possesses good ABD and attrition resistance. Average particle sues (by sedigraph) of 0.3 to 0.4 p with a 90% point of approximately 1 p are normal for FCC use. The iron and titania content of the clay can also be important. High iron or titania levels can lead to undesirable secondary reactions, such as gas and coke formation and increased CO combustion in the regenerator, when such clays are used in the manufacture of cracking catalysts. Titania levels below 3.0 wt-% and an iron content of 0.4 to 0.8 wt-% are considered acceptable. Because of environmental concerns, the crystalline silica content of the clay has also become an important parameter. Although crystalline silica is present in very small concentration in clay, recent government regulations have increased the importance of minimizing its presence.

IV. FORMATION OF THE COMPOSITE FCC CATALYSTS

Four primary operations are generally necessary in the preparation of composite catalysts once the components described in the section "FCC Catalyst Matrix" are prepared:

Mixing Spray Drying Exchange FinalDrying

If ion exchange is performed prior to spray drying, the post-spray dryer exchange and drying step is eliminated. Also, calcination steps may be involved: for example ACH bound catalysts are calcined directly after spray drying and may or may not be exchanged after the calcination depending on what process scheme is being used. In any of the schemes, the first step is component mixing.

A. Mixing

Rigorous mixing of the various components of the FCC catalyst is required to produce a hard, dense, homogeneous catalyst. The primary consequences of poor mixing are poor attrition and density as well as poor particle morphology. However, stratification of the active components caused by incomplete mixing can also significantly effect the activity and selectivity of the final catalyst. Little detail is

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given in the published literature as to the type of mixers used in various manufacturing schemes, and the theory of two or three fluid mixing is beyond the scope of this chapter. Mostly, the type and configuration of mixers used for a given FCC catalyst manufacturing process are arrived at over many years of experience. Generally the mixers are high shear because of the thkotropic nature of many of the slurries used in FCC manufacture. Also stirrer blades are designed to induce significant up and down mixing. Often recycle pumps or colloid mills are used to aid in top to bottom mixing.

B. S~rav Drving

Spray drying is the process step in FCC catalyst preparation that is responsible for the formation of the fluid particle itself. Of all the process equipment used to prepare FCC catalysts, the spray dryer is arguably the most impressive. It is large in size and is expensive. To a large extent, the ability of an FCC unit to circulate catalyst depends on how well the spray drying operation is done.

Two main types of spray dryers-mixed flow and parallel flow--are commonly used along with three different methods of catalyst slurry atomization into the hot dryer chamber. The process itself is basically simple: a slurry containing the ingredients of the cracking catalysts is atomized through either a spray nozzle or a fluted wheel rotating at 6000 t rpm into a heated chamber, where the aqueous medium is rapidly evaporated. A porous microsphere of thoroughly mixed catalyst ingredients is left after evaporation. The sphere size and density and the degree of sphericity are controlled by a variety of slurry feed properties and the atomizing technique.

Perhaps the most important control of particle size is the solids content of the spray dryer slurry feed and its viscosity. All else being equal, the particle size is directly proportional to the solids content of the feed slurry. However, control of the particle size, density, and sphericity characteristics also depends on the size and shape of the drying chamber as well as the atomization procedure used. A Boltzmann distribution of particle sizes is invariably obtained around a mean, which is usually set at approximately 70 p average particle size (APS). The APS is controlled by a variation in the slurry feed properties to the dryer and in the conditions of atomization.

Suppliers of FCC catalysts are generally asked to provide catalysts containing controlled particle size distributions (PSD) as a function of the weight percent catalyst contained in various particle size groups (Table 3).

Considerable PSD control is available, and different FCC unit configurations require different APS (and density) catalysts to fluidize well. In many cases, usually based on refinery experience, APS is carefully specified by the FCC catalyst user.

Atomizers used in FCC catalyst spray dryers fall into three general categories:

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I

Table 3 Weight Distribution of Microspheres

within a Particle Size Range

Particle Size Range, 1 Catalyst, wt-%

0-20 0-3

II 0-40 10-14

II 40-80 I 48-60 II II 80-100 20

II 105 t 5-10

two-fluid nozzles, single-fluid nozzles, and wheel atomizers. Two and single-fluid nozzles depend on pressure to atomize the spray dryer feed slurry. In the two-fluid case, air pressure is used to rapidly revolve a disk inside the nozzle when the slurry is passed into the nozzle chamber. The action of the revolving disk atomizes the slurry. The single-fluid nozzle is quite different. Here, hydraulic pressure (800t psig) is used to impinge the slurry on a fixed plate inside the atomizing chamber. The extreme force in the constrained volume atomizes the slurry.

Wheel atomizers also differ further yet in that their ability to atomize the feed slurry is dependent on the high speed at which the fluted wheel discharges the slurry into the drying chamber. Wheel speeds at the tip approach sonic velocity where the slurry is released. The result is droplet formation at that location. Variation in wheel speed, not surprisingly, has been shown to control catalyst particle size.

Both high pressure and wheel atomization are widely practiced commercially. Two-fluid atomization is used in many pilot plant and laboratory sized spray dryers. Commercial dryers sized to process 300 or more tons per day are thought to be in use. Because the overall worldwide FCC catalyst production capacity is estimated to be approximately 1800 tons/day, manufacturers have a strong incentive to use large dryers.

C. Ion Exchanee and Calcination

A 5.0 Si0,-A,O, ratio NaY zeolite has 55 exchange sites per unit cell, some 68% of these sites are accessible from the supercage (85). Approximately 32% of the sites are located in inaccessible places within the hexagonal prisms and beta cages (86). Removing most of the sodium from the zeolite is necessary in FCC

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catalysts because the sodium both poisons active sites and makes the zeolite much less hydrothermally stable.

Commercially, most of the sodium is removed by ammonium exchange using either nitrate or sulfate salts. In many cases, the zeolite is also rare earth exchanged, generally with a rare earth chloride solution. Rare earth stabilizes the zeolite, with respect to dealumination, when the zeolite is exposed to heat or steam. Increasing levels of rare earth result in higher equilibrated unit cell sizes (lower equilibrium silica-alumina ratio). This lower silica-alumina ratio in turn leads to higher catalyst activity, more hydrogen transfer, and a more paraffinic product stream (87). Prior to the phaseout of leaded gasoline, most FCC catalysts contained high levels of rare earth, which caused them to equilibrate at large unit cell sizes, Such large cell sizes result in high activity and gasoline yield. With an increasing need for high octane gasoline, rare earth levels were gradually decreased so that the unit cell size could be lowered and the hydrogen transfer reactions reduced to increase olefin yields and gasoline octane. This trend is expected to continue. Unit cell size and its relation to catalyst activity and selectivity is described in detail by Scherzer in Chapter 5.

After ion exchange and drying at 300-400°F, the zeolite or catalyst is usually subjected to some form of calcination to induce the migration of sodium ions from the inaccessible hexagonal prism sites to more accessible sites in the supercage. Sodium in supercage sites can be readily removed by exchange ions. Calcination is achieved in large rotary kilns, which can be fired either direct or indirectly. After calcination, a final exchange with either ammonium or rare earth ions is done to remove residual sodium ions. Fully exchanged and calcined Y zeolites contain from 0.2 to 1.0 weight percent sodium oxide.

1. In Situ and Ex Situ Exchange

Ion exchange of zeolite as applied to FCC catalysts can be done either in situ, with the exchange performed on the whole catalyst, or ex situ with the exchange done on the zeolite itself prior to inclusion in the matrix. The former type of exchange is required for catalysts prepared by in situ crystallization, in which pure unexchanged zeolite is never isolated. Also, catalysts prepared using silica sol binder derived from acidified sodium silicate require in situ exchange.

Silica sol binder is quite high in sodium and sulfate content and the exchange of the zeolite prior to inclusion in the matrix is done only if required to produce dealuminated zeolite. For instance, preparation of dealuminated zeolite via AFS technology requires formation of NH4Y prior to AFS treatment (16). Because back exchange of the sodium in the binder into zeolite exchange sites occurs, in situ exchange of the catalyst is also required.

Likewise, the rare earth exchange of the zeolite would be done in situ when using silica sol binder to prevent the back exchange of the sodium onto sites occupied by rare earth ions. The formation of pink salt (a sodium rare earth sulfate) may also occur. This pink salt renders a portion of the rare earth useless in terms of catalyst activity or selectivity.

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In situ exchange of catalysts also has some significant benefits in terms of material handling. Since the relatively large particle size of FCC catalyst (APS, approximately 70 p) makes its filtration rate quite fast, most commercial FCC catalyst manufacturers perform Fitration and ion exchange on belt filters. In situ exchange allows the filtration and exchange to be accomplished at significantly higher rates per square foot of filter area than would be the case if the zeolite alone were being treated.

In situ exchange does require sufficient catalyst microsphere green strength to allow stirring and pumping of aqueous solutions without significant breakup of the particles. Sufficient green strength is not present, when ACH or peptized alumina binder systems are used. Microspheres prepared using either of these binder systems are quite weak and require calcination to obtain sufficient green strength to be handled in slurry form. Zeolite Y with high sodium content would suffer significant surface area loss if calcined at temperatures high enough to produce adequate microsphere strength. Therefore, the zeolite is given appropriate ammonium, rare earth, and dealumination treatments prior to inclusion in the matrix.

In the case of peptized alumina binder, the matrix components are used in ammonium or rare earth exchanged forms, and spray drying can be the final step in catalyst preparation. A final calcination of the microsphere is known in some cases to improve microsphere attrition resistance. However, this improvement may also be accomplished by the high temperatures (> 1350°F) present in some FCC unit regenerators. When ACH is used as binder, some further treatment of the microsphere is required to remove the chloride ion present in the matrix. Chloride ion presents potential corrosion problems to the FCC catalyst user and also tends to reactivate agglomerated nickel on the FCC catalyst. Such reactivation increases gas and coke yield. Generally, chloride is lowered to <0.2 wt-% by calcination as the last preparative step.

The ex situ method of exchange has the advantage of better control over the degree of exchange of the zeolite. With in situ exchange some of the sodium removed by exchange or calcination is trapped on the alumina rich matrix after calcination and remains on the catalyst. This trapped sodium does not appear to affect catalyst activity or selectivity, but with in situ exchange, sophisticated techniques are required to determine if the sodium remaining in the catalyst is associated with the matrix or zeolite. Sodium analysis shows the bulk sodium content to be considerably higher than the sodium that is actually associated with a zeolitic site. The deleterious effects of sodium on cracking catalysts have been well known for years. Therefore a distinction must be made between bulk sodium and sodium associated with active sites. The latter plays the more important role in controlling catalyst activity and selectivity.

a. Ammonium Exchange

The theoretical isotherm for ammonium exchange of NaY at 77°F is shown in Figure 6 (88). The thermodynamics and kinetics of ion exchange in zeolite are

130

Figure 6: Zeolite Y Na' -> NH; Isotherm

25'Q 2.83 Si/AI; 0.1 Total Normality 1.01

0.8 Oe9I

NHJ

0.7

0.6

- -

0 0.1 0.2 0.3 0.4 0.5 0.6 0.7 0.8 0.9 1.0 NH,S

beyond the scope of this discussion. However, knowing that ammonium ion exhibits a type D isotherm is important (89,90). The ammonium ion is initially preferred, but the exchange does not go to completion. Xmax (the maximum percentage of zeolite site exchanged. by the exchanging ion) for NH,+ in NaY is around 0.70. This percentage corresponds to exchange of all sodium ions accessible from the faujasite supercage. NH,+ exchange for Na+ is exothermic and is thermodynamically favored at lower temperatures.

In practice, the ammonium exchange of the zeolite or catalyst is done on drum or belt filters and equilibrium is not reached. Hot ammonium ion (140 to 160°F) solutions are employed to aid in the kinetics of exchange. Ammonium exchange after calcination is practiced to remove residual sodium ions residing in inaccessible sites in the hexagonal prisms or beta cages. Because ammonium exchange is done before and after calcination on several belt filters, countercurrent exchange is considered the most efficient.

In a typical countercurrent scheme (Figure 7), fresh ammonium ion solution (usually ammonium sulfate) is used at the end of the exchange train, where the sodium is most difficult to remove. Spent solution or filtrate from this belt is cycled to the previous belt until the spent solution from the first exchange belt is discarded into a waste stream. Water washing is required between ammonium exchanges to remove adsorbed anions. Washing to lower anion concentrations is generally more of a problem when in situ exchange techniques are used and when the anion is sulfate.

131

Figure 7: Counter Current Ammonium Exchange with Continuous Exchange

Waste I :

c

Y

Stream

Reslurry

Ion

b. Rare Earth Exchange

The exchange of sodium or ammonium ions by rare earth impedes zeolite dealumination during hydrothermal treatment. Hence, equilibrium or steam deactivated catalysts containing highly rare earth exchanged zeolites exhibit unit cell sizes near 24.40 A, while very low levels of rare earth exchange result in unit cell sues less than 24.30 A. By preventing alumina loss from the zeolite framework under hydrothermal conditions, rare earth exchange increases activity per weight of zeolite and increases the rate of hydrogen transfer. The resulting loss in octane and olefinicity caused by the rare earth stabilization of unit cell sue has already been discussed. The majority of commercial FCC catalyst sold contains some amount of rare earth on zeolite. The rare earth is present mainly to moderate the activity loss observed in going to completely hydrogen exchanged forms.

Rare earth, which is generally a solution containing lanthanum, cerium, praseodymium, and neodymium with traces of other rare earths can be exchanged directly onto zeolite sites after the zeolite is incorporated in the catalyst before or after calcination. If exchange is done after high temperature calcination or other methods that result in dealumination, the maximum amount of rare earth exchange will be significantly lowered, because the number of exchange sites are lowered by dealumination.

Figure 8 shows the exchange isotherm for lanthanum replacing sodium at two temperatures (91). Again, a type D isotherm is observed. The steep slope at La(s)=O shows the rare earth to be highly preferred at low levels up to

132

Figure 8 La-Na Isotherms at 25°C and 0.1 Total Normality

0 0 0.2 0.4 0.6 0.8 1.0

JwmalolColloklandlntdaceScience s,

approximately 50-60 mole % exchange. Because calcination is required to allow penetration of rare earth into exchange sites in hexagonal prisms or Beta cages, the maximum exchange in the absence of calcination is approximately 70 %. Because of octane needs, no commercial catalysts designed for octane enhancement contain such high rare earth levels on zeolite. Furthermore, exchange is not done under equilibrium conditions, and much of the rare earth exchanges for ammonium rather than sodium ions.

At low rare earth levels based on zeolite exchange capacity, rare earth is highly favored over sodium or ammonium ions, and exchange is close to stoichiometric. At higher levels of exchange, excess rare earth must be used to reach the required Fmal exchange level. If ammonium sulfate is used as the exchange solution, washing of the catalyst between ammonium and rare earth exchange is critical to prevent formation of pink salt.

As with ammonium exchange, elevated solution temperatures (150 to 180°F) are used to improve exchange kinetics. The maximum level of rare earth exchange is also accomplished at higher temperatures. In addition, concentrated rare earth solutions are quite acidic and dilute solutions are generally used to prevent zeolite destruction.

D. Final Drving

In the case of in situ exchange a final drying step is required. This step is generally accomplished in a large rotary kiln. The micropsheres are subject to

133

temperatures of 300 to 450°F and have a loss on ignition (LOI) of 8 to 12% after drying. In the case of zeolite exchange, a final drying step is generally avoided, and the cake is reslurried and added to the spray dryer feed tank.

V. CHARACTERIZATION OF FCC CATALYST

As has been described above, FCC catalyst preparation involves the intimate mixing of many components. The quantity and quality of each component is critical to the physical and catalytic properties of the FCC catalyst. Proper and accurate characterization is thus essential to sufficiently define the product and insure quality catalyst.

Catalyst characterization as used in the industry describes both the physical attributes (analyses) and the catalytic performance characteristics (evaluation) of the materials under investigation. Physical measurements of FCC catalysts combine a wide range of disciplines from macroscopic (bulk) level down to the microscopic (molecular) level, and testing, such as elemental analysis, surface area, particle size,attrition, and bulk density. Catalyst evaluation, on the other hand, utilizes highly sophisticated, custom designed equipment, intended to mimic the commercial FCC process, or parts of the process, but on a laboratory scale. Activity, selectivity, and stability measurements fall into this category.

Because of this complexity, the characterization scientist must deal with an extremely broad range of instruments and techniques to fully define the system under study. In the specific case of FCC catalysts, for example, equipment currently in use, can range in cost from $0.30 glass bottle for water titrated pore volume to over $500,000 for a computerized circulating pilot plant. The fundamental goal of all testing is to provide laboratory information about the catalyst, which can then be used to accurately predict performance of the catalyst in the field.

A. AnalvseS

The first area, chemical analyses, spans the range from simple bulk techniques which yield total chemical composition, through intermediate techniques, that give indications of location of elements within the particles, to molecular methods that yield information regarding relationships among the important catalytic elements.

1. Bulk Measurements

In analyses of FCC materials three bulk techniques of elemental analysis predominate; atomic absorption spectroscopy (AAS), X-ray fluorescence spectroscopy (XRF), and inductively coupled plasma (ICP). Each has its own strengths and weaknesses, and in many respects are complementary.

A A S has been available for many years and its methods are well known. This technique measures the absorption of radiation by an ionized vapor of the sample under investigation. Its major advantages are relatively inexpensive equipment and well developed methodology. Its disadvantages are the need to completely dissolve the sample (very time consuming for FCC catalysts) and correctly match the

134

calibration standard solution to the unknown , and the fact that each element must be determined individually.

ICP uses the same characteristic radiation lines. However, because the temperature of the plasma flame is higher, emission radiation rather than absorption is used. Computer software is available to analyze most elements simultaneously, without the need to match the solution matrix of the standards and the unknown. However, as with AAS, the sample must first be dissolved before the analysis is performed.

XRF differs from ICP and A A S in that it uses emission lines from the innermost electron shells of an element. Because these lines are far removed from the valence electrons (those used in chemical bonds), chemical composition does not affect the analysis. Hence, samples need not be dissolved prior to analysis. In FCC catalyst elemental analysis, simple pressed pellets are suitable for routine determinations. The major disadvantages of XRF are the relatively high cost of the instrument ($150,000+) and the need to match the major element matrix in the calibration standards. Even relatively minor changes in matrix or particle size can effect the analysis. Matrix effects can be eliminated by fusing powder using a flux and high temperature into a clear ‘glass’ pellet. However, this technique is more labor intensive and precludes accurate sodium analysis at low concentrations. The latter is critical in determining catalyst quality.

Another important bulk sample measurement is resistance to particle degradation or as it is known in the industry, attrition index. Such degradation creates fines resulting in loss of valuable product from the stack, and environmental problems such as increased stack opacity. Sources of attrition in a fluidized FCC unit are the air grid, gas bubbles within the bed, the separation cyclones, rotary seals, elbows, and transfer pipe walls (92). There are many different types of attrition testing equipment currently in use to simulate these effects. Some are considered to be proprietary. In all cases, however, the goal is to estimate losses that will occur when the catalyst is used commercially.

Because it is impossible to quantitatively predict the level of losses from a particular unit using these laboratory tests, they serve to give valuable information on relative hardness values among several catalyst types. Forsythe and Hertwig described a simple test which was later modified by Gwynn (Figure 9) and has since been adapted in various forms by most FCC catalyst manufacturers and many catalyst user (93,94). The Gwynn design incorporates a separation chamber which effectively separates the fines (0 to 20 micron material) generated during the test period. This fines separation is important since these same fines will retard further attrition. It also closely approximates what occurs in a typical FCC unit. The equipment has been subsequently updated and utilizes a linear regression analysis of data. Fines generated during the test are collected at 1 hour, 3 hours, and 5 hours, and the cumulative loss from the catalyst bed is plotted versus time. From the regression analysis, the attrition loss rate (weight percent per hour) given by the slope of the line, can be separated from the initial fines and easily attrited material given by the intercept (Figure 10). Although a refinery will be interested in both

135

Fiaure 9: Fluidized Bed Attrition Apparatus (Gwvnnl

3 x 1/641h in. holes I

75 DSI re-d air SUDDIY

I Moore Flow Reaulalor

Rolameler - Figure 10: Typical Katalistiks' Attrition Test

136

numbers, the catalyst manufacturer would like to separate these two types of information as an aid in identification of specific manufacturing problems.

Surface area determination is another bulk technique that has undergone significant changes over the last several years. Total surface area measurement in the FCC industry has relied on the familiar single point BET model developed for macropore materials. However,modern FCC catalysts contain a mixture of catalytically active matrices containing macropores with an active zeolite containing micropores. Therefore, additional techniques are required to fully characterize these systems. The t-Plot method was developed to determine the surface area of the FCC catalysts matrix components or macropores (95). Results can differ from lab to lab depending on the pressure tables and algorithms utilized in the analysis. The difference between the macropore area and the total surface area determines the zeolite surface area. It is then possible with suitable standards to quantify the amounts of zeolite and matrix present in the original fresh sample as well as those same values in equilibrium catalyst or steam deactivated samples. This comparison can reveal the type and severity of deactivation occurring in a given FCC unit. Such information can then be used, along with other data to set up a laboratory deactivation procedure to simulate commercial deactivation (96).

Pore sue distribution measurements (both nitrogen adsorption and mercury intrusion) are an extension of nitrogen surface area measurements. These measurements yield information about the size of the pores within the catalyst particle. Usually the mercury technique is used for pores greater than 200 A in diameter and nitrogen pore size distribution is utilized for pores less than 300 A. Plotting of both the adsorption and desorption cycles can yield information about the shape of the pores as well as their size and frequency. This knowledge, coupled with information on the feedstocks to be cracked (hydrocarbon types, boiling range, and contaminant metals), assist the preparative chemist in designing catalyst for specific applications.

Further bulk physical characterization tests include apparent bulk density (ABD), and total pore volume. ABD measurements are useful in determining hopper inventories, catalyst circulation rates, and fluidization properties (97). The measurement determines the weight of an unpacked volume of FCC catalyst and is relatively simple to perform. Total pore volume is typically determined by a water titration technique. However, nitrogen adsorption measurements can also be used. Nitrogen adsorption has a practical upperlimit of about 600 A. Therefore, this method always shows less pore volume than the corresponding water titration method. Furthermore, nitrogen adsorption is more time consuming and requires more sophisticated equipment.

2. Molecular Spectroscopy

Besides the bulk properties of the catalyst, interactions on a molecular level are important in determining activity and selectivity. Molecular properties include acid site strength and distribution in both the zeolite and active matrix (98,99). Characterization of the interaction of the reactant molecules with the catalyst surface is also possible using molecular techniques (100,101). Methods have been

137

developed to study the oxidation state and method of migration of contaminant metals along the catalyst surface, as well as the reaction of the zeolite and matrix with contaminant metals (102). Obviously bulk techniques will yield very little information on a molecular scale.

A large number of techniques have been developed using standard methods of instrumental analysis to elucidate these interactions. Infrared, nuclear magnetic resonance, Raman, ESCA and other forms of spectroscopy have produced valuable information about reaction mechanisms and surface phenomena. Microscopy techniques such as STEM, TEM, SEM, and EDAX, as well as more sophisticated methods such as SIMS and EXAFS, have also given the FCC scientist valuable tools with which to tailor the catalyst to the needs of the end user. Detailed description of these methods is beyond the scope of this chapter but are dealt with in detail in Chapter 6.

B. Evaluation

Catalyst evaluation in the FCC industry is routinely performed using equipment that can be eight orders of magnitude (100 million times) smaller than the commercial units they are trying to simulate. The long term goal of catalyst evaluation, in the FCC industry, is to simulate commercial cat crackers in the laboratory by using the smallest scale equipment consistent with generation of reliable and accurate data. Accuracy is defined as agreement with commercial performance. Despite these enormous differences, such laboratory studies have led to significant improvements in catalysts performance. These improvements result in higher gasoline yields, higher octane, lower coke make, and better bottoms cracking.

FCC catalyst are evaluated by three test methods in most laboratories. These methods are microactivity (MAT), fixed fluidized bed, and circulating pilot plant. They differ primarily in scale. The MAT is the smallest, and the circulating pilot plant is the largest. They also vary in their thermodynamic and kinetic similarity to commercial operations. The MAT is the least similar to the process, whereas the circulating unit is a small version of the FCC unit. However, most circulating pilot plants are not run adiabatically. Prior to testing a catalyst in one of these units, a fresh catalyst must be pretreated to simulate the deactivation that occurs in commercial use. Generally this consists of steaming the catalyst at various temperatures, times, and steam partial pressures. In addition, the catalyst may be impregnated with nickel and vanadium to simulate the metal deposition that occurs on the catalyst in normal use. A complete discussion of the techniques and protocols used in catalyst deactivation and testing can be found in Chapters 7, 8, and 9.

c. owl litv Assurance

The concept of quality in FCC manufacturing, as with the chemical industry at large, is undergoing radical change. Focus has shifted from the traditional idea of meeting final inspection specifications to controlling process quality at every step in the manufacturing process. In an increasingly competitive global market, the cost of rework or disposal of off-spec material has become prohibitive. Total quality minimizes errors that necessitate disposal or rework and involves everyone in quality

138

issues rather than a traditional quality assurance group. In addition a total quality environment demands communication and interaction between supplier and customer. The measure of performance is meeting the customer’s expectations rather than meeting internal measures of success. Quality audits, quality circles, supplier quality, quality project management, statistical process control, service quality, ISO-9000 and personal quality are all integral parts of modern FCC manufacturing controls.

1. Statistical Process Control

Statistical process control (SPC) is an effort to shift the emphasis from final inspection by the quality control department to the employees responsible for the process. Properly implemented, SPC will not only improve the quality of the catalyst, but also detect problems in the process at an early stage. This approach naturally reduces the amount of off-spec material manufactured. To accomplish these goals, many tools are used: Pareto charts, cause and effect diagrams, scatter plots, designed experiments, and the workhorse of SPC, the control chart.

Many of the principles used for SPC control charts, which are generally based on abnormal distribution, were developed for the manufacture of machined parts, whose measurements are very precise, relatively inexpensive, and fast. Unfortunately, few of these attributes apply to the chemical industry in general and to the FCC industry specifically (103). These difficulties, coupled with relatively short run times, compel the FCC catalyst manufacturer to use control charts of individual measurements instead of the more typical R-bar (range-average) control charts. Nonetheless, with proper understanding, and use of the out-of-control triggers, significant reductions in process variation can be and are achieved. These reductions in process variation lead to more uniform FCC catalysts of higher quality and lower costs.

Figure 11 shows a typical control chart used at the mix tanks just prior to pumping the slurry to the spray dryer. The control characteristic in this example is pH (modified here because the actual values are proprietary). The flow control valves have been calibrated to allow the area operator to make adjustments (clicks) in line with the out of control condition. Because the adjustment rules were determined previously by the engineering group, the operator makes the changes to the process without having to check with supervisory personnel. The operator notes his response on the chart, and he continues to monitor the process with subsequent measurements. An essential aspect of SPC is to control the process by monitoring and adjusting those parts of the process that effect the properties of the fiished catalyst. In the previous example, mix-tank monitoring affects the attrition of the finished catalyst. Using multiple correlation techniques, several other areas were identified as critical and were similarly controlled. Figure 12 shows the results of this effort on the attrition of the fiished catalyst. Not only was the variation reduced (presently half of the variation is due to the laboratory test itself), but the catalyst hardness improved as well. All improvements were accomplished without capital investment or process modification.

139

Figure 11: Process Variable Control Chart

UCL

Target

LCL

Figure 12: Effects of Improved Process Control

0.7

0.6

si g 0 . 5 C 0

*= 0.4 .- CI

2 0.3

0.2

Improvement Transition

Beta Catalyst 162 Shipments

I I I I I I I I I I

0 25 50 75 100 125 150

Shipment Number

140

2. ISO-9OOO

Statistical process control and other quality tools are typically too detailed to share on a regular basis with the customer. A system is therefore needed to ensure that the customer is receiving the quality of product desired. In the past, the customer usually sent an auditor (or auditors) to their suppliers to ensure that the quality system was adequate to furnish the necessary product quality. This practice meant that the manufacturer could expect to have many audits and each audit would be somewhat different. To avoid these problems, a uniform set of minimum standards, the ISO-9000 series, is in the process of being adopted in Europe and to a lesser extent in the United States (104).

The ISO-9000 standards are really a series of standards KO-9001 to ISO-9003 that are targeted at various levels of a business operation. The ISO-9003 standard is aimed at a testing facility whose only products are the test results themselves. The ISO-9002 standard is used to register a manufacturing facility or specific production lines and incorporates additional areas for inventory control and customer interactions. ISO-9001 is the most detailed of the three and adds elements of design or development to 9002 and 9003. A certified auditor visits the site to determine compliance to the desired standard and if successful, registers the audited company. Follow up visits are performed every six months to insure continued compliance. Wide use of ISO-9000 in the FCC industry is already a fact.

VI. SUMMARY

In the preceding sections, the composition and techniques used to prepare FCC catalysts have been described. Essentially four different broad categories of preparative schemes are in commercial use at this time. Figure 13 schematically

Figure 13: Four Commercial Types of FCC Preparation

lnsau Clay sluny

T7 Calcine

9 Caustic Treat (100'F/6Hrs

SPZW

(-1 soor)

180'FH 2-1 4 Hm) v V

dash v V

Filtermash (?)/Dry

NH +/REd(?) Xchng

DrylCalcine

Acid Wash (7) NH;/RE4(?) Xchangewash V

Product

EmiBw!x J-uA NaY Xtal NaY Xtal NaY Xtal

NH,*Xchange (7) NH;Xchange NH,%hange

Mix Binder, Clay NH;,RE4(?)

SP~YDIY Mixxnder Mix %der

v v Dealurnination (7) Dealurnination(?) Dealurnination(?)

Also,, Zeolite XchangeMash XchangeMlash

V v NH;,REd

9

- e - v

Resluny

NH;/REd (?) Xchangemash

Dry/&cine

NH %E" Xchng dasWDw v

Product

etC. V

Spray Dw v Calcine

Further

V Product

V

Processing (?)

141

shows the beginning-to-end preparation of FCC catalysts based on these four schemes. Each of these catalyst preparative procedures has its own advantages and disadvantages. Also a brief description of analytical techniques and statistical process control has been given. The importance of statistical process control in the future of FCC catalyst manufacturing cannot be overestimated. In fact total quality, which includes SPC, will soon be an integral part of all manufacturing based business including FCC. Separate technology to affect catalyst porosity, physical properties, and activity and selectivity have grown up around each catalyst preparative technique. All catalysts must meet the requirements expected of modem FCC catalysts, that is, attrition resistance, high activity and stability, high octane capability, metals tolerance, and the ability to convert heavy oil molecules to lighter products and minimize coke formation. In the future, reformulated gasoline will require changes in FCC catalyst formulations to meet increased demands for light isoolefins for possible conversion into methyl tertiary butyl ether or ethyl tertiary butyl ether.

In addition, ultrahigh activity catalysts will be required. The catalysts will need to operate effectively at extremely short catalyst-to-oil contact times in the riser reactors. Such short contact time risers are currently under study by equipment designers. These new market demands may result in new FCC catalyst formulations in which the overall materials science involved in binders, zeolites, and active matrix components is stretched to another still higher plateau of catalyst preparation.

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71. D. DeCroocq, Catalytic Cracking of Heavy Petroleum Fraction, Golf

72. W. P. Hettinger, H. W. Beck, S. M. Kovach, U.S. Patent, 4,440,868 (1984). 73. D. E. W. Vaughan, R. J. Lussier, J. S. Magee, U.S. Patent, 4,176,090 (1979). 74. J. Shabtai, R. Lazar, A. G. Oblad, Proc. 7th Int. Congr. Catal.,

75. J. J. Pinnanvia, I. D. Johnson, U.S. Patent 4,621,090 (1986). 76. J. J. Pinnanvia, M. S. Tzou, S. D. Landau, U.S. Patent 4,665,045 (1987). 77. R. M. Lewis, R. A. van Saiton, U.S. Patent 4,637,992 (1987). 78. A. P. D. Hopkins, B. L. Meyers, D. M. van Duch, U.S. Patent 4,452,910

79. M. H. Stacey, Catalysis Today, 2, 1621 (1988). 80. J. Sterte, Preprint Div. Petr. Chem., ACS National Meeting, Miami, 1989,

81. J. R. McCauley, U.S. Patent 4,952,544 (1990). 82. G. Jungjie, M. Enze, Y. Zhiging, Canadian Patent 1,267,887 (1990). 83. R. J. Lussier, U.S. Patent, 4,940,531 (1990). 84. R. J. Lussier, J. Catal., 129, 225 (1991). 85. D. W. Breck, Zeolite Molecular Sieves, Krieger Publishing Co., Malabar, FL.,

1984, p. 551. 86. Ibid., p. 553. 87. L. A. Pine, P. J. Maher, W. A. Wachter, J. Catal., 85,466 (1984). 88. D. W. Breck, op. cit., p. 545. 89. F. Helfferich, Ion Exchange, McGraw-Hill, New York, 1982, p. 185. 90. D. W. Breck, op. cit., pp. 531-532. 91. H. S . Sherry in Molecular Sieves I, R. F. Gould, ed., ACS, Washington, DC,

92. British Materials Handling Board, 'Particle Attrition' Duns Tech Publication, 30

93. W. L. Forsythe and W. R. Hertwig, Ind. Eng. Chem., 41, 1200 (1949). 94. J. C. Gwyn, J. AIChE, 15, 35 (1969). 95. M. F. Johnson, J. Catal., 52,425 (1978). 96. J. S. Magee, J. J. Blazek, in Zeolite chemistry and Catalysis, J. A. Rabo, ed.,

ACS, Washington DC, 1976, p. 639. 97. M. F. Ratterman, Oil & Gus J . , 87 (1985).

National Meeting, Washington, DC, 1983, p. 884.

European Patent Application 8,630,021 (1986).

Insr., 49, 504 (1969).

(1984).

Rheinhold, New York, 1960.

Publishing, Houston, 1964.

Elsevier-Kodansha, Amsterdam, 1980, p. 8.

(1984).

p. 489.

1971, p. 357.

(1987).

144

98. H. Stach, R. Wendt, U. Lohse, J. Janchen, H. Spinder, Catal. Today, 3, 431

99. A. Aurox, C. Verdine, Catal. by Acids and Bases, 311 (1985). 100. H. Knozinger in Sugace Organometallic Chemistry, J. Basset, ed., 231, (1988). 101. W. H. McNeese, R. A. Klein in Statistical Method for the Process Industry,

102. V. Cadet, F. Raatz, J. Lynch, C. Marcilly Appl. Gztal., 68, 263 (1991). 103. ASQC Chemical and Process Industries Division Committee, Qualify Assurance

for the Chemical and Process Industry, ASQC Press, Washington (1987). 104. American Society for Quality Control, ANSUASQC QW 1987 Series, (1987).

(1988).

Marcel Dekker, New York, 1991.


145

CHAPTER 5

CORRELATION BETWEEN CATALYST FORMULATION AND CATALYTIC PROPERTIES

J. SCHERZER

UNOCAL Science & Technology Division 376 South Valencia Avenue

Brea, California 92621 U.S.A.

1. INTRODUCTION

The introduction of zeolitic cracking catalysts in the early 1960s, based on the discovery by Plank and Rosinski [l], revolutionized the petroleum refining industry. It resulted in a significant change in catalyst performance and in process technology, which lead to a dramatic increase in the profitability of the FCC process.

Compared to previously used amorphous catalysts, zeolitic cracking catalysts have the following characteristics: 1) high activity; 2) good activity retention (longer useful life); 3) good thermal and hydrothermal stability; 4) high gasoline yields; 5 ) low coke and gas yields; 6) good attrition resistance. Some of these catalysts have also good resistance to contaminants, such as metals, sulfur and nitrogen-containing compounds.

A comparison of product yields obtained over amorphous and zeolitic catalysts is shown in Table 1 [2]. The data show that at constant coke make,the zeolite catalyst gives considerably higher conversion and gasoline yield, while the dry gas and heavy cycle oil yields are reduced. However, the gasoline obtained with the zeolite catalyst has a lower octane rating, due primarily to lower olefinicity.

2. COMPOSITION AND CLASSIFICATION OF FCC CATALYSTS

Modem FCC catalysts consist, in general, of two major components: zeolite and matrix. Some catalysts also contain a third component: one or several additives, designed to boost gasoline octane, increase catalyst metal resistance, reduce SO, emissions, or facilitate CO oxidation. The additive can be incorporated into the catalyst particle or be used as a distinct physical particle [3].

146

TABLE 1

Comparison of Product Yields at Constant Coke Obtained Over Amorphous and Zeolite Catalysts[2]

Amorphous Zeolite Si02-A1203 catalyst Difference

Coke, wt% FF 5.0 5.0 0 Conversion, vol% 77.0 86.0 + 9

c3, vol% FF 1.4 1.6 + 0.2 Cf, vol% FF 14.2 14.0 -0.2 i c 4 , vol% FF 6.8 8.9 +2.1 n-C4, ~01% FF 0.6 0.8 +0.2

gaso., vol% FF 55.5 65.5 + 10.0

H2,wt% FF 0.07 0.04 - 0.03 c1+ c2, wt% FF 3.1 2.7 - 0.4

C;, vol% FF 9.9 8.5 -1.4

LCO, vol% FF 5.0 4.3 -0.7 HCO, vol% FF" 18.0 9.7 -8.3 Gaso. RON + 0 94.0 91 .O -3.0

"HCO = 100 - Cow. - LCO.

The zeolites used in FCC catalysts are mostly synthetic, faujasite type zeolites: Y and high-silica Y zeolites. In the past, X zeolites have also been used, but have been replaced by the more stable Y zeolites. Some commercial catalysts contain mixtures of Y and high-silica Y zeolites. The zeolites are being used mostly in the rare earth or ammonium exchanged form. Most commercial FCC catalysts contain between 15 and 40 percent zeolite, which is the major contributor to the catalytic activity and selectivity of the FCC catalyst.

The catalyst matrix consists usually of a synthetic and a natural component. The synthetic component in most commercial catalysts is amorphous silica, alumina or silica- alumina, while the natural component is clay. Thermally and/or chemically modified clays are also used. The matrix is responsible primarily for the physical properties of the catalyst, although it can also have a catalytic role. For example, FCC catalysts used for cracking heavy feedstocks (e.g., resid) usually have catalytically active matrices.

The additives show a wide variation in composition, depending on their role. For example, ZSM-5 zeolite is used as an octane-boosting additive. Antimony, bismuth and tin compounds are used for the passivation of heavy metals, such as nickel and

147

vanadium, deposited on the catalyst during the cracking of metal-containing feedstocks. Certain metal oxides (Al,O,,MgO) or mixed metal oxides, such as magnesia-alumina, mixed rare earths oxides-alumina or ceria-spinel, are used to control SO, emissions from the regenerator. Platinum is used as a CO combustion promoter. The composition of FCC catalysts, as well as the raw materials used in their manufacture, are shown schematically in Fig. 1.

RAW MATERIALS

SILICA

ALUMINA

SODIUM HYDROXIDE

RARE EARTH CHLORIDE

AMMONIUM SULFATE

- - - - - -

CLAY

ALUMINA

SILICA

ALUMINA

PLATINUM

RARE EARTH

ANTIMONY, ET. AL.

INTERMEDIATE FINAL PRODUCTS PRODUCT

1 10.50%

1- 1-

. \ CATALYTIC PROPS. = I FCC 1

PHYSICAL PROPS. SOME CAT. PROPS.

0 . 1 0 %

CO COMBUSTION

MET. RESIST. OCTANE BOOST

Figure 1 Composition of FCC Catalysts.

Regarding their main applications, commercial FCC catalysts can be broadly classified in three categories: (1) gasoline FCC catalysts, (2) octane FCC catalysts, and (3) resid FCC catalysts. Gasoline catalysts are used when the major objective is to maximize gasoline yields. Octane catalysts are used when the objective is to maximize gasoline octane or octane-barrels (the product between gasoline yield and octane number). Resid catalysts are used to crack resid feedstocks.

148

Gasoline FCC Catal st =F= large pore FAU

*Clay *REY (%RE203 >13) 0 ~ 1 2 0 3 Modified

*Si02-A1203

*RE, HY (10-13%RE203) *sio2

*(REX)

Figure 2 Composition of gasoline FCC catalysts [3].

I OCTANE FCC CATALYST 1

ZEOLITE COMPONENT MATRIX COMPONENT

FlplFqFl (4 < 6 h

HSY ZSM.5 SiOZ CLAY

RE, HY CLAY

(RE OR H-FORM) ZSM.11 A1203 MODIFIED

SAPO.1: S iO~A1203

ET AL.

Figure 3 Composition of octane FCC catalysts [3].

149

Resid FCC Catalyst 4s-l Zeolite: large pore large pore,

Additives: 'i Sb, Bi, Sn, &03, MgO,

Activated Clay c ~ T ~ o , , et al. CeOdspinel, PtlAlzO~, el al.

Pt,Pd

Figure 4 Composition of resid FCC catalysts [3].

The composition of these three categories of catalysts is shown in Figs. 2, 3 & 4. The zeolite component of the gasoline catalysts is commonly a rare earth, hydrogen Y zeolite (RE,HY), with a rare earth content between 10 and 13 percent RCO,. Such zeolites will give high gasoline yields. However, they will also generate high coke yields. The matrix plays usually a minor catalytic role in these catalysts and is therefore catalytically inactive (e.g., silica/clay matrix) or of moderate activity.

Octane catalysts contain mostly high-silica Y zeolites in rare earth or hydrogen exchanged form. They can also contain rare earth, hydrogen Y (RE,HY) zeolites with a lower rare earth content. Furthermore, an octane-boosting additive, such as ZSM-5 zeolite, is sometimes present in such catalysts. The matrix is catalytically active and usually contains amorphous alumina or silica-alumina, in addition to clay.

Resid catalysts usually contain zeolites similar to those present in octane catalysts: rare earth exchanged, high-silica Y and rare earth, hydrogen Y zeolites. The rare earth content of these zeolites is often high. The catalysts contain a large pore, active matrix, in which the active component is amorphous alumina, silica-alumina and, in some instances, modified clay. Furthermore, such catalysts contain metal passivators or traps, SO, abatement additives and CO combustion promoters. The metal passivators, additives and combustion promoters are generally blended with the FCC catalyst, often by the refiner himself.

3. THE CATALYTIC PROPERTIES OF FCC CATALYSTS

3.1. General characteristics of zeolite catalysts All FCC catalysts, regardless of their specific applications, are designed to have the

following catalytic properties: activity, selectivity and stability. Catalytic activity is due to the presence of acidic sites in the catalyst. It is determined by zeolite content and

150

type, by matrix type, and by the zeolite/matrix activity ratio. Catalytic selectivity is determined by several factors: zeolite type; acid site type (Broensted and/or Lewis), strength, concentration and distribution; pore size distribution in both matrix and zeolite; matrix surface area and activity; additives present; and contaminants present. Although the requirements are often seasonal and can change from refiner to refiner, a majority of refiners require catalysts with the following selectivity characteristics: high gasoline and light cycle oil selectivities, low coke and dry gas make, low decant oil make. Catalyst stability, which determines the catalyst longevity in the unit, is affected by both composition and structural characteristics of the catalyst components. Furthermore, process conditions in the FCC unit and feedstock will also affect catalyst performance.

In addition to the general catalytic properties described, it has already been shown that catalysts are usually designed to meet specific requirements, such as maximizing gasoline (gasoline FCC catalysts), maximizing octane-barrels (octane FCC catalysts) or being able to crack resid feedstocks (resid FCC catalysts). FCC catalysts can also be designed to maximize the yields of light hydrocarbons to be used in the making of alkylate and ethers.

I I

60 80 100 CONVERSION (WT%)

Figure 5 General correlation between product yields and conversion

151

The general correlation between product yields and conversion obtained with a zeolite catalyst is shown in Fig. 5. An increase in conversion results in an initial increase in gasoline yield, followed by a decrease due to overcracking. The LCO yield follows a similar pattern. Coke, dry gas, and C,+C, yields increase, while HCO yields decrease with increased conversion.

The PONA analysis of a gasoline obtained with a zeolitic FCC catalyst is illustrated in Fig. 6 [4]. C, hydrocarbons have the highest concentration among paraffins and olefins. As the molecular weight of these hydrocarbons increases, their concentration decreases. Among aromatics, C, and C, have the highest concentration.

.- Olefins Naphthenes Aromatics

10

a

6

4

2

0 5 7 9 1 1 5 7 9 5 7 9 6

i. 10

Carbon number

Figure 6 PONA analysis of gasoline obtained with zeolite FCC catalyst Astra-378. Conversion: 72 LV%; Feedstock: K-factor= 11.89; API gravity=27; Aniline pt:=9O0C.

For any given catalyst formulation, the change of process parameters during catalytic cracking will change the resulting conversion and product yields. For example, conversion increases with reaction temperature or catalyst-to-oil ratio and decreases with increasing space velocity.

3.2. Role and properties of zeolite

3.2.1. Zeolite type and composition The zeolite is primarily responsible for the catalyst's activity, selectivity and stability.

Most commercial FCC catalysts contain Y or high-silica Y zeolites, mainly in rare earth and/or ammonia-exchanged form. Conventional Y zeolites are synthesized with a framework SiOJA&O, ratio between 3 and 6 (usually about 5 ) while high-silica Y zeolites have a higher SiOJAl,O, ratio. Conventional Y zeolites are present primarily in gasoline FCC catalysts, whereas high-silica Y zeolites are used in the manufacture of octane and resid FCC catalysts.

152

Conventional Y zeolites are used in FCC catalysts mostly in the form of rare earth, hydrogen Y (RE,HY) zeolites. Such zeolites are prepared from NaY zeolites by ionic exchange with rare earths and ammonium salts in solution. The commercial rare earth salts are a mixture of lanthanum and cerium salts, with smaller amounts of neodymium and praseodymium. The presence of rare earth in the zeolite increases its stability and catalytic activity. The improved stability is attributed to the formation of polynuclear, rare earth containing hydroxy complexes in the zeolite sodalite cages. The improved activity is due to the higher number of Broensted acid sites, resulting from the partial hydrolysis of hydrated rare earth ions: RE(OHJ3+ + REOH" + H'.

More than half of FCC catalysts manufactured in the U.S.A. contain high-silica Y (HSY) zeolites. Such zeolites are obtained by partial dealumination of conventional Y zeolites. Good-quality HSY zeolites are prepared from well-crystallized NaY zeolites having a SiO&4lZO, mole ratio of 5 or higher. HSY zeolites used in FCC are made commercially by one of the following methods: a) calcination under steam of partially ammonium exchanged Y (NH,,NaY) zeolite, leading to the formation of ultrastable Y (USY) zeolite [ 5 ] ; b) acid leaching of USY zeolite, to obtain an ultrastable zeolite free of non-framework aluminum [6]; and c) treatment of NH,,NaY zeolite with a solution of (NH,),SiF, ( A F S ) , leading to the substitution of some framework aluminum with silicon from the reagent (AFSY zeolite) [7]. Other known methods described in the literature are usually confined to small-scale laboratory preparations.

HSY zeolites have a series of common properties, regardless of preparation method used (see Table 2). Other properties, such as the presence or absence of non- framework aluminum, pore system, stability or composition gradient, are strongly affected by the preparation method used.

TABLE 2 GENERAL CHARACTERISTICS OF HSY ZEOLITES

(VS. PARENT Y ZEOLITE!

1. INCREASED S102'A1203 RATIO IN FRAMEWORK

2. DECREASE IN UNIT CELL SIZE

3. DECREASE IN ION EXCHANGE CAPACITY

4. INCREASE IN THERMAL HYDROTHERMAL STABILITY

5. PREDOMINANT Si(0Al) GROUPS

6. COMPOSITION GRADIENT

7. INCREASE IN 21-1 VALUE OF XRD PEAKS

8. INCREASE IN FREQUENCY OF IR LATTICE VlBRATlONb

9. DECREASE IN INTENSITY OF ACID OH-BANDS IN I R SPECTRUIJO

10. DECREASE I N TOTAL ACIDITY

1 1 . INCREASE IN STRONG ACIDITY

12. DECREASE IN CATALYTIC SITE DENSITY

153

3.2.2. Zeolite activity and content RE,HY zeolites with over 10 wt% RE20, are highly active and are used primarily

in gasoline FCC catalysts. Such zeolites boost gasoline yields and make only a minor contribution to octane enhancement. As the rare earth content of RE,HY decreases, the contribution to octane enhancement increases, while activity and gasoline make decrease. Most commercial RE,HY-based catalysts contain between 15 and 25 percent zeolite.

HSY zeolites have a lower concentration of acid sites than conventional, rare earth exchanged Y zeolites. The lower acidity makes them catalytically less active. To compensate for the lower activity, HSY zeolites are used in considerably larger amounts resulting in catalysts with zeolite contents up to 35 or 40 percent. The activity of HSY zeolites can be increased by exchanging small amounts of rare earth into the zeolite IS]. Although less active than conventional Y zeolites, HSY zeolites retain a higher percentage of their initial activity under severe hydrothermal treatment, due to better hydrothermal stability. However, at high zeolite content the attrition resistance of many catalysts deteriorates. Therefore, when formulating a high-zeolite catalyst, the potential impact on attrition resistance should be taken into consideration.

3.23. Zeolite acidity Both Broensted and Lewis type acid sites are present in Y zeolites used in FCC

catalysts [9,10,11]. Broensted acidity is due primarily to acidic hydroxyl groups attached to the framework, while Lewis acidity is attributed mostly to non-framework aluminum species [12]. In a HY zeolite, the number of acidic hydroxyl groups attached to the framework is equivalent to the number of framework Al atoms. Amorphous silica- alumina present in steamed zeolites also contributes to zeolite acidity [ 131.

The concentration, strength and distribution of acid sites in the zeolite play a key role in determining its activity and selectivity. Although the reactions that take place during catalytic cracking of gas oil are rather complex and many of the primary products undergo secondary reactions (Table 3), most of these reactions involve carbocations as intermediates [14,15,16]. The initial carbocations can arise either at Broensted or Lewis acid sites on the zeolite. While cracking reactions require the presence of strong acidic sites, some secondary reactions such as isomerization, cyclization, and intermolecular hydrogen transfer take place at weaker acidic sites [3].

Hydrogen transfer plays an important role in the gas oil cracking process. It reduces the amount of olefins in the product through bimolecular hydrogen transfer, whereby reactive olefins and naphthenes are converted to more stable paraffins and aromatics [ 17,181. Further hydrogen transfer from aromatics, coupled with condensation and polymerization, can lead to the formation of coke [7]. It is assumed that these bimolecular hydrogen transfer reactions occur between molecules coadsorbed at adjacent acid sites, located at aluminum atoms bonded to the same silicon atom. These aluminum atoms are called next-nearest neighbors (NNN) [19,20]. In Y zeolites, the number of next-nearest neighbors can vary from zero to three (Fig. 7). Upon dealumination, the number of framework Al atoms with 0-NNN increases, while that with higher numbers of NNN decreases (Fig. 8) [20]. While the total acidity of the zeolite decreases with dealumination, the strength of the acid sites associated with the remaining aluminum atoms increases in the order 3-NNNc2-NNNc 1-NNN<O-NNN. The strength of acid sites is further increased by the interaction of framework Broensted sites and the Lewis sites associated with nonframework cationic aluminum species, resulting in "superacid" sites [21,22].

154

- cracking , LPGOlefins

-+ Naphthenes C y c h t i o n

Isomcdzation , Branched HTrlMfCI , Branched olefins paraffins

H Tnnsfer -b Paraffins

C y c h t i m Coke

Condensation ' Dehydrogcnation

TABLE 3 Main Reactions in FCC Catalysis

Paraffm

Olefinsa

Cracking , Olefm

b Cyclo-olefm b Aromatics Dch ydrogcnatim Jkhydmgmstion

Isomuization Naphthenes with different rings

Naphthenes

+ Olefins side-chnin , Unsubstituted m&g aromatics

Tranmlk ylaticn Aromatics I - Different alkylaromatics

Jkhydrogcnation Alkylation I Condensation ' ~ydroaenation' Poly aromatics Coke

"Mainly from cracking. very little in feed.

2-N"

I -Si -

i 0

I 0 - S i - 0

I

0-NNN

I -Si-

I 0 I

0 - S i - 0 I

0 I I

I I I I

0 I

' L o 0 0 0 I I

0 - S i - 0 0-Si-0 I I

0 I

- Si - I

Next nearest neighbor (N") concept. Figure 7

155

A 0-NNN B 1-NNN C 2-NNN D 3-NNN

100

E 4-NNN

24.25 24.35 24.45 24.55 UNIT CELL SIZE, A

Strength of Acid Sites: O-NNN>l-NNN>PNNN, etc.

Figure 8 Acid site distribution in HSY zeolites as a function of unit cell size [20].

The decrease in the number of framework Al atoms and of associated acid sites upon zeolite dealumination decreases the probability of bimolecular hydrogen transfer reactions. It was shown that the lower density of acid sites in USY zeolites compared to that in RE,HY is responsible for the reduced rate of conversion of olefins to paraffins and of aromatics to condensed polycycles [7]. This explains the higher content in olefinic hydrocarbons in the gasoline fraction obtained from gas oil cracking over dealuminated zeolites, as well as the lower coke yield. In addition to reduced acidity, advanced dealumination also reduces the adsorption capacity of the zeolite for olefins, thus further reducing the rate of hydrogen transfer reactions [23].

3.2.4. Framework SiOJAI,O, ratio and unit cell size The catalytic performance of the zeolite is strongly affected by its framework

SiOJAl,O, ratio, since this ratio determines to a large extent the concentration, strength and distribution of acid sites in the zeolite. Due to the correlation between framework SiOJAl,O, ratio and zeolite unit cell size (Fig. 9), and the easy measurement of the unit cell size, the latter parameter is often used to predict the catalytic activity and selectivity of the zeolite [20].

The unit cell size of a good-quality NaY zeolite is about 24.65 f 0.02 A. While fresh RE,HY has a unit cell size in thk range of 24.60 to 24.65 A (depending on SiOJAl,O, ratio and rare earth content), in equilibrium FCC catalysts the unit cell size of the zeolite is usually in the range of 24.50 to 24.55 A. Fresh, commercial HSY zeolite has a unit cell size of about 24.5 A corresponding to a framework SiOJAl,O, ratio of 11 or 12. Upon steaming, the unit cell undergoes further shrinking due to further framework dealumination. The equilibrium unit cell depends upon the severity of steaming and upon the type and number of cations present in the zeolite. Higher steaming severity results in a lower equilibrium unit cell size and lower activity [24]. An increase in rare earth content results in an increase in zeolite activity and equilibrium unit cell size [20,25]. The rare earth content of rare earth exchanged HSY zeolites used in FCC

156

catalysts is usually less than 6 percent RE,O,. Such a rare earth content increases the catalytic activity of the zeolite without a significant boost in hydrogen transfer reactions. Rare earth exchanged HSY zeolites equilibrate at a unit cell size of about 24.30 to 24.35 A. In the absence of rare earth, the HSY zeolites equilibrate at a unit cell size of about 24.25 0.02 A.

The correlation between the number of framework Al atoms per unit cell and unit cell size is shown in Fig. 10. Based on statistical calculations, Peters et. al. [26] have shown that below 24.48 A, the Y zeolite contains primarily isolated (0-NNN) and paired (1-NNN) Al atoms in the framework. The maximum number of isolated Al atoms corresponds to a unit cell size of about 24.30 A (Fig. 10). Such a unit cell size strikes a balance between catalytic activity and the suppression of hydrogen transfer reactions. At lower unit cell sizes the activity of the zeolite is very low due to the low concentration of active sites. At higher unit cell sizes hydrogen transfer reactions will be favored due to the rapid increase in the number of paired Al atoms.

UNIT CELL SIZE (A)

Figure 9 Correlation between zeolite unit cell and structural SiIAl ratio.

157

J -1 w 0

2420 2430 2440 2450

UNIT CELL SIZE, A

Figure 10 Aluminum site density distribution in Y zeolite as a function of unit cell size [26]. (Reprinted with permission from Chemical Engineering Science, Copyright 1990, Pergamon Press)

The decrease in total acidity with decreasing unit cell size results in a decrease in the catalytic activity of the zeolite (Fig. 11,a) [20]. Since the rate of hydrogen transfer reactions also decreases with unit cell size, HSY zeolites favor less hydrogen transfer reactions than REY or RE,HY zeolites. Therefore, the decrease in unit cell size and the corresponding increase in framework SiO&4l20, ratio favors the formation of compounds with high RON and MON in the gasoline fraction (Fig. 11, b and c). It also enhances gas make (Fig. 11,d) and LPG olefinicity, while reducing coke make (Fig. 12,a) [27]. However, the decrease in unit cell size also results in somewhat lower gasoline selectivity. This has been attributed to secondary cracking of olefins in the gasoline boiling range with formation of LPG hydrocarbons. The decrease in gasoline selectivity has also been attributed, in part, to cracking over non-framework aluminum sites [28]. The decrease in unit cell size further results in an increase in LCO selectivity and a decrease in DO selectivity (Fig. 12, c and d), showing that lowering the zeolite unit cell size will facilitate the cracking of "bottoms" to LCO. The better conversion of bottoms accomplished by decreasing the unit cell size is due primarily to the higher severity required to maintain constant conversion with a less active zeolite. The above conclusions are supported by MAT data shown in Table 4, in which product yields and gasoline quality obtained with REY and USY-based catalysts are being compared [29]. The general correlation between zeolite rare earth content, unit cell size and cracking selectivities is shown in Table 5.

158

85, I I I I I I I I ,

c e P d

24.20 24.24 26.28 24.32 24.36

Unit cell size, A

24.20 24.24 26.28 24.32 24.36

Unit cell size, A

95

94

92

91

90

891 I I I I 1 I I 1 I 24.20 24.24 26.28 24.32 24.36

Unit cell size, A

24.20 24.24 26.28 24.32 24.36 Unit cell size, X

Figure 11 Correlation between zeolite unit cell size and catalyst performance [20]. (Reprinted with permission from Journal of Catalysis. Copyright 1984, Academic Press).

The effect of zeolite unit cell size on FCC gasoline composition is shown in Fig. 13 [30]. The feedstock used has the following characteristics: API gravity: 22; Aniline point: 770C; K factor: 11.5. At constant conversion (70 vol%), the increase in equilibrium unit cell size from a low-rare earth exchanged USY zeolite (u.c.=24.26A) to a rare earth exchanged Y zeolite (u.c.=24.39A) results in an increase in paraffins, a sharp decrease in olefins and a slight increase in aromatics.

The effect of unit cell size on hydrocarbon distribution in FCC gasoline is shown in more detail in Fig. 14 [30]. Among paraffins and olefins, the C, molecules have the highest concentration, regardless of unit cell size. As the number of carbon atoms per molecule increases, the concentration decreases. The concentration of paraffins with the same carbon number increases with increasing zeolite unit cell size, while that of olefins decreases.

159

6.0 (a)

4.0 24.20 24.25 24.30 24.35 24.40 24.45 24.50

UNIT CELL (A)

50 t 45 40 L 24.20 24.25 24.30 24.35 24.40 24.45 24.50

UNIT CELL (A)

18.0

17.5 - $ 17.0 L 0 16.5 0

16.0

15.5

15.0

-I

24.20 24.25 24.30 24.35 24.40 24.45 24.50

UNIT CELL (A)

24.20 24.25 24.30 24.35 24.40 24.45 24.50

UNIT CELL (A)

Figure 12 Effect of unit cell size on coke and liquid product yields. Conversion: 65V%; RT=527"C; 4 catloil ratio; 30 WHSV; steam deactivated catalysts 1271

Among aromatics, C, and C, molecules have the highest concentration, regardless of unit cell size. The concentration of aromatics with the same carbon number generally increase with increasing unit cell size.

3.2.5. Sodium content A high sodium content is detrimental to zeolite stability, activity and, in the case of

HSY zeolite, its octane-enhancing capability [20]. The lower activity and octane- enhancing capacity is attributed to the partial neutralization of strong Broensted acid sites by sodium ions. In commercial FCC catalysts, the sodium content of the zeolite is minimized, usually below 1 percent Na,O. This is accomplished by ionic exchange of the zeolite with solutions of ammonium and/or rare earth salts. However, there are reports that the presence of small amounts of sodium in the zeolite may have a beneficial effect on stability [31] and selectivity [32].

160

Figure 13 Effect of zeolite equilibrium unit cell size on FCC gasoline composition. Conversion: 70V%, RT=527'C; pilot plant data [30]

PAIR AFFl NS I l l l l l l l l l l

5 12

F 6

z 10

I - 8 6

z w 4 0

Q

2 2

8 0

CARBON NUMBER " ' 03 ' 0 4

NAPHTHENES

OLEFINS

n c

. . . . 05 0 6 07 0 8 09,

CARBON NUMBER

AROMATICS

CARBON NUMBER CARBON NUMBER

Figure 14 Effect of zeolite unit cell size on PONA distribution in FCC gasoline. Conversion: 70V%, RT=527"C; pilot plant data [30].

161

TABLE 4 Comparison of Yields and Gasoline Quality Obtained

with Conventional REY Catalyst and USY Catalyst [2q

REY catalyst USY catalyst (deactivation: (deactivation: S-13.5 stm.a) S-20 stm.b)

Pilot unit conditions

Conversion vol% Hydrogen, wt% FF

Total C3k, vol% FF

Total C4's, vol% FF C,=/i-C, Ratio

40 WHSV, 4 do, 51O"C,

C1+ Cz, wt% FF

Cy, ~01% FF

C: Gasoline, vol% FF Octane No. RON + 0 MON + 0 Gravity, "API Aniline Pt., "C Bromine No.

Light cycle oil, vol% Gravity, "API Aniline Pt., "C

Gravity, "API Aniline Pt., "C

338°C residue

Coke, wt% FF

MWR feed' 72.5 0.020 1.28 7.9 6.0

13.6 0.83

59.0

86.0 78.0 57.4 31.7 31.0 18.1 18.4 16.7 9.4 3.6

39.4 4.6

72.5 0.020 1.13 9.0 7.6

15.1 0.91

58.0

90.4 80.0 57.3 26.7 50.0 19.5 20.1 23.9 8.0 3.1

34.4 4.0

"732°C 8 h, 100% steam, 15 psig. b827"C, 12 h, 20% steam in air. 'Feedstock: 23.9 "API gravity, 92°C aniline pt.. 11.9 UOP K factor; IBP: 201°C: FBP: 552°C.

3.2.6. Non-framework aluminum species During preparation of HSY zeolites by steam stabilization of ammonium exchanged

Y zeolites, as well as during the catalytic cracking process, framework dealumination takes place. The resulting non-framework aluminum species and some amorphous silica-alumina ("debris") affect catalytic activity and selectivity [28]. Although the catalytic role of nonframework aluminum during gas oil cracking is still debated, it is generally agreed that its effect is both positive (higher activity, higher bottoms conversion) and negative (higher coke and gas make). Some studies claim that the presence of nonframework aluminum in steamed Y zeolites reduces gasoline selectivity [28], while other claim an improvement in gasoline selectivity [33]. The steaming severity of the zeolite will also affect catalyst activity and selectivity, since higher severity

162

TABLE 5

Ciucking Selectivities of Diferent Y Zeolites and ujActive Matrix

ACTIVE USY REUSY REHY REY !MATRIX

I

I I

I I

Unit cell size 4 decrease I

Framework Si/AI 4 increase RE content decrease I +-

Dry yield C3/C4 yield C3/C4 olefins Coke/conversion Gasoline selectivity Octane potential LCO selectivity DO selectivity

Low High High

v. Low Moder. High

Moder. Moder.

Low Moder. Moder. v. Low High

Moder. Moder. Moder.

Low Moder. Moder.

Low High Low Low High

Low Low Low

Moder. High Low Low High

I I I I I I I I I I I I I I I I

High High High High Low High High Low

results in more advanced framework dealumination and a smaller unit cell size, as well as some loss in crystallinity.

In order to reduce the non-selective cracking, some catalyst manufacturers minimize the amount of non-framework aluminum species in ultrastable Y zeolites by acid leaching [34]. Some or all non-framework aluminum species can be removed by such treatment without harm to the zeolite framework. In the FCC unit, further zeolite dealumination takes place, but the total amount of non-framework species in equilibrium catalysts is lower when using acid leached USY zeolites.

The non-framework aluminum content is also reduced by using HSY zeolites prepared by dealumination with (NH,), SiF, [28,35]. While fresh (NH,), SiF, treated Y zeolite (AFSY) has practically no non-framework aluminum, steaming leads to the formation of non-framework aluminum due to hydrothermal framework dealumination [36]. However, the concentration of non-framework "debris" in steamed AFSY zeolites is smaller than in steamed USY zeolites. According to Rabo et. al. [28], catalysts containing steamed AFSY zeolites have better gasoline and light cycle oil selectivity, and make less coke than those made with standard USY zeolites. The better performance of steamed AFSY zeolites is attributed to the lower concentration of non-framework "debris," which catalyze secoadary reactions that result in lower gasoline and higher coke yields.

HSY zeolites with the same unit cell size but prepared by different methods have different activities and selectivities [35,36]. The observed differences have been attributed primarily to differences in the distribution of aluminum in the zeolite. For example, USY zeolites have an aluminum-enriched crystal surface, while AFSY zeolites have an aluminum-deficient crystal surface. Since most cracking reactions occur close to the external surface of the zeolite crystals, the difference in aluminum concentration near the crystal surface can account for differences in selectivity.

The differences in selectivity between USY and AFSY zeolites may also result from differences in the pore size distribution. While USY zeolites have both micropores and mesopores, AFSY zeolites have primarily micropores. Such differences in pore size distribution will result in different diffusion rates of hydrocarbon molecules during the gas oil cracking process.

163

3.2.7. Gasoline RON and MON enhancement The use of HSY zeolites in FCC catalysts gives a significant improvement in

gasoline RON, and a more moderate improvement in MON. HSY based catalysts increase gasoline olefinicity, which enhances RON, but has little impact on MON (Fig. 15) [37]. The increase in olefinicity also increases octane sensitivity. MON is enhanced primarily by increasing branching and aromaticity of gasoline hydrocarbons. It was shown that addition of small amounts of rare earths to HSY zeolites enhances the formation of branched and aromatic hydrocarbons in gasoline [37,38]. Such zeolites equilibrate at unit cell sizes varying from 24.30 to 24.35 A. Low levels of rare earths in HSY zeolites reduce somewhat the RON, but decrease octane sensitivity and boost activity as well as gasoline selectivity. Both RON and MON are further enhanced by the decrease of HSY zeolite crystallite size, since such a decrease favors the formation of branched hydrocarbons at high conversions [39]. The zeolite is often embedded in an active matrix, which enhances gasoline olefinicity and octane. FCC gasoline RON and MON can also be enhanced by using an octane-boosting ZSMJ additive (m infra). However, the correlation between catalyst composition and MON enhancement is not yet well understood and needs further investigation.

In addition to using octane-enhancing FCC catalysts or additives, FCC gasoline octane can also be enhanced by changing certain process parameters in the FCC unit. For example, an increase in conversion, reaction temperature, recycle ratio, or a reduction of the gasoline cut point results in higher octane ratings [3]. An increase in feedstock aromaticity (lower UOP "K" factor) will also increase the FCC gasoline octane rating [40].

96

95

94

93 $1 92

91

90

89

0.5 0.7 0.9 1.1 1.3 1.5 GASOLINE OLEFlNlClTY

Figure 15 Dependence of RON and MON on gasoline olefinicity (commercial data) [37]

3.2.8. Crystallite size The catalytic performance of the zeolite is affected by the size of the zeolite

crystals. By decreasing the size of the crystals, their exterior surface increases. This allows more cracking of large molecules on the zeolite crystal surface to gasoline range

164

products, resulting in higher activity and gasoline selectivity. The diffusivity of hydrocarbon molecules into smaller crystals is also improved [41]. The lower hydrothermal stability of smaller (<lpm) zeolite crystals can be improved by using zeolites with high SiOJAl,O, ratios, such as HSY zeolites [42]. Such crystal size reduction enhances the cracking of resid feedstock, and has a beneficial effect on gasoline octane.

TABLE 6 Effect of Residue on Cat Cracker Yields [43]

Feed + 10%

Gas Oil VTB +20% VTB

Cat/oil ratio over base Conversion: vol% FF Gas yields: wt% FF

c,: wt% Hf: wt% Total C3’s: vol% Total C4’s: vol%

Gasoline: vol% LCO DO C3 + liquid products Coke: wt% FF

Product quality Gasoline

Research clear octane no. Motor clear octane no. Aromatics: vol% Olefins: vol% Saturates: vol%

API gravity Aromatics: vol% Olefins: vol% Saturates: vol% Cetane number

LCO

Decanted oil API gravity

Base 80.2

2.6 0.05

11.6 18.3 62.3 14.8 5.0

112.1 3.7

90.8 78.3 27.0 33.0 40.0

18.9 72.0 0.0

28.0 <20

1.2

77.3

3.0 0.05

10.5 16.1 60.4 16.1 6.6

109.9 4.7

91.4 78.3 29.4 35.0 35.5

19.3 70.3 0.0

29.7 20

1.7

-1.0 75.8

3.2 0.05

10.0 15.2 58.6 16.2 8.1

108.1 6.1

91.6 78.8 26.8 42.9 30.3

22.7 57.9 0.0

34.4 26

6.8

3.2.9. Resid cracking enhancement Addition of resid to FCC feedstock results in reduction of conversion and gasoline

yield, while increasing dry gas, coke, LCO and DO yields (Table 6) [43]. It also results in higher gasoline octane and higher gasoline aromaticity. Heavy metals (especially Ni

165

and V) from the feedstock are deposited on the catalyst, with a deleterious effect on catalyst stability and selectivity.

The deleterious effect of resid on catalytic cracking can be minimized by adjusting the FCCU operating conditions and by using resid FCC catalysts. To be effective, a resid FCC catalyst must have the following characteristics: 1) low coke and dry gas selectivity; 2) high thermal and hydrothermal stability; 3) resistance to deactivation by metals (mainly by Ni and V); 4) resistance to nitrogen poisoning; and 5) ability to crack selectively heavy hydrocarbons.

Resid catalysts usually contain HSY zeolites in rare earth exchanged form, although conventional rare earth exchanged Y zeolites have also been used [44,45]. The HSY zeolite provides the catalyst with the high thermal and hydrothermal stability required to withstand the severe regeneration conditions applied to resid catalysts. The ability to crack the large, heavy molecules from the bottoms fraction is significantly enhanced by using an active matrix. Furthermore, the reduction of zeolite crystal size and the formation of large pores (mesopores) within the zeolite crystal by steam-dealumination will also facilitate the cracking of large molecules. The increase in exterior surface of the zeolite crystals due to size reduction allows more cracking of large molecules to gasoline range products. Forming a network of large pores within the zeolite crystal reduces the diffusional limitations in the crystal.

HSY zeolites have lower coke selectivity and better metal resistance (especially towards V) as compared to conventional Y zeolites. High levels (over 25 percent) of HSY zeolite are used in resid catalysts, in order to provide sufficient activity to the catalyst that would otherwise readily deactivate under resid cracking conditions. Rare earth exchange is used to increase catalyst activity and metal resistance. The sodium content of the zeolite is minimized in order to enhance catalyst stability. The zeolite is incorporated in an active matrix (vide infra).

3.2.10. Zeolite mixtures In some instances, mixtures of several zeolites can be used in the same catalyst

particle, in order to obtain the desired activity and/or selectivity. For example, ammonium or rare earth exchanged HSY zeolites are blended with RE,HY zeolites, in order to enhance catalytic activity as well as gasoline selectivity. By optimizing the ratio between the two zeolites, octane-barrels can be maximized. The optimum ratio depends on feedstock, operating conditions and desired product slate. The RE,HY zeolite can contain different levels of rare earths, which will convey different activities and selectivities to the catalyst. In general, a higher rare earth content in RE,HY results in higher activity and gasoline selectivity, but also in higher coke make. A more detailed description of the effect of HSY and RE,HY zeolites on catalyst performance can be found in [3].

3.2.11. Non-faujasite type zeolites in FCC catalysts FCC catalyst formulations with non-faujasite type zeolites or mixtures of Y zeolites

and non-faujasite zeolites have also been described. The use of ZSM-5 zeolite as an octane-boosting additive has already been mentioned. Several patents describe the use of AlP0,-derived molecular sieves in catalyst formulations, resulting in an increase in gasoline octane rating and higher olefin/parafin (C3+ C,) ratios [46,47,48). Mordenite, by itself or in combination with USY zeolites, can be used as a catalyst component to boost gasoline octane [49]. Offretite [50], silicalite [51] and different members of the ZSM-5 family of zeolites [52] have been reported to enhance gasoline octane and c3+C4 yields. The use of large-pore R zeolite in FCC catalyst formulations results in enhanced

166

gasoline octane and olefinicity, as well as lower gasoline aromaticity [53,54,55]. The selectivity advantages of some of these zeolites over Y type zeolites may prove especially valuable in the EPA mandated reformulation of gasoline.

33. Role and properties of matrix The matrix has a significant impact on the physical properties of the FCC catalyst

and, to a lesser extent, on its catalytic properties. It has the following physical functions: a) the matrix components facilitate the formation of catalyst particles with suitable size and shape during spray drying (primarily microspheres with 9 =60-80pm); b) the matrix acts as a binder for the zeolite particles in the catalyst microspheres, forming a hard, attrition resistant catalyst; c) acts as a diffusion medium for the feedstock molecules and its conversion products, providing good accessibility to the zeolite crystals through the matrix pore system; d) facilitates heat transfer from regenerator to reactor and protects the zeolite from structural damage; e) acts as a sink for sodium ions and other contaminants.

Such matrices can have the following catalytic functions: a) facilitate molecular "traffic" to the zeolite by cracking large oil molecules to smaller ones that are accessible to the zeolite; b) improve bottoms upgrading by cracking the larger molecules in the bottom fraction; c) improve LCO quality by increasing its content in aliphatics; d) improve gasoline octane due to lower hydrogen transfer rate vs. cracking rate; e) improve the metal resistance of the catalyst by cracking the heavy, metal containing molecules, and binding some of the metals, such as vanadium, to the catalyst matrix; f ) improve the nitrogen resistance of the catalyst by reacting with nitrogen compounds from the feedstock; g) in some instances (e.g. in the presence of active alumina) reduce SO, emissions from the regenerator. However, active matrices also enhance low-seledivity cracking, leading to the increase in coke, dry gas and CJC, olefins, often at the expense of gasoline [3].

33.1. Matrix composition Most commercial catalysts have semi-synthetic matrices, consisting of a synthetic

compound (usually amorphous silica, alumina or silica-alumina) and a natural component (usually clay). In catalysts prepared by the in situ zeolitization of calcined clay, the matrix usually has only a natural component (calcined clay).

33.1.1. Amorphous silica containing matrices The silica precursor used in the formulation of catalysts with such matrices is usually

a silica hydrosol. The characteristics of the silica sol used, such as particle size, concentration and counter-ions present can affect the physical properties of the catalyst.

The catalytic role of such a matrix is less significant than that of alumina or silica- alumina matrices. However, upon steaming a catalyst containing a USY zeolite embedded in a silica matrix, silica from the matrix can react with nonframework aluminum from the USY zeolite, forming a catalytically active Si0,-A,O, phase [33]. Furthermore, silicon transport from the matrix to the zeolite during steaming can result in its incorporation into the zeolite framework in the vacancies left by the expelled aluminum atoms, leading to an increase in zeolite stability [56]. As in any other matrix, the pore size distribution in the silica matrix will affect the diffusion rates of hydrocarbon molecules into the zeolite catalyst particles and thus affect catalytic performance.

Many matrices are catalytically active.

167

33.1.2. Active alumina containing matrices Among the different commercial aluminas available, pseudoboehmite is more

frequently used in catalyst formulations. The catalytic activity of pseudoboehmite- derived alumina is due to its acidity. Its presence in the catalyst not only increases catalytic activity, but also improves attrition resistance [57], serves as a metal trap [58], and reduces SO, emissions [59]. Aluminum chlorhydrol is also used as an alumina precursor in catalyst formulations [60].

33.13. Amorphous silica-alumina containing matrices Amorphous silica-alumina is an active matrix component encountered in many

commercial FCC catalysts. The physical and catalytic properties of the catalyst are strongly affected by the composition, preparation conditions and incorporation method of this synthetic matrix component into the catalyst. Its activity is due to the presence of Broensted and Lewis acid sites. Changes in the SiO&Il,O, ratio in amorphous silica- alumina result in activity and selectivity changes [61]. Low-alumina (10-15 percent Al,O,) gels have lower pore volume, lower average pore diameter and are generally less active than high-alumina (-25 percent Al,O,) gels. Recently it was shown that with a heavy aromatic feedstock, gasoline yields and octane numbers are maximized by using catalysts consisting of calcined REY and an amorphous silica-alumina matrix that contains between 70 and 90 percent Al,03 [62].

The formation of meso- and macropores in the gel is favored by synthesizing the gel at high pH, in concentrated solution, in the presence of pore-regulating agents. Similar to other amorphous components, steaming of amorphous silica-alumina results in a partial loss of activity, loss of surface area, collapse of micropores and increase in the average pore radius. Matrices with a high content of large pores are generally more stable under steam. Amorphous silica-alumina matrices can serve as a source of volatile silica species that are transported under steam from the matrix to the zeolite, where they replace the expelled aluminum atoms in the zeolite framework [56].

33.1.4. Matrices with modified clays Some thermally and/or chemically modified clays are active matrix components.

Mild acid leaching of clays removes the contaminating iron, exchanges some of the metal cations with protons, increases the clay porosity and thus makes the clay catalytically more active. Some of the aluminum can also be extracted by acid leaching. Due to their good hydrothermal stability and moderate activity, such acid treated clays have been used in the formulation of different FCC catalysts [63].

Kaolin converted to a spinel/mullite mixture by calcination at 1000oC and subsequently leached with caustic has also been used as a catalyst component, especially in resid FCC catalysts [64]. Catalysts containing acid-leached metakaolin have similar properties.

33.2. Effect on catalytic activity and selectivity Catalyst matrices can be catalytically inactive or active. Inactive matrices often

consist of amorphous silica and clay. In active matrices, the catalytically active component is usually alumina, silica-alumina or modified clay.

In catalysts with inactive matrices only the zeolite determines the activity and selectivity of the catalyst. An inactive matrix is devoid of acid sites and therefore has less impact on catalyst activity and selectivity. However, it has already been shown that the interaction under steam between zeolite and a silica-containing matrix can affect catalyst selectivity and increase zeolite stability. Furthermore, the pore size distribution

168

of the matrix will affect the diffusion rates of hydrocarbon molecules in the catalyst particles.

- z c 35 I 0 W

5 n

(a) d 30 F

f g 35

W

a CI

30

10

- z r c

; 5

f! : 10

8

(b)

W Y

5

Figure 16

/;* , ,

IEY

I I 1

0 ISY

'REUSV P cREyy

r REUSY

I I I

CONVERSION (WEIGHT %)

Effect of matrix type on gasoline and coke selectivity, 'I, U, 0 - active matrix; V, 0, o - inactive matrix. (a) effect on gasoline yields; (b) effect on coke yields. (CREY = calcined REY). [66]

In catalysts with active matrices, the activity and selectivity of the catalyst is determined by both zeolite and matrix. The active matrix contains acid sites usually associated with aluminum atoms. Catalysts with active matrices are often used in FCC units where higher conversion and improved gasoline octane cannot be achieved by operating at higher severity due to unit limitations. It is also used to enhance G + C 4 yields and olefinicity, increase LCO selectivity and decrease DO selectivity (higher bottoms conversion). However, due to some nonselective cracking over the active matrix the gasoline yield may decrease and the coke make increase as illustrated in Table 7 [65] and Fig. 16 [66].

169

TABLE 7 Effect of Matrix on Product Selectivity (MAT data)

REY REY zeolite in zeolite in

active inactive matrix matrix

amorphous low-SA

SA matrix, m2/g Conversion, wt% c3, wt% c;, wt% cq, wt% c,=, wt% CSltot. c3 ca/tot. c4 Gasoline, wt% Coke, wt%

130.0 65.0

1.4 3.2 6.0 3.5 0.70 0.37

43.5 6.2

30.0 65.0 1.3 2.8 5.7 2.9 0.68 0.34

46.5 4.5

Matrix pore size distribution and pore volume also have an impact on catalyst activity and selectivity. A properly designed matrix should have sufficient large pores to allow the diffusion of large gas oil molecules to the zeolite crystals. An active matrix should also have sufficient surface area (>lo0 m2/g) combined with sufficient accessible active sites, in order to facilitate the cracking of the large gas oil molecules. Catalysts with active matrices are commonly used in the cracking of resid feedstock and in octane catalysts.

3.4. Role of zeolite/matrix ratio In addition to zeolite type and unit cell size, the zeolite/matrix activity ratio plays

a key role in the performance of FCC catalysts. The ratio is designed to enhance the positive contribution of the active matrix, while minimizing non-selective cracking. The effect of zeolite/matrix ratio on product yields is shown in Fig. 17 [67]. At high zeolite/matrix ratios the selectivity pattern approaches that of pure zeolite cracking, while at low zeolite/matrix ratios the selectivity pattern is dominated by the matrix. At constant catalyst activity, a decrease in zeolite/matrix activity ratio results in an increase in LCO, coke, dry gas yields, and a decrease in bottoms yields. The gasoline octane and olefidparaffin ratio in LPG also increase under these conditions. At high zeolite content, the increase in matrix activity has little effect on gasoline selectivity.

The cracking of resid feedstock requires a catalyst with low coke selectivity, i.e., a catalyst that gives low coke yields at constant conversion. An increase in zeolite/matrix activity ratio will reduce catalyst coke selectivity. However, such an increase will also reduce the ability of the matrix to crack heavy oil molecules. Therefore, a balanced zeolite/matrix activity ratio is required. The optimum ratio depends on feedstock composition, process parameters and desired product slate.

140

120

100

16

AMORPHOUS CRACKING

LEOLlTElMATRlX SURFACE AREA - OF STEAMED CATALYST - ZEOLITE CRACKING

Figure 17 Effect of zeolite/matrix ratio on product yields at constant conversion (60 wt%, MAT data). Test conditions: 3 cat./oil ratio, 16 WHSV, 500°C reaction temp. Feed: 22.5 MI; 11.5 UOP K [67].

3.5. Role of Additives FCC additives with a cracking or non-cracking function can be incorporated into the

catalyst by the manufacturer, or can be added separately by the refiner. Most refiners

171

prefer the latter approach, since it is more economical to use the additive on an as- needed basis. The additives most frequently used in FCC catalysis are: (a) Octane- boosting additives; @) SO, reducing additives; (c) Metal passivators and traps; and (d) CO combustion promoters. A brief discussion of the role of these additives follows. More details on this subject can be found in chapters 10, 13 and 14 of this volume, as well as in [3].

3.5.1. Octane-boosting additives The most frequently used additive consists of ZSM-5 zeolite embedded in a

matrix, and is commonly used as a distinct physical particle in conjunction with a FCC catalyst [3,68]. ZSM-5 is a member of the pentad family of high-silica zeolites and has relatively small pore openings (4=5.5&. It increases gasoline octane primarily by shape-selective cracking of low-octane, straight chain and monomethyl aliphatics to lighter products [69,70,71]. This results in the enrichment of gasoline in higher-octane branched paraffins, olefins and aromatics. It also entails the loss of some gasoline and the formation of LPG olefins. These olefins are used in alkylation units to make high- octane alkylate, thus compensating for the loss of FCC gasoline.

In addition to cracking gasoline range low-octane components, steamed or equilibrated ZSM-5 also favors the isomerization of cracking products, such as n-olefins to higher-octane branched olefins [69]. The additive represents only a small percentage (usually 1 to 3 percent) of the catalyst inventory in the unit and increases both RON and MON at low concentrations.

3.5.2. SO, reducing additives [72,73] Such additives are used in conjunction with FCC catalysts when cracking high-

sulfur feedstocks, in order to reduce SO, emissions from the regenerator below acceptable limits. The additives are mostly inorganic oxides (e.g. Alz03, MgO) or mixed oxides (e.g. spinels, rare earth oxide mixtures, vanadia-alumina, ceria-promoted spinels). The additive promotes the oxidation of SOz to SO3 in the regenerator, followed by the formation of metal sulfate type compounds on the additive. In the reactor (and stripper) the metal sulfates are reduced to hydrogen sulfide, which leaves the FCC unit with the cracked products and is subsequently removed with an amine scrubber. The additive represents usually a few percent of catalyst inventory and is often added by the refiner shortly before catalyst utilization.

3.5.3. Metal passivators and metal traps These additives are used primarily in resid cracking and are designed to minimize

the deleterious effect on catalyst and its performance by the heavy metals present in the feedstock, such as nickel, vanadium, iron and copper. In the absence of such additives, nickel deposited on the catalyst causes a significant increase in coke and hydrogen make, while vanadium causes partial destruction of the zeolite under steam, in addition to increasing coke and hydrogen make. Nickel is often passivated with antimony [74] or bismuth [75] containing additives, while vanadium is passivated with tin compounds r761. Metal traps are usually physically distinct particles that trap the contaminant metal. Numerous metal oxides (e.g. Al20,, TiOd, mixed oxides (e.g. BaTiO,, CaZrO,) and natural clays (e.g. metakaoh, sepiolite) have been recommended as metal traps [77,78,79]. These materials are usually mixed with a binder and spray dried. The resulting microspheroidal particles are blended with the catalyst.

172

TABLE 8

Characteristics of Equilibrium FCC Catalysts [8O]

K c w u e Gas oil operation operation

Activity (conversion) Carbon factor Hydrogen factor Surface area Pore volume Apparent bulk density Average particle size ' 4 1 2 0 3

Na Fe C V Ni c u Sb

45-75 0.74.0

4-22 50-100

0.3-0.4 0.8-0.99 67-85 30-50

0.4-1.0 0.3-1.0

0.03-0.6 1000-8000 5oo4ooo 20-300 0-1000

65-80 0.5-1.3 0.9-5 .O 70-180

0.2-0.4

60-80 30-58

0.1-0.6 0.3-0.7

0.03-0.3 200-1000 40-500 20-100

0

0.8-1.1

The composition and physicaYproperties of equilibrium catalysts from a residue operation using high-metal feedstock are significantly different from those obtained from a light gas oil operation (Table.8) [80]. In the residue operation the equilibrium catalysts generally have lower activity, higher hydrogen and carbon factors, and considerably higher levels of metals (Na, Ni, V, Fe, Cu) as compared to equilibrium catalysts from light gas oil operations. The lower activity and higher level of contaminants on the residue catalysts require a higher catalyst make-up rate in the FCC unit.

3.5.4. CO combustion promoters To facilitate the oxidation of CO to CO, in the dense phase of the regenerator,

platinum- or palladium-based combustion promoters are being used [81,82]. Maximizing CO oxidation in the dense phase of the regenerator prevents uncontrolled CO oxidation (afterburning) in the dilute phase, which can cause metallurgical damage to the equipment. Furthermore, the heat generated during maximized CO oxidation in the dense phase is transferred to the circulating catalyst, thus providing additional heat for the endothermic cracking reactions in the riser. Maximizing CO oxidation also minimizes flue gas CO emissions.

4. CATALYST DEACTIVATION

During the catalytic cracking process, FCC catalysts gradually lose their activity. Catalyst deactivation with time on stream can be caused by (1) structural changes in the catalyst; (2) coke deposition; (3) contaminants (poisoning). Catalyst deactivation can

173

be reversible, e.g. deactivation due to coke, which can be burned off in the regenerator; or irreversible, e.g. due to structural changes in the catalyst during operation. Deactivation by contaminants can be reversible, e.g. deactivation by nitrogen compounds; or irreversible, e.g. deactivation by sodium or vanadium ions.

4.1. Structural changes in the catalyst Structural changes can occur in both zeolite and matrix. The changes take place

primarily during catalyst regeneration, when the catalyst is exposed to high temperatures (up to 7600C) and steam. Severe regeneration conditions cause zeolite dealumination, followed by loss in zeolite crystallinity. This, in turn, causes a loss in zeolite surface area and acidity, with a corresponding loss in activity. Furthermore, the presence of contaminants such as sodium or vanadium has a deleterious effect on zeolite stability, since both metals can cause a loss in zeolite crystallinity under regeneration conditions.

High temperatures and steam will cause structural changes in the matrix, especially in its synthetic component. Under these conditions, the matrix surface area decreases, small pores collapse and the average pore diameter of the amorphous component increases. In the case of an active matrix, such changes result in a loss of activity.

During operation, a gradual physical degradation of the catalyst particles also takes place. Attrition leads to a gradual decrease in the average particle size and can lead to the eventual disintegration of the particle into powder. This results in catalyst losses, which are compensated by the corresponding addition of fresh catalyst to the FCC unit.

4.2. Coke deactivation The catalytic cracking of FCC feedstock is accompanied by the formation of

carbonatious deposits on the catalyst, which block its active sites and pores. However, the presence of coke is necessary in the FCC process, since coke burning in the regenerator provides the heat required to compensate for the loss of heat during cracking. Over the years, the general trend in the refining industry has been towards FCC catalysts with low coke selectivities, i.e. towards catalysts with reduced coke-making tendencies. Effective octane and resid FCC catalysts have low coke selectivities.

4.2.1. Origins of coke The coke formed in the FCC unit can have the following origins [83,84]:

- -

- -

Catalytic coke produced from the cracking reactions that occur at the acid sites of the zeolite and matrix, Contaminant coke produced by heavy metals deposited on the catalyst (Ni, V, Fe, Cu); Cat-to-oil or occluded coke resulting from the carryover of hydrocarbons in the catalyst pores; Feed residue coke contributed by carbon residue in the feed.

In conventional gas oil cracking, catalytic coke contributes 50 to 65 percent of overall coke. In resid cracking, contaminant and residue coke are major contributors to the coke load on the catalyst.

4.2.2. Nature and role of coke The mechanism of coke

formation is rather complex. It is generally agreed that coke has a pseudographitic structure, obtained by extensive dehydrogenation of polycyclic fused-ring aromatics [85,86]. Other species, with a higher hydrogen content, are also present.

Coke is formed on both zeolite and active matrix.

174

I3 inaccessiba to N2

CJ inaccassible to n-haptona

Figure 18 Schematic representation of coke distribution in the H’, zeolite ,Jr (a) low coke content, and (b) high coke content [88]. (Reprinted with permission from Journal of Catalysis. Copyright 1987, Academic Press).

Zeolite acidity and pore structure affect coke composition, distribution, and coking rate [87,88]. Highly acidic Y zeolites with a 3-dimensional large-pore network and large cavities have a strong tendency to form coke. However, the coking tendency of ZSMJ is very limited, due to its reduced acidity and narrow pores. Shape-selective zeolites such as ZSMJ, erionite and ferrierite have low coke selectivity, i.e., the coke yield obtained with these zeolites at constant conversion is low compared to that obtained with an acidic Y zeolite. The coke deposition on a Y zeolite is illustrated in Fig. 18 [88]. Coke formation is enhanced by certain metal contaminants (especially nickel).

Since coke formation on Y zeolites involves hydrogen transfer reactions, suppression of such reactions (e.g. by decreasing acid site density through increased framework Si/Al ratio) will reduce the coke forming tendency of the zeolite. This explains the low coke selectivity of catalysts containing HSY zeolites. An increase in zeolite/matrix activity ratio can further reduce coke selectivity. Contaminant coke produced by metals is minimized by using metal passivators and low-surface area matrices. A lower matrix surface area reduces the dispersion of the contaminant metal, thus minimizing its deleterious effect. Occluded coke that results from the carryover of hydrocarbons in the catalyst pores can be reduced by increasing the average pore diameter of the catalyst matrix and by reducing matrix surface area. Coke make can be further reduced by adjusting operation variables during the FCC process. For example, reducing contact time between catalyst and feed in the reactor reduces coke make.

43. Catalyst poisoning FCC catalysts are readily poisoned by substances whose molecules react with the

catalyst’s acidic sites, thus inhibiting further reactions at these sites. Catalyst poisoning often encountered in the FCC process is due to basic nitrogen compounds in the feedstock. Such compounds act as temporary poison to FCC catalysts by neutralizing

Catalyst design can have a significant impact on coke selectivity.

175

acid sites, thus reducing catalytic activity. The activity decreases with increasing nitrogen content in feed. At constant conversion, the gasoline yield decreases, while coke, hydrogen and light hydrocarbon yields increase with increased feed nitrogen content [89]. Most of the nitrogen-containing compounds in the liquid product are in the heavy oil fractions (DO and LCO). Some of the nitrogen ends up in coke on the catalyst and is eliminated during regeneration.

The proper design of the FCC catalyst can mitigate the deleterious effect of nitrogen compounds in the feed. To be effective, the catalyst should contain a high concentration of zeolite and an active matrix with a broad pore size distribution. The active matrix will neutralize and crack the nitrogen compounds, thus protecting the zeolite from poisoning.

The effect of sulfur compounds in the feedstock on catalyst activity and selectivity is less significant. However, some of the sulfur ends up on the catalyst, and will be converted to SO, during regeneration.

Alkaline metals present in the feedstock also have a poisoning effect on the catalyst. Sodium ions in the feedstock will neutralize the acidic OH groups in the catalyst and reduce its activity. Sodium ions also have a deleterious effect on gasoline octane.

Catalyst additives are deactivated with time on stream. When using SO, reducing additives, the deactivation consists in reduced ability to capture SO, under regeneration conditions. Metal traps are deactivated with increasing contaminant metal loading. Platinum used as a CO oxidation promoter is partially passivated in the presence of antimony, used as an additive for nickel passivation [go]. Lead and sodium also passivate platinum.

5. DESIGNING FCC CATALYSTS

The design of a FCC catalyst is determined by the following factors: feedstock type, products desired (yields and quality), unit design, environmental impact, and cost [91]. For example, the processing of a resid feedstock requires a different catalyst than the processing of a light gas oil; maximizing gasoline yields requires a different catalyst than maximizing gasoline octane; limitations on the unit air blower capacity or LPG capacity will affect the design of the FCC catalyst; a well designed FCC catalyst will minimize SO, and CO emissions during the cracking process.

When designing a FCC catalyst, both the mechanical/physical and catalytic properties should be taken into consideration. The following mechanical/physical properties are important: attrition resistance, thermalhydrothermal stability, particle size distribution, pore size distribution, surface area, bulk density and X-ray crystallinity [3]. Some of these physical properties, such as thermalhydrothermal stability, pore size distribution and surface area affect the catalytic performance of the catalyst. Other properties, such as attrition resistance, particle size distribution and density affect primarily the mechanical behavior of the catalyst in the FCC unit.

When designing the catalytic properties of a catalyst, catalytic activity, selectivity and stability should be considered. The correlation between catalyst formulation and catalytic performance is shown in Fig. 19.

The figure illustrates the correlation between zeolite Y unit cell size, zeolite/matrh activity ratio and catalytic performance. The catalysts can be divided in several groups, based on composition and performance. Catalysts in group A consist of zeolites with small unit cell sizes and a low zeolite/matrix activity ratio. Such catalysts are best suited for cracking bottoms, since both low-unit cell size zeolites and active matrices favor the

r---------- 1 I CAT. WITH I ! *BEST BOTTOMS !

I A HlGHESTC3 t C 4 I I HIGHEST OCTANE I I I L----,,..---J

CRACKING

E

r---------- 1 I CAT. WITH 1 I GOOD BOTTOMS I

s ! CRACKING ! I LOWERC3+ C 4 I I HIGHESTCOKE I I MAKE I L--,-,,,,--J

1

r - - c ~ ~ ~ ; ~ - - - 'I I MAKE I ! LOWESTCOKE 1

c I MODERATE I I GASOLINE MAKE! . MODERATE I --------,- 7 L - -02TANE - - - - -I

I I

CAT. WITH

I OCTANE-EEL. I

I I

-,,,-,,---J

HIGHEST

r - c ~ ~ i j ~ r - - - - 1 I- HIGHEST GASOLINE! ! MAKE I

L _F_R&W_KING - - - - J

LOWESTOCTANE I D ! LOWESTC3t C4 I

I* LOWEST BOlTOMS

2 3 4

I I I IWER Z/M ACTIVITY RATIO HIGHE ZIM f-- MATRIX ACTIVITY INCREASE Z/M - COKE MAKE INCREASE

Figure 19 Optimized catalyst compositions for specific applications.

cracking of heavy hydrocarbons. Such a composition also maximizes total C3+C4 as well as C3+C4 olefin yields, which is advantageous'when alkylation capacity is available. The low-unit cell size zeolite and active matrix will also maximize gasoline octane. G+C, yields and gasoline octane can be further enhanced by using ZSMJ additive.

Catalysts in group B consist of zeolites with a large unit cell size and an active matrix. Such catalysts are able to crack bottoms, although to a lesser extent than the previous category. They also generate more coke and less C,+C4 hydrocarbons. The catalysts can be used for bottoms or resid cracking when limitations exist on the C3+C4 capacity. Such catalysts have good gasoline selectivity.

Catalysts in group C consist of low-unit cell size zeolites and low-activity matrices. The gasoline selectivity and gasoline octane are moderate. Due to the low coke make of such catalysts, the FCC unit can be operated at higher conversion, which results in higher gasoline yields and better octane numbers. Due to the low-activity matrix such catalysts are not well suited for bottoms cracking.

Catalysts in group D contain active zeolites with relatively large unit cells and low- activity matrices. Such catalysts have the highest gasoline selectivity, but gasoline octane, C3+C4 yields and the ability to crack bottoms are low. Such catalysts are used to maximize gasoline make, but are unfit for bottoms cracking.

Catalysts in group E consist of zeolites with medium-unit cell sizes and matrices of medium activity. Such catalysts, which have a composition and properties between those of group A and D, are designed to maximize octane-barrels. The catalysts have a moderate ability to crack bottoms and give moderate yields of C3+C, hydrocarbons.

The catalyst selection is often determined not only by the desired product yields and quality, but also by unit constraints [92]. For example, when blower capacity is a

177

limiting factor and the objective is maximizing gasoline yield and octane, a catalyst from group C is best suited. Due to its low coke make, such a catalyst can be run at high conversion, resulting in higher gasoline yields and octane. When alkylation unit capacity is the limiting factor ( q + C 4 yield limitation) and the objective is maximizing gasoline yield and octane, a catalyst from group E is selected. (Catalysts from group D would maximize gasoline and minimize q + C 4 yields, but would also minimize gasoline octane.) Similar considerations apply when compressor limitations exist, requiring the reduction of dry gas and q + C 4 yields.

The environmental impact of the FCC catalyst during its use in the cracking unit is of considerable importance. A well designed catalyst will not only reduce SO, and CO emissions, but will also contribute to the making of quality products, which have a diminished environmental impact (vide infra).

6. CHALLENGES AND OPPORTUNITIES FOR FCC TECHNOLOGY IN THE 1990s

Tougher environmental regulations targeting automobile emissions will change the nature of the fuels produced in the 1990s as well as the corresponding refining processes. Important features of reformulated gasoline for the 1990s will include lower aromatics (especially benzene), lower olefins, lower gasoline vapor pressure, and mandated oxygen content (Table 9) [4]. Diesel fuels will have lower sulfur and aromatics content.

TABLE 9

U.S. Gasoline Composition (1 990) [4]

Estimated Possible Unleaded Reformulated Typical

Pool Specs. FCC Gasoline

Octane, (R+M)/2 RVP, psi. (summer) Aromatics, vol. % Benzene, wt. % Olefins, vol. % 90% Pt., O F Sulfur, ppm. Oxygen, wt. %

87-88 8-12

30-35 1-2

10-15 330-350 200-400 0.1-0.3

86-87 8 max. 6-9

25 max. * 20-30 1 m a . * .8

300 max. 380 5 max. 25-40

250-300 max. 500-700 2 min. *

* EPA mandated for ozone non-attainment areas. For CO non-attainment areas wt.% oxygen is the mandated minimum.

2.7

The reformulation of gasoline will cause significant changes in the FCC operation. Aromatics content of full-range gasoline will have to be reduced from about 30 to a maximum of 25 volume percent in reformulated gasoline. Benzene will have to be reduced to a maximum of 1 volume percent. Since FCC gasoline represents about one- third of the gasoline pool content, a reduction of FCC gasoline aromatics may be

178

required. This can be accomplished by reducing the gasoline end point, since such a reduction decreases the yield in heavy gasoline where the aromatics are concentrated. However, such a gasoline end point reduction decreases significantly the volume of FCC gasoline. A further reduction in aromatics can be achieved by distillation or extraction.

An alternate route to reducing the aromatics content is hydrotreating the FCC feedstock, especially highly aromatic feedstock. Cracking the hydrotreated feedstock will result in higher gasoline yields and lower sulfur content in liquid products. Aromatics in the gasoline pool can be further reduced by reduction of reformer severity. Benzene from gasoline can be removed by hydrogenation, hydrogenation and isomerization, or by extraction. It can also be reacted with light olefins from the FCC unit, to form alkylbenzenes [93]. Such an approach allows the production of reformulated-type gasoline and low-sulfur diesel fuel.

Newly emerging FCC catalyst technology may also contribute to the reduction of aromatics in gasoline. A recent report [30] describes a FCC catalyst containing a non- faujasite zeolite, capable of producing gasoline with a lower aromatic content than Y zeolite based catalysts.

The olefin content of FCC gasoline varies presently from 25 to 40 volume percent. The olefins are concentrated mainly in the light gasoline fraction (C,-C, range) and increase the gasoline octane rating. A possible reduction to 5 volume percent olefins will require significant changes in the FCC process in the 1990s. To reduce FCC gasoline olefinicity, the FCC unit may be run to be more selective to CJC, olefins, which can be used in the alkylation and MTBE (methyltertiarybutyl ether) or ETBE (ethyltertiarybutyl ether) units. Under these conditions, the FCC unit would become a major producer of light olefins. The higher CJC4 olefin selectivity can be achieved by using catalysts with HSY zeolites and active matrices, as well as by operating at higher conversions andlor reaction temperatured The use of ZSMJ additive will also enhance CJC, olefin selectivity. The reduction in gasoline yield that would occur under these conditions would be compensated by the higher alkylate yield. Alkylate is an attractive reformulated gasoline component, since it is low in olefins and aromatics, it has a low vapor pressure and has a high octane blending value.

The mandated minimum oxygen content of 2 weight percent will be supplied primarily by ethers, such as MTBE and ETBE. These ethers are valuable gasoline blending components due to their octane boosting capacity, their oxygen content, and their low vapor pressure. MTBE and ETBE are prepared by reacting isobutylene with methanol and ethanol, respectively. Therefore, isobutylene from the FCC unit will be a valuable product. Isobutylene yield and selectivity can be increased by raising reactor temperature, using ZSM-5 additives, and/or using catalysts with low hydrogen transfer capability (e.g., catalyst with HSY zeolites and active matrices). The effect of catalysts with different hydrogen transfer capabilities on isobutylene yields and selectivities is shown in Figures 20 and 21 [94].

Among the gasoline-range olefins, C, (especially isoamylenes) are potentially the most valuable. Their formation is enhanced by using HSY containing catalysts and/or ZSM-5, as well as by higher reactor temperatures. After separation, these olefins can be used in the production of alkylate, as well as in the production of TAME (tertiaryamylmethyl ether) and TAEE (tertiaryamylethyl ether). Both ethers are octane boosting and RVP reducing additives [95].

The reduction of gasoline RVP requires the reduction or elimination of C, and C, hydrocarbons from FCC gasoline. As already shown, the separated C, and C, hydrocarbons can be used in the production of alkylate and MTBE, ETBE or TAME.

179

v 1.6 l l b 1.4

.- 1.0 0.0

t? 1.2

30 40 50 60 70 80 Conversion

Figure 20 Isobutylene yields obtained with catalysts having different hydrogen transfer capabilities [94]

h

> -.

0.20

0.15 30 40 50 60 70 80

.-

Conversion

Figure 21 Isobutylene selectivity as a function of catalyst hydrogen transfer capability P41-

In conclusion, the 1990s will see significant changes in FCC technology as a result of new environmental regulations.

References

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4

5

6

C. J. Plank and E. J. Rosinski; Chern. Eng. Progr. Symp. Ser. 73 (63), 26 (1967). J. J. Blazek; Oil & Gas J., @ (459, 66 (1971). J. Scherzer; "Octane-enhancing zeolitic FCC catalysts," Marcel Dekker, Inc., New York, NY, 1990. G. M. Stokes, C. C. Wear, W. Suarez, G. W. Young; Oil & Gas J., July 2, 1990, p. 58. C. V. McDaniel and P. K. Maher; Conf. Molec. Sieves, 1967, SOC. Chern. Ind., London, 1968. J. Scherzer; J. Catal. 54- 285 (1978).

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70 71 72 73 74

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AM-88-73.

J.S. Magee and M.M. Mitchell, Jr. Fluid Catalytic Cracking: Science and Technology Studies in Surface Science and Catalysis, Val. 76 0 1993 Elsevier Science Publishers B.V. All rights reserved.

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CHAPTER 6

INSTRUMENTAL METHODS OF FCC CATALYST CHARACTERIZATION

ALAN W. PETERS

W. R. Grace & Co. - Conn. Columbia, MD 21044

1. INTRODUCTION

The FCC Catalyst and the FCC Process The FCC (Fluid Catalytic Cracking) catalyst is a coarse powder consisting of particles in

the 40 to 100 micron size range with an average size of about 65 microns. It contains 20% to 50% of an active zeolite, usually faujasite. with a surface area of 900 meters2/gram. In some cases the catalyst may also contain a moderately high surface area alumina or silica alumina matrix of -- 150 meters2/gram. The balance is a low surface area inert filler, typically clay, and a low surface area amorphous silica or silica alumina binder (1). The filler and binder provide the physical properties such as attrition resistance and density that are required for operation in a commercial fluid catalytic cracking unit. The properties and preparation of the catalyst are described in more detail elsewhere in this book.

Instrumental methods are used to characterize the FCC catalyst particle both fresh and after deactivation. The catalyst properties, such as zeolite content, the unit cell size of the zeolite, and the content and location of the active components determine the activity and selectivity of the catalyst. Catalyst stability is measured by comparing the fresh properties of the catalyst with the properties of the operating or equilibrium catalyst. In most advanced technology catalysts the zeolite is designed to hydrothermally dealuminate in a controlled and stable way to the intended unit cell size and surface area. The hydrothermal environment of the unit is in a sense part of the catalyst preparation. Also during use contaminant metals contained in the oil at the part-per-million level may deposit on the catalyst and cause changes in activity and selectivity. Control of these contaminants is one of the most important issues in catalytic cracking. It is important for the catalyst user, the manufacturer, and the researcher to be able to follow the chemical and physical changes that take place during the operation of the catalyst in the FCCU (FCC Unit). The operation of the FCCU is briefly described below. A more detailed discussion of the FCCU process is given by Venuto and Habib (2) and in other chapters of this book.

During operation the catalyst alternately passes between the reactor and regenerator, spending a few seconds in the reactor and about 15 minutes in the regenerator during each cycle. The reaction occurs when the hot regenerated catalyst is mixed with the relatively cooler oil at the bottom of the riser at a mix temperature of about 500°C - 550°C (930°F - 1020°F). The oil expands as it heats up and converts to lighter products. Both the oil and the catalyst move up the riser into a reaction vessel where the catalyst and hydrocarbon are separated. The products include gasoline, light gases, some unconverted oil, and coke embedded on the catalyst. The activity of the catalyst refers to the total amount of conversion to all light products including gasoline, lighter gases, and coke. The selectivity refers to the distribution of products. Desirable selectivities are for less coke, light gas, hydrogen, and

184

high boiling heavy slurry oil, for more gasoline with a higher octane, and more light cycle oil with a higher cetane index. Requirements for a more environmentally acceptable gasoline involve selectivities for less aromatics, less sulfur, more light olefins and less heavy olefins in the gasoline. Activity and selectivity are unfavorably affected by metal contaminants present in the oil, especially vanadium and nickel, and including iron, copper, and sodium. These metals form a residue on the catalyst along with the coke. Unlike coke, these residues cannot be removed during operation. Instrumental methods of analysis are used to determine the location and chemistry of these contaminants.

Before it is sent to the regenerator the catalyst is separated from the reaction products in the cyclones and passes through a stripping section where steam removes most of the remaining hydrocarbon from the catalyst. Stripping efficiency can depend to an extent on the pore structure of the catalyst. The catalyst discharges into the regenerator along with injected air. The coke on the catalyst is burned to CO, COz, water, and trace amounts of sulfur and nitrogen oxides. The heat of this reaction is enough to increase the temperature of the catalyst in the regenerator to the 720OC- 800°C (1320'F - 1450'F) range. About 20% of the regenerator off gas is water.

It is the presence of water at these temperatures in the regenerator in combination with the metal contaminants on the catalyst that cause loss of activity and changes in the selectivity of the catalyst. Both the zeolite and the matrix component can undergo significant changes, some desirable and some undesirable. The zeolite can dealuminate resulting in lower activity but improved (lower) selectivity for coke and improved selectivity for the production of higher octane olefinic gasoline. Both the matrix component and the zeolite can deactivate by losing crystallinity and/or surface area, resulting in generally lower activity and poorer selectivity. Also, both the matrix and the zeolite can undergo changes in the pore size distribution. The zeolite in a well designed catalyst will dealuminate to the desired unit cell size without loss of structure. It wjll form a system of mesopores interconnected to the zeolitic micropores.

The steam and temperature in the regenerator also has an effect on the vanadium on the catalyst. In this environment the vanadium destroys zeolite structure and surface area, reducing activity as well as causing changes in selectivity. The presence of nickel primarily causes increases in coke and hydrogen. The effects of nickel can be alleviated by adding antimony or bismuth compounds to the feed oil. A desirable catalyst will survive relatively high amounts of vanadium and nickel with acceptable losses in activity and selectivity. Currently vanadium levels of 7000 ppm and nickel levels of 4000 ppm are reasonable practical limits. As work continues in controlling the location and chemistry of these contaminants, the metal tolerance of catalysts will continue to increase.

Summary of the Use of Instrumental Methods of Analysis Instrumental methods are used to characterize the changes that occur in the catalyst during

the FCC process. These changes are then related to desirable or undesirable changes in the selectivity and activity of the catalyst. Consequently analytical information concerning both the fresh catalyst and the operating catalyst is important to the refinerhser as well as to the manufacturer and researcher. The various instrumental methods of analysis are divided into five somewhat arbitrary groups. The techniques within each group tend to provide similar or complementary information, make use of similar instrumentation, and involve similar sample preparation techniques. The five groups are

Diffractive Analysis (XRD, High Energy XRD, SAXS, Neutron Diffraction, EXAFS)

Diffractive analysis is used to identify the structure, composition and amounts of crystalline materials such as zeolite and aluminas. In the case of EXAFS the local arrangements of atoms can be identified in the absence of a long range crystalline structure.

185

Pore Structure Analysis, Adsorption (Nitrogen, Mercury, Water, He, Argon, and Organic Materials)

Adsorption methods are used to determine the pore structure, surface area, pore volume, and the distribution of small and larger pore structures in the catalyst.

Surface Analysis and Imaging (SEM,TEM, STEM, EDX, XPS (ESCA), Auger,SIMS, WDS (EPMA), STM, AFM )

Surface analysis provides information characterizing the morphology of the sample, provides element mapping and elemental associations, and can provide an estimate of the oxidation state of the elemental components.

Spectroscopic Analysis (lnfra red, Fl'IR, Raman, DRIFTS, Mossbauer, MASNMR, uv-visible)

Spectroscopic analysis is used to estimate the strength, number, and type of acid sites (Lewis or protonic), describe aluminum coordination and acidity, and provide descriptions of silicon and aluminum distributions.

(TPD, TGA, DSC, TPR, TPO, Microcalorimetry) Thermal Analyses

Thermal analysis is also used to estimate the number and strength of acid sites, more generally the strength of adsorption, thermal and hydrothermal stability, and stability to oxidationheduction.

2. DIFFRACTIVE ANALYSIS (XRD, High Energy XRD, SAXS, Neutron Diffraction, EXAFS)

Except for EXAFS and neutron diffraction, this group of techniques involves the diffraction of high energy radiation (x-rays) from the atomic lattice of crystalline compounds present in the sample. In the case of XAFS there is not necessarily a lattice, but there is a local symmetry or arrangement that gives a similar result., the presence of peaks at certain positions characteristic of the atomic distances and symmetry. In the case of neutron diffraction there is a lattice, but high energy neutrons are used rather than x-rays. These techniques are used to identify and to measure the relative amounts of crystalline materials in the catalyst. XRD techniques are commonly used to estimate the aluminum content in the zeolite by measuring the unit cell size.

Powder XRD (X-ray Diffraction) FCC catalysts may currently contain either or both of two different zeolites, faujasite and

ZSM-5. and may contain a crystalline alumina as a matrix component. Further, the catalyst may contain clay, and the clay can transform to spinel or mullite. The zeolite may also transform into mullite and crystobalite. All of these phases are crystalline and can be identified and quantified by XRD.

An explanation of the principle of the technique is illustrated in Figure 1. The crystalline compound forms a series of repeating planes with a spacing of a few tenths of a nanometer. The radiation is scattered by each atom in the plane. Part of the wave is reflected, but most goes through to the next plane where part is reflected, etc. The wave front interacts with the electrons associated with each atom on the plane. Each atom re-emits a spherical wavelet. The re-emmitted waves will be in phase and will give a peak at the detector for a particular angle, 26, such that the distances traveled by the wavelets differ by an integral wave length. At any other angles, even ones only slightly different, the amplitude of the wave reflected from neighboring atoms will be slightly out of phase, the wave from second neighbors will be

186

twice as much out of phase, and finally the negative amplitude of the reflected wave from some nth neighbor will cancel the reflection from the first and so on until there is a net cancellation. Consequently, the angle 26 is a measure of the distance between planes, and the intensity of the signal is a measure of the scattering power (number of electrons) and the density of the atoms in the diffracting plane. For a given material, the intensity is a measure of how much material there is and of how perfectly the planes are ordered.

Figure 1. Pictorial representation of X-ray scattering intensity reinforcements responsible for the distinctive X-ray patterns of crystalline materials. After H. P. Mug and L. E. Alexander, X-Ray Diffraction Procedures, John Wiley & Sons, 1974, p. 121.

Identification of c The basis of he XRD technique is that crystalline materials have peaks at values of 26

such that both the values of 26 and the intensities are characteristic of the structure of the material. Figure 2 shows the powder patterns of two zeolites, faujasite and ZSM-5, used in cracking catalysts as well as a pattern characteristic of a clay (kaolin) often used as a filler. Crystalline or amorphous aluminas may also form a part of the matrix of the cracking catalyst. The XRD scan of a cracking catalyst may show peaks due to all or several of these components. Amorphous materials, of course, do not have an X-ray pattern and cannot be identified by XRD.

A major feature of current XRD systems is the existence of search routines capable of identifying zeolites, clays, or other catalytic components from an XRD scan. The JCPDS (Joint Commission for Powder Diffraction Standards) files include XRD patterns for about 20,000 inorganic materials available in a computer searchable format (6) and include materials that may occur in FCC catalysts such as zeolites, clay, and aluminas. Although currently only two zeolites, faujasite and ZSM-5, are being used commercially, other zeolites will almost certainly be used in the future. Several collections of zeolite XRD scans are available in the literature (7-9). A collection of XRD scans for aluminas is published by Alcoa (lo), and collections of scans for various clays have also been published (11). An improved search procedure using the full XRD scan rather than just a few major peaks has been described (12) and is commercially available through the authors from Pennsylvania State University along with an extensive and current zeolite data base. Nearly 300 programs for the analysis of powder diffraction data have been recently reviewed (13). Many of these

187

140-

120-

100

- 80-

3

: sc:

40

2c

programs are included in the Penn State package. Using these and other files it is possible to obtain an identification of any crystalline material provided its XRD pattern is available in a computer or hand searchable form.

-

-

-

'"ooo~i 14000 I 2a. Dealurninated USY

Degrees. 2-Theta

t

0 15

2b. ZSM-5

u 35 45

Degrees. 2-Theta

lC000

, o o o ~ " " ' I' " " ' " " ' " " " " " " " '-7

Figure 2. X-Ray patterns for a) dealuminated USY,

c) Clay (Kaolin). b) ZSMJ,

Degpees. 2-Theta

188

XRD intensities can be used to estimate the amount of a material present in a catalyst. The XRD intensities are compared to some standard defined as 100% crystallinity. There is a procedure defined as a standard test, ASTM D 3906, for the measurement of crystallinities of faujasite containing material (14). This technique is often used to measure the relative amounts of zeolite such as faujasite in either a fresh catalyst or in commercial or lab deactivated catalysts. The stability of the zeolite is the per cent zeolite retention. The presence of varying amounts of exchange cations such as sodium or rare earth can significantly alter intensity relationships. Further, the recently observed presence of an extensive system of internal surface defect structures (mesopores) in the framework of hydrothermally dealuminated zeolites is expected to cause an intensity loss. Atoms near the internal surface will be slightly displaced from their ideal lattice positions and so there will be some interference of amplitudes. In zeolites with a high mesopore surface area of -100m2/g, 10% or more of the atoms may be at a surface. These atoms may not be adequately counted even though they are a part of the zeolite structure. For these reasons quantitative analysis by XRD can be inaccurate and should only be used to compare samples of similar types of zeolite.

Qystallite size Verv small crvstallites will have fewer diffracting Dlanes and so the angle over which

amplitldes add Gill be larger. This is known as l i k broadening, and can' be used as a measure of crystallite size in the 0.01 p to 1 p size range (15) . Since line broadening is associated with a loss in peak height, intensity measurements for analytical purposes should be obtained from peak areas. Realistic estimates of crystallite size require that the optics of the instrument be adequate to eliminate instrumental broadening. Usually this involves the use of a primary monochromater as discussed in more detail in the section concerning high resolution XRD. Since most samples of zeolite are agglomerates with a rock pile morphology, crystallite size is not the same as particle size. Particle size and crystallite size are the same only if one has a collection of single crystals without intergrowths.

un i t Cell Size Measurement i n F a u i w In a zeolite framework consisting of silicon and aluminum, the aluminum is the

catalytically active ingredient. The activity, selectivity, and stability of the zeolite are all related to the framework aluminum content of the zeolite. Chemical analysis will give the aluminum content of the framework in a pure zeolite sample if there is no other source of alumina. However, the aluminum removed from the framework by dealumination, either during manufacture or during deactivation, remains within the structure in some form as nonframework alumina. Further the catalyst will contain other sources of alumina including clay and matrix components. It is desirable to be able to measure the amount of framework alumina in a zeolite independently of the occurrence of other forms of alumina in the catalyst or the zeolite. A number of instrumental methods of measuring framework aluminum content have been developed. It is possible to estimate the aluminum content of the framework by measuring the frequency shift of the 800 cm-' and 1050 cm-I i. r. absorption bands (16, 17). and also by 29Si MASNMR (1 8) as well as by the XRD methods discussed below. These other procedures are discussed in the section on spectroscopic analysis.

The most common method of measuring the aluminum content of the framework and the only one directly applicable to the catalyst is based on the measurement of the unit cell by XRD. Since the Si-0-A1 bond is longer than the Si-0-Si bond, the unit cell increases slightly with aluminum content. Consequently the unit cell size is a very important parameter. The measurement of unit cell size by XRD has been standardized as ASTM Test D-3942 (14). In the case of faujasite there are several correlations of unit cell size and alumina content in current use, one developed by Breck and Flannigan (19), and another more recently by Fichtner-Schmittler (20) and by Sohn (17). The correlations of Breck and of Sohn are shown

. .

189

graphically in Figure 3. The earlier correlation of Breck and Flannigan was developed for low SUAl ratio (1.5-2.5) as-synthesized samples of sodium exchanged faujasite using direct chemical analysis as the reference, while the two latter correlations were developed from dealuminated and decationized samples of faujasite using 29Si MASNMR results as the Si/Al reference.

Using these correlations there are simple but useful relationships between the unit cell size determined by XRD, the silicodaluminum atom ratio of the framework, R, and the number of aluminum atoms per unit cell. Provided there is no nonframework alumina present, it is possible also to calculate the number of aluminum atoms per unit cell from the chemical analysis, below, and to compare the results with an estimation of the aluminum content in the framework, Table 1.

By Chemical Analysis R=Si/AI = (%SiO2~51)/(%Al203~60) Si02/Al203 = 2R # AYunit cell = 19U(R+1)

Table 1. The measurement of aluminum content per unit cell in faujasite by XRD analysis using the published unit cell correlations, where a is the unit cell in nanometers.

# AVunit cell = rn

1152(a - 2.4191) Es

1124(a - 2.4233) sahn

1071(a - 2.4238)

There is no single c o m t correlation since cation exchange and even the degree of hydration can significantly affect the unit cell size. As-synthesized sodium faujasites will have a unit cell of about 2.460 nm to 2.475 nm. After stabilization and dealumination the unit cell will be 2.440 to 2.460 nm. and after dealumination in the FCCU the unit cell will typically be 2.422 to 2.432 nm. The unit cell size of an equilibrium catalyst can be lower than the zero aluminum limit of the Sohn and the F-S correlations. A unit cell size of 2.417 nm, lower even than the minimum predicted by the Breck correlation, has been reported in the literature (21). One reason for lower than expected unit cell sizes is that as the catalyst dealuminates, some of the aluminum occupies exchange sites (22,23). Cation exchange can reduce the unit cell size by as much as 0.004 nm compared to the same decationated zeolite. Removal of cations such as sodium from as-synthesized faujasite or nonframework alumina from steam dealuminated faujasite results in an increase in unit cell by 0.002 to 0.004 nm. The older Breck relationship may be more useful for deactivated FCC catalysts, while the more recent correlations may be more appropriate for experimental or otherwise decationated zeolites.

For other zeolites having a higher as-synthesized silicon to aluminum ratio this kind of relationship has not so far been useful. Zeolites such as ZSM-5 contain small amounts of aluminum, and small variations in aluminum content do not produce enough of a change in the unit cell size to be easily and quantitatively measurable.

Resolution XRD A limiting factor in the accuracy of the unit cell size measurement is the resolution of the

XRD unit. The X-ray source, the Cu K alpha emission line, is a doublet, and is close to the beta emission line. Most commercial X-ray units have a secondary filter placed after the sample that eliminates the beta line, but not the K alpha doublet. A primary filter will eliminate the doublet but requires a higher intensity source to compensate for the intensity

190

loss associated with the primary filter. The use of an intense monochromatic beam available from a synchrotron source at one of the storage ring facilities, for example The National Synchrotron Light Source at Brookhaven National Laboratory, allows an order of magnitude improvement in the accuracy of the unit cell size measurement (24). A monochromatic and intense X-ray beam is useful in other contexts as well. It is possible to detect small amounts of crystalline impurity phases or to deduce crystal structures from high resolution powder data or from micrometer size single crystal data (25,26).

80

- Breck, Ref. 19 - - - Sohn, Ref. 17

60 ._...______________._ 1 40

20

0 2.42 2.43 2.44 2.45 2.46 2.47 2.48

Unit Cell, nm

Figure 3. Correlations between unit cell size in nanometers (10 a = 1 nm) and aluminum content, aluminum atoms per unit cell, one unit cell contains 192 total atoms, excluding oxygen. From references 17 and 19.

SAXS (Small Angle X-ray Scatterinel Small angle XRD scattering has been used in the past to evaluate the distribution of

relatively large structures such as particle size distributions or pore size distributions. In the case of pore size distribution the pores are fiiled with liquid containing a heavy atom. Since scattering power increases with the number of electrons per volume, a heavy atom will contribute a large scattering cross section (27). Recent experiments with light scattering have shown that the fractal dimensionality of a particle can be determined by scattering experiments (28). The fractal dimensionality is related to the particle shape.

Neutron Diffraction The physics of neutron diffraction is very similar to XRD, Figure 1. The neutron beam is

obtained from a nuclear reactor. The flux or number of particles per unit of time is lower than for the X-ray beam, so the analysis of complex structures can be difficult. Unlike the X-ray beam the neutron beam is strongly scattered by certain light elements including deuterium. It is possible to obtain structural details using neutron diffraction not normally obtainable by ordinary XRD methods. A recent study used neutron diffraction to locate the protons in the

191

faujasite structure (29). Neutron diffraction experiments have also been used to show that the aluminum distribution is disordered in faujasite (30). A review of neutron scattering experiments has recently appeared in Science (31).

EXAFS (Extended X-ray Absorption Fine Structure Spectroscopy) XANES (X-ray Absorption Near Edge Spectroscopy)

These techniques use sychrotron radiation as an X-ray source, available at some of the national laboratories such as Brookhaven. At the energy required to remove a core electron from an atom there is a sharp increase in X-ray absorption, called the absorption edge. Near the absorption edge there are intensity peaks and on the edge there are oscillations that result from an interference of amplitudes between the ejected electron and the electrons associated with the local surrounding atomic distribution, Figure 4a. The frequency spectrum of the oscillations can be Fourier analyzed to give interatomic distance information shown as a peak or series of peaks, Figure 4b. This technique is primarily applicable to catalysts containing deposited or impregnated metals where one is interested in the local coordination or bond distances around the catalytic metal atom.

The technique has been applied to an analysis of vanadium chemistry on FCC catalysts. The results show that vanadium is not present as VzOs on the catalyst after calcination or steaming, but instead is present as a highly dispersed oxide species such as VOd3- (32). XANES results have further shown that vanadium interacts very strongly with proposed passivators such as magnesium (33).

i t I f 1 I I 4a Vanadium edee X-Ray clrlll

4

1 Absoiption &fIi

:Ld 0 5 10 15 20 25

- Y

'E

4h. Vanadium Radial Structure Function

0 2 4 6 8

Energy leVl X10 Radius. i

Figure 4. a) X-Ray absorption coefficient at the vanadium edge for a vanadium impregnated FCC catalyst as a function of energy. b) The waves on the top of the edge are Fourier analyzed to give the Radial structure function as a function of distance from the vanadium atom. The peaks occur at distances where coordinating atoms occur, in this case oxygen, and are characteristic of the particular compound V04, D. J. Sajkowski, S. A. Roth, L. E. Iton, B. L. Meyers, C. L. Marshall, T. H. Fleisch, and W. N. Delgass, Appl. Catal., 1989,51,255.

3. PORE STRUCTURE ANALYSIS BY ADSORPTION AND ABSORPTION (Nitrogen, Mercury, Water, He, Argon, other Gases and Organic Materials)

Adsorption methods are used to characterize the pore structure of the catalyst surface and also the thermodynamics of the interaction between the surface and the adsorbent. Besides a determination of the surface area, pore volume, and an indication of the distribution of pore sizes, it is possible to use adsorption methods to estimate the amount of zeolite in either the fresh catalyst or the deactivated catalyst. Another application, referred to as pore gauging,

192

permits an estimation of the size of the pore opening in experimental zeolites. This method involves adsorption experiments with organic molecules . Gregg and Sing (34) have given a comprehensive review of the determination of the surface area and pore size distributions in catalysts as have Lowell and Shields (35). The Proceedings of recent IUPAC Symposia provides an update on many topics of current interest including those discussed here (36.37).

Nitrogen Nitrogen adsorption is used to determine the total surface area of a catalyst, and can also

be used to estimate the small pore zeolitic component and a larger pore component. A procedure for the determination of the total surface area using the BET (Brunauer, Emmett, Teller) method is described as ASTM test D 3663 (14). The sample is calcined to remove moisture, exposed to a measured volume of nitrogen at some constant pressure P, and the volume adsorbed by the catalyst is measured. An adsorption curve, referred to as the adsorption isotherm, is generated by starting at some low pressure and measuring the volume adsorbed. The pressure is increased slightly, and the new adsorbed volume is measured. The adsorption data can provide an estimation of the relative amounts of small zeolitic pores and larger pores and is widely used in the analysis of FCC catalysts to separate the zeolite and matrix contributions to the surface area (38). This method is referred to as the t-plot method and has been developed as an ASTM standard test, D 4365 (14).

The basis for these methods is the nitrogen isotherm. Microporous systems have what is called a type I isotherm. Catalysts containing both zeolitic micropores as well as mesopores have a combination of a type I isotherm with a type IV isotherm, Figure 5a. Adsorption in the small pores of the zeolite mostly occurs at a low pressure p relative to atmospheric po, p/po < 0.1, while adsorption by the larger pores of the matrix or the mesopore system occurs in the region p/po > 0.1. The adsorption data can be recalculated and plotted in such a way that the amount of nitrogen adsorbed is plotted against the thickness t of the adsorbed layer. For thin layers, both micropores and mesopores contribute, but for thicker layers, the micropores are filled up and only the mesopores contribute. Figure 5b shows the conversion of a conventional plot of the amount adsorbed with partial pressure to a t-plot of the amount adsorbed with the thickness of the layer.

The t-plot method is frequently used to estimate the amount of zeolite present in a cracking catalyst. In the case of a faujasite dealuminated to a low unit cell size the method can be misleading since the zeolite can form a varying degree of mesoporosity during dealumination depending on the preparation and method of dealumination, e.g. with SiF,' or with steam (39). Other work has demonstrated that catalytic activity, especially for bottoms cracking, is associated with the development of mesoporosity (40). In the case of catalysts containing significant mesoporosity the t-plot method will count the mesoporosity as matrix surface area, when in fact it is zeolitic. As much as 20% and as little as 5% of the zeolite surface area at low unit cell size can appear as mesopores. The t-plot method has also been used to obtain an estimate of the particle size of as-synthesized zeolites (41). Since there is no mesoporosity in this case, the mesopore surface area measures the external surface area of the zeolite, typically 2 to 20 m2/g, depending on the particle size. Differences in particle sizes can be independently if only qualitatively confirmed by SEM pictures. Zeolite particle size can have important selectivity and stability consequences (41).

The nitrogen desorption curve is sometimes used to determine a pore size distribution. Although this procedure has also been standardized as as ASTM test D 4641 (14). the results are not entirely reliable. It has been noted that 4.0 nm diameter pores are frequently observed as peaks in the desorption curve of a variety of materials. This is due to hysteresis in the adsorptiorddesorption isotherm at a P/PO of - 0.4. probably due to the surface tension of the nitrogen film, Figure 5a. This discontinuity in the isotherm is reflected in the desorption pore size distribution (42).

The nitrogen adsorption methods tends to be unreliable for very large pores over 60 nm in diameter where adsorption occurs at pressures near atmospheric. In cases of very low surface

193

area catalysts krypton adsorption is used, ASTM standard D 4780. This test is not normally applicable to the typically high surface area FCC catalysts.

a) Isotherm / /

/ 140

4 0 F " " " " ' 1 " " " " ' L " " ~ " " " " " " " 1 ' " " 0 . 0 0 . 2 0 . 4 0 . 6 0 . a

PIP0

0 5 10 15 20

t function

Figure 5. a) The desorption isotherm of a typical cracking catalyst containing a region characteristic of microporous zeolites, and a region, p/po > 0.1, dominated by mesoporosity. The isotherm is of type IV, characteristic of a sample such as a cracking catalyst containing both mesopores and micropores. after S. J. Gregg and K. S. W. Sing, Adsorption, Surface area, and Porosity, Academic Press, New York, 1982. p. 4. Also shown as marked is an example of nitrogen adsorption/desorption hysteresis showing the ubiquitous anomaly at p/po of about 0.4 on the desorption branch due to the surface tension of the nitrogen film. The anomaly can produce artifacts in pore size distributions measured from desorption data, S. J. Gregg and K. S. W. Sing, Ibid., p.161. b) A t-plot of the isotherm in Figure 5a. The thickness of the layer of nitrogen being adsorbed is plotted against the amount adsorbed. The slope at a given p/po (or thickness) is the amount adsorbed per layer and is a measure of the surface area at that point. S. J. Gregg and K. S. W. Sing, Ibid., p. 97.

Mercury Mercury intrusion is especially useful for determining the volume and surface area of

larger pores, typically over 20 nm to whatever size is appropriate. Powders will give spurious intrusion peaks in the 10oO nm range caused simply by the spaces between particles. The size and shape of the particles in the sample will set the upper limit of usefulness of this method. For this test also there is an ASTM standard, D 4284 (14). An analysis of the hysteresis can

194

supply additional important information about the shape of the pores, ink bottles, cylinders, etc. (43). In using the mercury method, an important parameter used in the calculation of the pore size distribution is the wetting angle between the mercury and the pore wall. In practice, the wetting angle is set to a value that gives good agreement with the nitrogen pore size distribution over a range of 20 nm to 60 nm where both techniques are reliable. For silicdalumina an estimate of the wetting angle is about 140'.

Water The incipient wetness method measures the total pore volume by measuring the amount of

water absorbed in the pores. Water is added to the sample until the particles just begin to stick together and form a cake.

Helium Helium pycnometry is a helium adsorption technique for determining the skeletal density

of porous solids (44). The helium atom is sufficiently small that it will enter into the smallest pores. The result is a measurement of the skeletal volume. If the weight of the sample is known then the skeletal density can be calculated. In the case of zeolites this result gives the framework density. Furthermore. from the skeletal density and the surface area it is possible to estimate the average pore size. This is a result frequently used in reaction engineering estimates of diffusivities (45).

Low Pressure Adsorption Methods involving the use of argon at very low pressures have recently become available.

These methods are useful in the estimation of pore size distributions in the small pore range of 2 nm or less pore diameter. Zeolites, for example, have pore sizes in the 0.5 to 1.5 nm range and some zeolites may have several types and sizes of pores. In these cases low pressure isotherms may be used to characterize zeolites or mixtures of zeolites. For example, using this technique M. Davis was able to show that the molecular sieve MCM-9 is a mixture of VPI-5 and SAPO-11 (46). Other catalytic materials such as pillared clays and carbon can also have distributions of very small pores. The application of these techniques to carbon has been recently described by Carrott and Sing (47).

Organic compounds Adsorption isotherms of organic compounds have been used to characterize the pore size

of zeolites. A discussion of the use of different size molecules to characterize the dimetisions of the zeolite pores has been given by Szostak (48). The pores of zeolites are formed by circumscribed rings of -0-T-0-T- where T is either silicon or aluminum. In a small pore zeolite the rings contain eight T atoms, in a medium pore zeolite such as ZSM-5 the rings contain ten T atoms, and in a large pore zeolite such as faujasite or mordenite the rings contain twelve T atoms. The method is based on the idea, for example, that a small pore zeolite with eight membered rings and a pore diameter of about 0.4 nm will absorb n-hexane, but not a bulkier cyclohexane. Table 2 gives the approximate kinetic diameter of selected molecules used as adsorbents as well as approximate pore sizes of small, medium, and large pore zeolites. Recently M. Davis prepared a new zeolite, VPI-5, containing 18 membered rings with 1.2 nm pores. This material will absorb triisopropylbenzene with an estimated kinetic diameter of 0.85 nm (49).

Organic compounds can also be used to measure surface areas, and isotherms generated using organic compounds can have catalytic significance. The original observation of the existence of a mesopore system in a dealuminated faujasite was based on mercury intrusion results and n-hexane isotherms (5031). Subsequent work with a wide range of organic absorbents showed that part of the pore volume of dealuminated faujasite is in the mesopore range. In this work the thermodynamics of adsorption of a wide range of organic materials on variously dealuminated zeolites was determined. General relationships were developed that will allow the development of (P,T) absorption isotherms from a limited amount of data (52).

195

Table 2. a series of different sized molecules at P P o - 0.4. Ring size 8-Ring 10-Ring 12-ring Pore size small medium large

Pore gauging. Molecular sieve pore size estimation from adsorption properties of

-Pore diameter (nm) 0.40 0.53-0.6 0.6-0.75

Adsorbate s i z d m l Does it dssxb? H 2 0 0.26 Yes Yes Yes n-hexane 0.4 Yes Yes Yes cyclohexane 0.6 no Yes Yes n e o pe n tan e 0.62 no no Yes

4. SURFACE ANALYSIS AND IMAGING (SEM,TEM, STEM, EDX, XPS (ESCA), Auger, SIMS, WDS (EPMA), STM, AFM )

The electron microscope with attachments for elemental analysis (EDS. WDS) is the most commonly used instrument for surface analysis. This technique involves bombarding the surface with an electron beam and detecting the backscattered electrons. This technique can be used to observe the surface topology of the catalyst including the occurrence and location of zeolite and matrix materials embedded in the catalyst particle. Since the beam penetrates. the microprobe provides an analysis of a small volume of the sample, - 1 pm3. Consequently the location and distribution of contaminant metals and other elements is determined in a surface layer with a thickness of about 1 micron. In the case of the transmission electron microscope (EM) the electrons are transmitted through a thin section of the sample. This technique has been used to image zeolite lattice structures and can be used to identify very low zeolite levels in a catalyst particle.

The other techniques use either an X-ray beam (XPS), an ionic or an atom beam (SIMS, FAB-MS) or do not use a beam (STM, AFM) to bombard the sample. The idea is to bombard the surface and to observe a reflected or an emitted particle. In this case the observed surface is only a few atoms thick. The energy that is lost or emitted will be characteristic of the type of atoms on the surface and will depend on the oxidation state of the atoms. XPS and Auger spectroscopy provide information concerning the location and sometimes the oxidation state of metals on the catalyst particle. These and other methods have been extensively reviewed by Briggs and Seah (53) and more recently by Fiem (54).

In using most of the surface techniques the sample is typically embedded in a matrix, often epoxy, and is then cut and polished in such a way as to expose a section through the sample. Elemental migration from one part of the sample to another or between sample and matrix due to high pressures or to abrasion and high temperatures during polishing can be a problem unless special care is taken. The sample may also be cut in a thin section (microtomed) for transmission electron microscope studies. What is exposed and analyzed may be the surface of the particle or may be the interior of the catalyst particle exposed by the sectioning procedure. Frequently argon etching of the exposed surface is used to expose a fresh layer of catalyst. This procedure is referred to as depth profiling. Surface techniques are therefore not restricted to the surface, but rather refer to the fact that an analysis can be obtained for a small and very precise region of the sample rather than for the bulk.

If an electron beam is used or if electrons are emitted, the development of an electric charge by the sample is an important consideration. Excessive charging of the sample lowers resolution and can divert the electron beam. Nonconducting samples such as FCC catalysts are coated with a conductive material such as carbon or a metal in an attempt to prevent the

196

buildup of charge during bombardment with the electron beam. Frequently charge will build up no matter what, so electron bombardment methods can be difficult with nonconducting samples. '

In some cases the sample can acquire a positive charge. The SIMS technique uses a positive ion beam, and in the case of XPS the ejected photoelectron leaves a positive hole. A solution in this case is to flood the sample with electrons using the electron flood gun to neutralize the charge build up as it occurs. The ability to neutralize positive charge build-up on non conducting materials like FCC catalysts is an advantage in applying XPS and SIMS.

THE ELECTRON MICROSCOPE AND MICROPROBE - High Resolution Catalyst Topology and Elemental analysis

Electron microscope (SEM, STEM, TEM) There is a group of methods used for surface analysis that involve the bombardment of the

surface with an electron. The type of information obtained and the resolution or the spot size depend on the energy and resolution of the beam and on the detection system, and the depth of penetration depends on the energy. The electron microscope is typically used to observe the structure and topography of catalytic materials at high resolution. The images occur as light and dark regions. Dark regions can be holes or may indicate the presence of lighter elements, while light regions may represent a clump of heavier elements. Under the best conditions the electron microscope (STEM or TEM) can resolve features well below one nanometer in size. The electron microscope can be used to observe the pore openings of zeolites such as faujasite, beta, and ZSM-5 as well as other zeolites.

A commonly used instrument is the SEM, the Scanning Electron Microscope. It typically operates at a magnification of 20,O)lO or more and can resolve structures of 0.1 p or less. Since FCC particles are about 70 pm in size and are composed of O.1pm to lpm size particles of zeolite, clay, and matrix materials, the SEM can be used to observe the distribution of the zeolite, clay and matrix particles in the catalyst particle where the clay and zeolite particles can sometimes be identified by their shape (1). The SEM is also commonly used to identify the morphology of catalytic components such as zeolites. Zeolites can occur in a variety of sizes and shapes. These parameters are known to have effects on selectivity, for example as observed in the case on ZSM-5 (55). In preparing zeolites frequently small amounts of another phase may occur as an impurity. The SEM may be used to identify impurities by differences in morphology compared to the main component.

The TEM (Transmission Electron Microscope) and the Scanning Transmission Electron Microscope (STEM) operate at a higher magnification, about 100,OOO or more, and can resolve features as small as 0.2 nm, about the size of an atom. These machines are frequently used to observe structures at the molecular level in both zeolites and in catalysts. TEM micrographs are used to observe the existence of intergrowths and faulting in pure zeolite samples. Intergrowths involving several structurally similar materials are commonly observed in ZSM-5 crystals and more recently in beta zeolite. The occurrence of significant mesoporosity (1 1) has been subsequently observed directly in TEM micrographs (56-58). Figure 6. The sensitivity of TEM is illustrated by recent work by Beyerlein (59) in carrying out an analysis of a FCCU deactivated catalyst. He was able to separate various fractions of the catalyst by age using a sinWfloat technique. He found that the oldest fractions having spent the longest time in the unit were the most severely deactivated. These fractions did not contain zeolite by XRD, but small zeolite fragments were clearly visible by TEM. The TEM or STEM can also be used to observe the growth and size of metal crystallites. Activity is proportional to the dispersion of the metal particles which is assumed to be inversely related to the size of the crystallites. The TEM has been used to observe nickel crystallites on metal contaminated FCC catalysts (60).

197

Figure 6. A TEM micrograph showing both the zeolite lattice and larger mesoporous holes resulting from hydrothermal dealumination, F. Mauge, A. Auroux, J. C. Courcelle, Ph. Engelhard, P. Gallezot, J. Grosmangin, Studies in Surface Science and Catalysis, Vol. 20, Catalysis by Acids and Bases, Elsevier, New York, 1985, p. 94.

Both the STEM and the SEM are frequently equipped with elemental analysis capability, EDS or Microprobe. High resolution elemental analysis capability has permitted the use of the STEM to observe, for example, the aluminum profiles across a faujasite particle resulting from the various dealumination techniques. Low temperature solution phase dealumination by diammonium silicon hexafluoride creates an aluminum gradient such that the framework of the outside of the zeolite is silicon rich with additional silicon deposits occurring on the zeolite (61). In this case the external surface of the zeolite is more strongly dealuminated than the interior. As discussed in the appropriate section, XPS has been used to show that during hydrothermal dealumination the aluminum from the interior migrates to the external surface.

Elemental Analysis (Electron Microprobe (EPMA, WDS), EDS, XES) It is often desirable to be able to analyze the chemical composition of the structures

observed using the electron microscope. For example, one may want to know the locations of the zeolitic and matrix components, or the location of various contaminants or poisons such as nickel, vanadium, and iron. For this purpose the SEM often has an attachment capable of providing elemental analyses of a portion of the surface. The bombarding electron beam produces a characteristic X-ray emission spectrum from the inner electronic shells of the elements in the sample. The technique of measuring the X-ray emission spectrum of a sample

198

is called XES, or X-ray Emission Spectroscopy. The instrument that measures the the x-ray emission spectrum using electron bombardment is called an electron microprobe, and the technique of analysis using the electron microprobe is called EPMA (electron probe microanalysis). There are two different detector systems in use for measuring the X-ray spectrum, EDS or WDS. The EDS or energy dispersive spectrometer attachment typically has poor energy resolution so that peaks of some elements will overlap with others. If one is interested in the elemental composition of some region of the sample, the electron beam will focus on that spot, typically about 1 p in size and 1 p deep. X-ray emission intensities will be acquired in enough channels to provide the qualitative amounts of a relatively large number of elements over a few minutes.

The WDS (wavelength dispersive spectrometer) has much better energy resolution and can resolve each element in the presence of others. Typically the WDS is used in a scanning mode where an FCC particle of 50-80 p is scanned and the profile of specific elements are obtained. For example, this technique has been used to identify the location of zeolite (regions of high silicon) and to associate the poison vanadium with various catalytic components such as rare earth (62) and alumina (63). The observation that in the presence of steam vanadium migrates both within a particle and between particles was based in part on WDS results (63). A draw back of the WDS and EDS methods is that sensitivities are different for different elements and therefore a calibration procedure is required.

XPS (X-ray Photoelectron Spectroscopy) ESCA (Electron Spectroscopy for Chemical Analysis) UPS (Ultra violet Photoelectron Spectroscopy)

XPS, sometimes identified as ESCA, involves bombarding a sample with X-ray radiation and measuring the energies of the emitted photoelectron. These energies are closely related to the binding energy of the core electrons in the particular compound. Each element has a characteristic spectrum of photoelectron emission energies. In a compound or in a solid, these binding energies are modified by the effective charge on the atom and to an extent by the surrounding atoms or ions. A positive charge will shift the binding energies higher, while a negative charge will produce a shift to lower binding energy. The XPS spectrum allows one to identify the major elements in a sample, and more importantly, from the position of the peak to estimate the oxidation state of the element.

There are two different kinds of XPS instruments. XPS with a high spatial resolution may include a high intensity synchrotron radiation source. More commonly, the XPS instrument is used to observe the surface without spatial resolution and to provide information concerning the chemistry of the observed elements. XPS spectra has been used to identify the oxidation state of vanadium (+5) and nickel (+2,+3) on cracking catalysts (64).

Recent XPS studies have experimentally confirmed that the acidity of zeolites is a property of the lattice associated with oxygen polarizability rather than a local phenomenon (65,66) and support previously proposed relationships between acidity, lattice electronegativity, and aluminum content (67,68).

Steam dealumination produces a uniform framework dealumination profile in that an .equal amount of aluminum is removed from the surface and the interior of the zeolite particle. The XPS results further show that the alumina formed as a result of dealumination migrates to the outside of the zeolite during the steaming and does not remain in the pores (69-70). XPS has also shown that low temperature solution phase dealumination gives a silicon rich zeolite surface (71). as previously discussed (61). Taken together, these results have considerably clarified the events that occur both during catalyst preparation and during zeolite hydrothermal dealumination and deactivation in a FCC regenerator.

UPS is similar to XPS except the lower energy of the photon in the ultraviolet region produces photoelectrons from the valence shells. Consequently the results are more sensitive to information concerning the details of the chemical bonding. UPS is not commonly used in the analysis of FCC catalysts.

199

Auger Electron Spectroscopy (AES) Auger Spectroscopy involves bombardment with electrons of a sufficiently high energy

to ionize core electrons and an observation of the consequent relaxation process, the process responsible for removing the energy involved in creating the ionized state. Part of the energy is removed by a transition of an outer shell electron to the lower energy core shell. The excess energy is lost either by the emission of an electron, the Auger electron, or by the emission of a photon. The standard X-ray fluorescence technique detects the photon. The Auger spectrometer analyzes the energy of the electron, and based on the analysis can identify the element and, sometimes, the oxidation state.

Cracking catalysts are nonconducting, so the electron beam from the Auger spectrometer produces significant charging of the sample. Since Auger and XPS can provide similar information, XPS is more commonly used. Auger in combination with Argon etching has been used to show that at 725 OC nickel antimony alloys even with a low antimony content are enriched in antimony at the surface (72). This result provides some insight into the mechanism for antimony passivation of nickel on FCC catalysts.

Figure 7. Cross sections of equilibrium catalysts containing zeolite, clay and alumina particles analyzed by the SIMS technique. An elemental mapping of b) silicon, c) lanthanum, d) vanadium, and e) nickel show that both nickel and vanadium appear on the surface of the particles, nickel more so than vanadium, from D. P. Leta, W. A. Lambetti, M. M. Disko, E. L. Kugler, and W. A. Varady, Fluid Catalytic Cracking II, M. L. Occelli, ed., ACS Symposium Series 452, American Chemical Society, Washington D. C., 1991. p. 276.

200

SIMS (Secondary Ion Mass Spectroscopy) SIMS is an ion sputtering technique. A beam of positive ions (eg. argon, oxygen, cesium)

ionize the surface and knock surface ions free. These ions are collected and analyzed in a mass spectrometer. The mass spectrometer gives the atomic weight of each ion observed. Since the detector normally used is sensitive to single ion events, the method provides a complete and a very sensitive elemental identification of all elements present. In a scanning mode it can provide elemental maps of the catalyst surface with 0 . 1 ~ spatial resolution. Since some elements sputter much more easily than others, one drawback of the technique is that the surface composition can change during the analysis. For those elements that sputter most easily the technique can be very sensitive, orders of magnitude more sensitive than the electron microprobe.

Recent SIMS results on equilibrium FCC catalysts show the association of vanadium with alumina, with zeolite, and with rare earth (73) in a commercially deactivated equilibrium catalyst. Nickel is associated with alumina in the catalyst particle. Both vanadium and nickel appear to be deposited on the outside of the catalyst, Figure 7. Vanadium is more mobile and penetrates more easily (74,75). A study of laboratory impregnated and deactivated catalysts using SIMS gave similar but not exactly the same results (76). Both nickel and vanadium concentrated to an extent on the surface. This work is especially important in attempting to understand and to simulate catalyst poisoning by nickel and vanadium in the laboratory. The metals initially lay down on the outside of the particle and migrate to the interior. Nickel migrates much more slowly than vanadium, and both react strongly with aluminum.

FAB-MS (Fast Atom Bombardment Mass Spectroscopy) In this technique a beam of energetic atoms knocks ions from the sample. This technique

has been used to make the first observation of the enrichment of aluminum at the surface of hydrothermally dealuminated zeolites (77, 78). Since the sample is bombarded with neutral species, sample charging is less of a problem.

STM (Scanning Tunneling Microscopy) AFM (Atomic Force Microscopy)

In both techniques the tip of the spectrometer is placed very close to the sample and can follow the contour of the sample at an atomic level. In the case of STM the tip distance is maintained by resistance to electron tunneling from the tip to the sample. In the case of AFM the tip contains an optical balance that is very sensitive to changes in force. It is known that as neutral molecules approach each other there is an initial attractive force followed by repulsion. These forces are the Van der Waals or Leonard -Jones forces. The tip approaches the surface so closely that it is sensitive to these forces and scans the surface in such a way that this force is maintained constant. Since the surface is essentially defined by the Van der Waals distance, the path of the tip defines the surface with atomic accuracy.

STM is applicable to conducting materials such as metals with a reasonably smooth surface. Since FCC catalyst materials have a rough surface and are non conducting powders, this method is has not been thought to be applicable. Recent results have shown that in the presence of ambient air an image of a silicalite surface can be obtained. Silicalite is a high silica zeolite with the structure of ZSM-5. The resolution is sufficient to show individual silicon atoms in six membered rings (79). AFM has also been used to obtain high resolution images of zeolite surfaces, showing pore openings and individual tetrahedra as well as the location and orientation of adsorbed organic molecules (80).

The Usefulness of the Surface Techniques The morphology and structure of the FCC catalyst is important, especially the location of

the catalytic components, the pore structure, and the location and chemistry of coke and metallic poisons. The instruments of most value so far have been the SEM (morphology), the electron microprobe (location and analysis of the chemical components), and XPS for a determination of oxidation states and chemistry of various elements on the catalyst.

201

Auger spectroscopy has been less useful primarily because the catalysts are insulating and charging effects tend to be relatively severe. The information obtainable by Auger spectroscopy (elemental analysis, oxidation states) is more readily available by XPS where charging can be compensated using the electron flood gun.

SIMS equipment is expensive and it is difficult to provide quantitative elemental analyses at this time, so its use has been limited. It can be especially destructive of the surface, depending on the mode of operation. However the potential for new information is so great that one can expect an increasing number of applications especially in research connected with catalyst deactivation. STM and AFM techniques are just beginning to find use. The roughness of the FCC catalyst particle has so far ma& these techniques difficult to apply.

5. SPECTROSCOPIC ANALYSIS (Infrared, FTir, Raman, Diffuse Reflectance, MASNMR, uv-visible, EPR)

Spectroscopic analysis deals primarily with the acidic properties of the catalyst including the zeolite and matrix components. The results provide information concerning the number, the strength, and the types (Br0nsted (protonic) or Lewis) of acidic sites present. Infrared and uv-visible techniques are used either to directly observe the properties of the acidic -OH group or to observe interactions between the acidic surface and an adsorbed probe molecule. Probe molecules include ammonia, pyridine, benzene, hydrogen, and other more or less basic molecules.

In silica alumina materials proton donor sites are usually associated with an aluminum atom connected to a silicon atom through an oxygen bridge as in the following structure, Figure 8.

si H ' 0

\

Figure 8. Schematic illustration of the acidic bridged OH function believed to be responsible for the acid cracking activity of zeolites.

The proton bonded to the oxygen can be characterized by an OH stretching frequency in the infrared and is distinguishable from other non-acidic or less acidic OH groups. Centers of both Brgnsted and Lewis acidity can be indirectly characterized by spectroscopically monitoring the interaction of the solid acid with various probe molecules. Frequently proton donor acids associated with an -OH group are called Bransted acids. In this discussion they will be referred to as protonic acids.

MASNMR (Magic Angle Spinning Nuclear Magnetic Resonance) also provides information concerning acidity, but from a different point of view. The chemical shifts of silicon, aluminum, and protons are observed. Since zeolites and cracking catalysts are solid acids one would expect the 'H spectrum to be the most useful. However the acidic properties of zeolites are determined by the lattice and by the environment of the proton rather than by

202

the properties of the proton itself. Consequently MASNMR is most often used to describe the silicon and aluminum containing structures that appear to determine the density and strength of the acid. . Fierro (81) has provided a recent review of the various spectroscopic techniques used for measuring acidity, The measurement and nature of acidity in solid acids such as zeolites has been discussed by Tanabe (82), by Rabo and Gajda (83), by Vedrine (84) and by C o m a (85). With the exception of MASNMR, spectroscopic methods are generally applicable to cracking catalysts. Thermal methods discussed in the next section are also used to characterize acidity, especially in zeolites. The division into thermal and spectroscopic methods is somewhat artilkid since these methods are frequently combined.

Infrared; FTIR; Diffuse Reflectance, DRIFTS Infrared, FTir (Fourier Transform Infrared), and DRIFTS (Diffuse Reflectance Infrared

Fourier Transform Spectroscopy) techniques all involve spectroscopic observations in the infrared region using slightly different instruments. Both infrared and FTir operate in the transmission / adsorption mode. The sample is prepared, the beam is transmitted through the sample, and wavelengths where absorption occurs are observed as absorption peaks (or transmission valleys). An infrared spectrometer will scan over the region of interest using a diffraction grating and a series of slits and mirrors to control the wavelength. The FTir instrument splits the beam into two components. By slightly changing the path length of one of the beams and then recombining the two beams an interference pattern is created. The interference pattern contains a spectrum of frequencies for which absorption is measured. A large number of scans are acquired over a period of time to build up the spectrum. Fourier transform techniques are used to transform the frequency spectrum into the conventional series of peaks. If the sample is spread on a reflection ball and the intensities of the reflected infrared are detected, the technique is DRIFTS.

. . . otonic Aciditv

Absorption bands due to the stretching vibration of the OH bond can be observed directly in the 3500 cm-I to 3750 cm-' region of the infrared (86.87). Some assignments are listed in Table 3. Assignments of strong or weak acidic character are based on combined infra redhase titration experiments. An intensity reduction is correlated with titration by some strong base such as ammonia or pyridine. The strongly interacting or acidic OH bands will show up as peaks in the difference spectrum. Typically there are weakly or nonacidic silanol OH groups that appear near or above about 3700 cm-'. These are considered to be lattice termination groups. Strong acidity is associated with OH groups in the 3610 cm-' to 3650 cm-L region. Weaker acidity is associated with groups in the 3550 cm-I range. An example of a typical spectra is shown in Figure 9. Steam dealuminated faujasite contains a strong protonic acid with a band at -3610 cm-I (88) that has been associated with the formation of a nonframework species in severely dealuminated faujasite (89-91). A band in the same region has been associated with strong acidity in ZSM-5 and other high silica zeolites (92). Infrared spectra have been used to confirm the suggestion by Kuhl (22) that some of the nonframework aluminum can occupy exchange sites in the zeolite (23).

Infrared spectra have been used to characterize the occurrence of hydroxyl nests, a lattice defect formed as a result of dealumination without silicon insertion or lattice rearrangement. The hydrolysis of the silicon aluminum bonds and the removal of the aluminum leaves four OH groups in the hole previously occupied by the aluminum atom. Defects of this type have been characterized by a broad infrared absorption in the 3000 cm-1 to 3750 cm-1 range (93). Hydroxyl nests are observed by comparing absorbances at a wave number such as 3710 cm-1 without overlapping contributions from other structures.

203

a c C n .c L C U s n

3900 3500

a

3610 3 6 3 0 3560 A

3 6 1 0

J- 3610

2-

b

B

A L

K

L

I I

1700 1500

c m-1

I LPV LPV

1400 1500 1600 1700

Frequency (crn-')

Figure 10. Pyridine adsorbed on silica alumina at 20O0C in successive doses 1-4 and (top curve) after desorption at 400OC. The large peak at 1450 cm-I is associated with Lewis sites and a much smaller peak at 1540 cm" is associated with protonic or Bronsted acidity. The figures and explanations are from N. C-Martinez and J. Dume- sic, J. Catal., 1990,125,427.

Figure 9. Infrared difference spectra of dealuminated faujasite before and after pyridine adsorption and evacuation 1) at 25OoC, 2) 35OoC, and 3) 400°C in two different infrared regions including a) the OH stretch region showing the occurrence of conventionally observed bands at 3630 cm-' and 3560 cm-1 and a strong protonic band at 3610 cm-I and b) the lower frequency spectrum of pyridine showing a large peak associated with protonic sites at 1540 cm-I and a smaller peak at 1450 cm" associated with Lewis sites, from G. Gar- ralon, A. Coma, and V. Fomes, Zeolites, 1989,9,84.

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J .ewis and Rot- . . . Direct observation of the OH stretch can only provide information concerning the

protonic acidity. Breck and Skeels have shown that dehydroxylation can occur at about 115O0F (62OOC) (94). The result of dehydroxylation is the generation of Lewis sites. The infrared spectrum of absorbed pyridine, an aromatic organic base, provides a method of distinguishing between protonic and Lewis types of sites. The technique has been used to study sites on matrix materials as well as zeolites. The pyridine is adsorbed, the sample evacuated to desorb excess pyridine, and the infrared spectrum is obtained. Pyridine adsorbed on a Lewis site gives a peak at about 1450 cm-'. A band at 1545 cm-' occurs in the case of the pyridinium ion and is characteristic of a protonic acid. A non-diagnostic band occurs at 1485 cm-' as a result of adsorption on either Lewis and protonic sites (95,96). Figure 9 shows an infrared difference spectrum of pyridine absorbed on a silica alumina catalyst both in the OH region and in the pyridine region. The peak assignments are given on the figure. Dealuminated zeolites contain significant protonic activity while a silica alumina catalyst contains primarily Lewis sites (97), Figures 9 and 10. -

One objection to infrared studies involving the absorption of strong bases like pyridine and ammonia is that these compounds are strongly basic as well as highly polar and will absorb on nearly anything including polar as well as weakly or strongly acidic sites. Consequently the acidity that is observed may not be the acidity that is catalytically important. An alternative is to use much weaker bases which will form a complex only with very strong acids.

Titration of the surface with a weaker base such as benzene shifts the band due to the interacting -OH group to a lower frequency. Since the shift is expected to be more or less proportional to the degree of interaction, the strength of the acid is correlated with the frequency shift of the - 3610 cm-' hydroxyl band in the infrared (98). Shifts to lower wave numbers by 300 cm-I to 350 cm-' are typical of strong acids, while weaker acids give smaller shifts in wave numbers. Similar experiments using hydrogen as the probe molecule are useful for probing the strength of Lewis sites. The adsorption of hydrogen on a strong Lewis site is associated with peaks in the 4000 cm-I to 4080 cm-' region. Lewis sites are observed on dealuminated zeolites by this method (99).

gv-visjble: weaklv basic dyes a nd aromahc Drobes One proposal is to use as a probe a series of molecules such as benzene, toluene, and

xylene (100). A sufficiently strong acid will form a sigma complex observable in the near uv at a wavelength of about 335 nm. Benzene requires a stronger acid than does toluene, and toluene requires a stronger acid than xylene. This method has been used to experimentally rank zeolites by their strong acidity. The ranking is the same as that obtained by catalytic methods such as a measurement of the turnover number for the cracking of n-hexane. Dealuminated faujasite, the zeolite used in cracking catalysts, has only intermediate strength sites and no strong sites. The number of intermediate strength sites increase with dealumination. Mordenite and ZSM-5 contain strong sites. It may be that high gasoline selectivity is associated with sites of moderate strength.

An approach suggested by Benesi (101) is to mix a weakly basic dye used as a Hammett indicator with the solid acid catalyst. If the acid protonates or interacts with the dye a visible color change will occur and one may infer the existence of acidic sites in the appropriate range of strengths. A series of dyes has been developed and has been used to characterize the acidity of various cracking materials. Recent work uses some of the same dyes but uses a uv spectrophotometer in place of the eye to measure the occurrence or non occurrence of the appropriate peaks in the uv region due to protonation (102). The results have shown that this is a reliable approach. The transitions associated with the protonic form of the dye occur in

205

the uv region. Since the eye is not sensitive in the uv region, visual color observation is misleading and spectroscopic observation is required. Acid strength is given in terms of the Hammett function Ho where 100% sulfuric acid has an Ho of -12. Super acids are stronger than 100% sulfuric and have Ho < -12. The results using the spectrophotometer show that mordenite is a moderate super acid. It contains strong acid sites marginally stronger than 100% sulfuric acid with an Ho value of - -13. Faujasite has an acid strength equivalent to about 95% sulfuric acid, Ho - -9 to -1 1. Silica alumina (13% alumina) is equivalent to about 70% sulfuric acid, Ho - -4 to -8.

One drawback of this method of characterizing FCC catalysts is the observation that the regenerator of an FCCU typically operates above 1250'F (68OoC), while Skeels and Breck have shown that zeolites dehydroxylate at about 1150'F (620'C). On the other hand, steam is present and hydrogens are available for rehydration. Consequently the sites observed by room temperature measurements may not be the active sites present in the operating catalyst. Infrared measurements in the OH stretching region during the cracking of cyclohexene made at operating temperatures show the presence of strong proton acidity in situ on a dealuminated zeolite catalyst and suggest the involvement of these protons in coke formation (103).

Raman Raman spectra have been used primarily to observe vibrations characteristic of the silica

alumina framework. There are peaks than can be related to the silicdalumina ratio of faujasite (16, 17). Raman spectroscopy can be used to identify zeolite framework structures, and also the structure of templates that may be present.

MASNMR (Magic Angle Spinning Nuclear Magnetic Resonance) NMR has become a favorite tool of organic and physical chemists for structure

determination and identification of organic compounds. Any nucleus with spin one half produces a very sharp resonance in solution. 13Carbon and 'H have been especially useful since these are the major components of organic compounds. Solid state catalytic applications are more recent. For these applications the nuclei studied include 27Al, 29Si, 'H, and 129Xe. Since FCC catalysts are solid acids composed predominately of silica and alumina, the applicability of 27Al, 29Si, and 'H MASNMR is obvious. The acidic properties of these materials are a result of an aluminum atom tetrahedrally connected by means of an oxygen bridge to four silicon atoms. In this situation the aluminum has a formal negative charge balanced by a positively charged proton located on or near one of the bridging oxygen atoms, Figure 8. 27Al MASNMR has been used to characterize the Occurrence of least four types of aluminum structures in zeolites and other catalytic materials including three types of nonframework alumina and two types of classically acidic tetrahedral forms, one framework and one probably nonframework. 29Si MASNMR has been used to quantify the Occurrence of silicon with zero, one, two, three and four aluminum neighbors and to indirectly infer the degree of aluminum site isolation. Since 29Si has a relatively low (4.7%) natural abundance, the acquisition of a spectrum may require several hours and thousands of scans.

More recently 12%e, 15N, and 31P NMR techniques have become useful. Nitrogen containing compounds are frequently used either as templates, adsorbents, or as exchange cations, and 129Xe is used to investigate the pore geometry of zeolites as well as catalysts. Phosphorus appears in the framework of a new class of molecular sieves containing phosphorus as well as aluminum and silicon. The term zeolites is normally restricted to silica alumina containing structures, while the term molecular sieve is used more broadly where the framework may contain significant amounts of other elements, eg. phosphorus.

MASNMR is used to characterize the chemical bond involving a particular element, where the chemical shift is characteristic of the number and identity of the bonding atoms.

206

Table 4 gives correlations between chemical shifts and the bonding environment for silicon, aluminum, and phosphorus bonded to each other. SiOP bonds are normally not observed. While the chemical shifts are characteristic of the number and chemical identity of nearest neighbors, for a given chemical environment small variations in the chemical shift are observed that are a result of variations in the bond angle and other bonding parameters. For each environment the range of shifts observed is listed in Table 4.

Table 4. and 31P.

MASNMR chemical shifts (ppm) associated with structures containing 29Si, 27Al,

2 9 s ,,(OAI) -tive to TMS (Tr-

n = Q 1 2 3 4

105 to 107 95 to 105 88 to 96 86 to 92 80 to 86 ~

27-tirelative to ~1 3+

Coordumn

Qsi m m l2S.i mi

. . Tetrahedral Octahedral -oordina&

20 to 35 -5tolO -20 50 to 70 35 to 45

Octahedral u

-15 to 0

The basic interaction is between the magnetic moment of the spinning nucleus and the magnetic field of the instrument. In the case of spin one half there are only two orientations, parallel or anti-parallel to the field. The sample is placed in the field of the probe, a radio- frequency oscillator. When the frequency of the probe oscillator corresponds to the energy of transition from parallel to anti parallel, the nucleus starts to precess and to change its orientation. This is associated with a change in magnetization measured in the probe. The position of the NMR signal of a bare nucleus would depend only on the nuclear magnetic dipole moment. However, in a chemical compound the nucleus is surrounded by moving electrons which also respond to the magnetic field in such a way as to produce shielding or deshielding effects. Depending upon the electronic environment, the local magnetic field at the nucleus will be slightly changed. The transition frequency will be shifted by an amount depending on the electronic environment of the particular nucleus. It is the magnitude of this shift, called the chemical shift, that provides useful chemical information.

In a powder there are several sources of line broadening. Molecular orientation varies in a solid powder, and since the chemical shift depends on orientation, the signal will be broad, a result of chemical shift anisotropy. A second source of broadening is the dipolar interaction. Each nucleus interacts with other nearby nuclei. Since the nucleus is a small magnet, nearby nuclei contribute to changes in the local magnetic field. The degree of interaction depends on the relative orientations of the interacting nuclei to the field. The dipolar interaction is only a factor for neighboring nuclei. It is not a factor in the case of low natural abundance isotopes

207

such as 29Si. A third source of broadening is a result of the quadrupolar interaction between a nucleus of spin greater than 1/2 and a non spherically symmetric distribution of valence electrons. The interaction results in a line broadening due to unresolved multiplets.

In solution the average orientation for each molecule is zero over the time required to obtain the signal. For spin 1/2 the peaks are sharp and well resolved and the position of the peak relative to some standard gives the chemical shift information. In a solid the orientation is fixed, leading to a broad signal that obscures the chemical shift information. However, if the sample is spun at an angle of 54.7Oto the imposed magnetic field. the orientation effects will average to zero. Such an experiment is called Magic Angle Spinning Nuclear Magnetic Resonance (MASNMR). The application of MASNMR gives relatively sharp resonance lines containing the desired chemical shift information in a variety of solids, including catalysts and zeolites. Comprehensive reviews have been provided by Fyfe (104) and more recently by Engelhardt and Michel(lO5).

sof 2 9 w ~ ~

The silicon chemical shift in alumino-silicates depends on the number of neighboring aluminum atoms. There are five types of silicon that can be distinguished by their chemical shift in the 29Si MASNMR of faujasite. They correspond to silicons bonded through oxygen linkages to 0, 1, 2, 3, or 4 aluminum atoms in the structure of the faujasite schematically shown in Figure 11. The chemical shift depends to a lesser extent on structural details such as the bond angle and bond length associated with the particular structure. In favorable circumstances the variation in chemical shift can allow an estimate of the bond angles in a particular zeolite structure (106). Consequently there is a fairly broad range of chemical shifts that can be associated with each of the five types of silicon chemical environments, The chemical shift range appropriate to different silicon environments is provided in Table 4 (107). Amorphous non-zeolitic silica can often be further identified at a higher shift if present.

The usefulness of the 29Si NMR spectra depends on certain characteristics of the zeolite structure. The structure consists of silicon and aluminum tetrahedra linked through oxygen. The tetrahedral aluminurn has a formal negative charge neutralized by a cation, typically sodium or hydrogen. The hydrogen exchanged form is the acidic form. It has been proposed that aluminum tetrahedra will not link together because of the repulsion of the two adjacent negative charges. This rule, known as Lowenstein’s rule, is based on generalized bonding considerations suggested by L. Pauling and is obeyed for zeolite structures (108).

Consequently an aluminum will always be surrounded by four silicon tetrahedra. The total number of aluminum atoms per unit cell, #Al/uc, in the structure is therefore one fourth of the total number of silicon - aluminum bonds. Since the number of silicon atoms in various possible environments is proportional to the MASNMR intensities,

#Al/silicon atom = Si(lAl)/4 + Si(2Al)/2 + 3Si(3Al)/4 + Si(4Al).

where the numbers of silicon in the various environments are given by the respective MASNMR intensities normalized to a total intensity of one. The silicon to aluminum ratio R of the framework is the inverse. The number of aluminurn atoms per unit cell is given by #Al/unit cell = 192/(R+1). This :lationship has been used to define the relationship between the unit cell size and the alum urn content of zeolites, Table 1. Unfortunately cracking catalysts contain sources of sili, n other than the zeolite. These other sources such as clay and the silicon components of tl binder broaden the peaks and interfere with the application of 29Si MASNMR directly to cr; dng catalysts.

As a result of Lowenstein’s rule the distribution of aluminum atoms can be indirectly inferred from the distribution ol the silicon intensities. The distribution of aluminum is important since issues such as coke selectivity, hydrogen transfer activity, and acid site

208

strength may be related to the degree of site isolation. As the SUAI ratio of faujasite increases there is a tendency towards site isolation reflected in the loss of intensity in peaks associated with Si(4Al) and Si(3Al) relative to the intensities expected of a random distribution (111 - 113). At higher SUAI ratios persistent intensity associated with the Si(2Al) peak shows that a degree of site pairing exists even at low unit cell sizes (110). It has been proposed that hydrogen transfer, coke selectivity, and octane selectivity can be correlated with the relative amounts of paired and isolated sites (1 14). A difficulty with this analysis is that the chemical shift due to an SiOH bond is similar to that due to SiOAl (1 15). Since dealuminated zeolites contain a large internal surface terminated by SiOH groups, the quantitative estimation of the number of silicons linked to zero, one or two aluminum atoms is uncertain.

Si(2AI)

-94

i--, , , , , , , , r , , , , , , , , , , , , . .

PPm -80 -90 -100 -1 10

Figure 11. 29Si MASNMR spectrum of faujasite, SUAl - 2.5, taken at 104.3 MHz with peak and structure identifications.

Cross polarization techniques are often used to indicate the existence of contributions to the 29Si NMR spectrum from surface or defect structures where any surface terminated by SiOH is a defect structure. In this experiment the intensity of the SiOH group is enhanced relative to the intensity of the Si-0-Si and Si-0-Al groups because of the transfer of magnetization from the 'H nucleus to 29Si. Cross polarization experiments have shown that hydrothermally dealuminated zeolites contain relatively large amounts of terminal SiOH groups, presumably on the internal surface. These results are consistent with the observation by other methods of the occurrence of significant mesoporosity in zeolites.

209

Applications of 27Al MASNMR 27Al MASNMR has been used to characterize at least four types of aluminum structures in

zeolites including three types of nonframework alumina and two types of classically acidic tetrahedral forms. It has also been used to characterize the occurrence of aluminum species in silica alumina materials other than zeolites. The chemical shift range appropriate to different silicon environments is provided in Table 4.

32.3

3.9

I 13.5 KHz

7.8 KHz

,1

300 200 100 0 -100 -200 - 3 0 0 PPm

Figure 12. 27Al MASNMR spectra of hydrothermally treated faujasite taken at 104.3 MHz and at different spinning rates showing the variations in intensity for the various alumina species. From A. W. Peters, W. C. Cheng, M. Shatlock, R. F. Wormsbecher, and E. T. Habib, Jr., in Guidelines for Ma&xbg the Properties of ’ D. Barthomeuf, E. G. Derouane, and W. Holderich, eds., Plenum Press, New York, 1990, p 365.

210

Tetrahedral aluminum as it occurs in the faujasite framework has a chemical shift of about 60 ppm relative to the octahedrally coordinated aqueous aluminurn cation used as a standard (1 16). This aluminum species can occur in other zeolite framework structures such as beta and ZSM-5 at a slightly lower shift of 52-57 ppm (1 17,118). Alumina, silica alumina, and hydrothermally treated zeolites may contain octahedral (six-fold) aluminum characterized by a peak with a chemical shift of about 0 ppm (1 19, 120). Pentacoordinated alumina occurs at a shift of about 30 ppm in the mineral andalusite. in fine particle alumina, in severely hydrothermally dealuminated zeolites, and in other hydrothermally treated materials such as clay and silica alumina gels (121, 122). The spectra in Figure 12 are characteristic of hydrothermally treated faujasite. The relative intensities of the three peaks can depend strongly on the spinning rate (1 14).

More than one aluminum species with a shift in the tetrahedral region has been observed, although with difficulty. Since the 27Al nucleus has a 512 spin, the spectrum is broadened by quadrupolar effects. More than one resonance with about the same chemical shift is difficult to observe under normal conditions. However, it is possible to set up the MASNMR instrument in such a way as to distinguish resonances from non-equivalent atoms that happen to have nearly the same chemical shift. In a two dimensional experiment the chemical shift spectrum is observed under a variety of pulse times. Atoms in different chemical environments will have different quadrupole coupling characteristics and will have different intensities under different pulse conditions. The line shape due to the different components of the resonance will change with pulse time leading to a two dimensional plot (123). Quantitative spectra are usually obtained at shorter pulse times (- 15' angle). Recently double rotation (DOR) experiments have been described that reduce quadrupolar broadening as well as ansiotropic effects ( 124).

60.1

120 80 LO 0 -LO 120 80 LO 0 -LO PPfl PPH

Figure 13. 27Al MASNMR spectra of faujasite dealuminated to a) 2.422 uc and b) 2.420 uc showing a doublet in the tetrahedral region at 60.7 ppm and at 54.3 ppm for sample a). From A. Corma, V. Fornes, A. Martin&, and J. Sanz, Fluid Catalytic Cracking, M. Occelli, ed., ACS Sym. Ser. 375,1988, p. 22.

21 1

In severely dealuminated faujasite a non framework tetrahedral species has been observed by 2 dimensional MASNMR as well as conventionally as a slight split in the spectrum of specially prepared samples, Figure 13 (125, 126). In these materials tetrahedral aluminum occurs with both a 60 and a 55 ppm shift. The occurrence of this peak is independent of the dealumination method (1 15).

4 l l h a W a 3 I - . .

Applications of 31P MASNMR are relatively recent and are used to characterize the coordination of phosphorous in catalysts and in zeolites. Incorporation of phosphorous into S A P 0 or ALP0 types of zeolites has become common. Phosphorous is believed to form bonds only to aluminum (not silicon) in these structures. The presence of phosphorous shifts both the octahedral and tetrahedral 27Al MASNMR signals about 20 ppm in the negative direction, Table 4. The tetrahedral 31P peak occurs at a shift of about -20 to -30 ppm compared to phosphoric acid. As multi-nuclear probe cross polarization techniques become more generally available it will be possible to identify bonds between a phosphorous with a particular shift and another element, say aluminum, also with an identifying shift.

-s . . of 1 2 9 ~ ~ NMR

129Xe NMR, recently reviewed by Dybowski, Bansall and Duncan (127), is also an adsorption method in that Xe is adsorbed on the surface. The observed chemical shift relative to free Xe is related to the size and the shape of the pore in which the Xe is confined. Recent work has shown that molecules in very small zeolitic pores spend most of the time close to the surface of the pore as a result of conventional attractive Lennard-Jones type attractive forces. This interaction is specific to small pores and is known as the confinement effect (128). J. P Fraissard and coworkers have shown that the chemical shift observed for 129Xe is sensitive to interaction with the pore wall (129). I2%e NMR has been applied to the characterization of the pore structure of zeolites and other materials. An application to

Table 5. MASNMR. The micropore size is given in the number of atoms of Xe per micropore and as the volume A3 per micropore. The micropore capacity is the number of micropores per gram X volume per micropore as mmoles Xdgrarn (cm3/grarn). From Reference 132.

Analysis of the pore structure of fresh and equilibrium FCC catalysts by 129Xe

Number of Micropore size Micropore % Zeolite Micropores, # Xe and Capacity

walk ( A3/micropore)

Catalyst A Fresh 169 9.1 (406) 1.54 (0.041) 26 Equilibrium 108 8.4 (375) 0.91 (0.025) 16

Catalyst B Fresh 133 9.1 (406) 1.20 (0.032) 20 Equilibrium 81 8.2 (366) 0.67 (0.018) 12

Reprinted from: T. T. P. Cheung, J. Catal., 1990,124,5 11.

212

dealuminated faujasite showed that the 129Xe NMR results could be used to provide information concerning changes in the size of the microporous zeolite cages during dealumination as well as to observe the development of mesoporosity (130, 131). The change in size as measured by a change in the chemical shift is presumably a result of alumina deposits formed during dealumination or partial framework collapse. Subsequent results obtained on a series of fresh and equilibrium FCC catalysts showed that 12'Xe NMR could be used to describe catalyst stability by counting the number of surviving microporous cages (132), Table 5. In addition an estimate of the change in the size of the cage is obtained from an estimate of the number of '29Xe in each cage. Changes in the pore structure of the zeolite in a coked FCC catalyst coupled with low pressure argon adsorption showed a change in the chemical shift that correlated with the chemistry of the coke. The adsorption results showed some diminution in cage size associated with the coke deposits (133).

Applications of 'H MASNMR While MASNMR techniques were applied to obtain proton ('H) magnetic resonance spectra very early in the development of the MASNMR technique, the interpretation of these spectra is not yet entirely settled. Peaks observed in the 1.5 to 2.5 ppm shift region relative to a TMS (trimethylsilane) standard have been attributed to surface silanol groups (134, 135). and a peak at 6-7 ppm has been attributed to a hydrated Lewis site (136). Peaks in the 4-5 ppm chemical shift range have been identified as being associated with protonic acidity and seem to require the presence of at least some water. Current peak assignments have been reviewed along with a discussion of correlations between protonic acidity and chemical shift (137). The 'H NMR technique is difficult and the spectra are often not well resolved. Some of the experimental difficulties are discussed in a review by Freude (138).

6. THERMAL ANALYSES (TPD, TGA, DSC, TPR, TPO, Microcalorimetry)

With the exception of microcalorimetry, the thermal analysis methods involving temperature programming usually give kinetic information rather than equilibrium information. This kinetic information concerns rates of adsorptioddesorption (TGA, TPD), rates of oxidationheduction (TPO, TPR), and rates of phase changes and sintering (DSC). Consequently the analysis of the results is an important issue.

Often the results are analyzed as though they gave equilibrium information. Ammonia TPD may be interpreted as though the temperature of desorption were related to the strength of the ammonium bond to the assumed acid site. The experiment gives this result only indirectly by answering the question, how fast does ammonia desorb from the sample as we increase temperature? Attempts to measure the stability of catalysts or zeolites using DSC results have the same problem. The sample is rapidly heated and one observes a peak or a valley at the temperature of sintering or of a phase transition. However, it is the activation energy and the rate of sintering that has physical and chemical meaning, not the temperature. The rate parameters are obtained by observing changes in the peak shape and in the apparent temperature of sintering as the heating rate is varied. Consequently the kinetic parameters must be extracted from the DSC curve to make a meaningful comparison (139). The same is true of TPD experiments (140). This is normally the situation for methods involving temperature programing. Only recently has this kind of an analysis been done for ammonia TPD experiments on zeolites where the result is a desorption activation energy. The recent review by Bhatia, et. al. emphasizes kinetic analysis (141). The monograph by Turi reviews thermal methods, but does not discuss applications to catalysis (142). Modern commercial thermal analysis equipment often includes software for the analysis of the desorption kinetics or the kinetics of phase transformations.

213

Microcalorimetric methods involve equilibrium measurements and give heats of adsorption directly. The interpretation is straight-forward. The strength of the interaction between the solid acid and the adsorbate is measured by the heat of adsorption. However, data collection is much more time consuming. The measurement of the heat of adsorption for a single system requires many thermal measurements. The heat generated is measured for each incremental amount of adsorbate added to the system, and each measurement may require hours in order to achieve thermal equilibrium.

TPD (Temperature Programed Desorption) In this experiment a base such as ammonia is adsorbed on the solid acid and the system is

evacuated to remove excess ammonia. The sample is heated and the thermally desorbed ammonia is detected as it is removed either by a mass spectrometer or by some other detector. In the case of zeolites such as faujasite or USY the result is usually two peaks, Figure 14. These two peaks are ascribed to the occurrence of strong and weak acidity corresponding to the high and lower temperature peaks respectively. Sometimes these peaks are further identified with (or confused with) protonic and Lewis acidity.

Figure 14. Ammonia TPD of a) dealuminated faujasite with high (<350"C) and lower temperature peaks, A. Corma, V. Fornes, F. V. Melo, and J. Herrero, Zeolites, 1987.7.559 and b) ZSM-5 showing both the low temperature alpha peak at about loO°C, the low temperature beta peak at about 200 to 250°C and the well separated high temperature gamma peak at > 400°C characteristic of the stronger acidity of the ZSM-5, from N.-Y. Tops@, K. Pedersen and E. G. Derouane, J. Catal., 1981,70,41.

The intent of the test is to measure the number of acid sites and to estimate the strength of the interaction with the adsorbate. In measuring the number of sites it is assumed that each adsorbed ammonia accounts for one site. The strength measurement assumes that the heat of adsorption of the ammonia on the acid catalyst is proportional to the strength of the acid, and that the temperature of desorption is proportional in some way to the heat of adsorption. The test is normally used in a semi quantitative way to count the numbers of strong and weak sites, where the site strength is qualitatively given by the peak temperature of desorption. The interpretation of the comparison curves in Figure 14 is that ZSM-5 contains stronger acid sites than faujasite.

214

Ammonia TPD experiments suffer from most of the problems generally associated with temperature programming methods. as well as additional ones associated specifically with ammonia TPD. The analysis is difficult since there are a number of experimental issues to be dealt with. There is an obvious diffusion issue both through the bed and within the catalyst or zeolite particle. During the experiment it is likely that the ammonia will be desorbed and readsorbed many times depending on the bed thickness, particle size, the partial pressure of ammonia in the bed, and the residence time of the ammonia in the bed.

The identity and the origin of the observed high and low temperature peaks is not clear. In the case of faujasite, the area under both peaks, the total amount of ammonia absorbed, is used to count the total number of sites per gram (143). It correlates with the total alumina content. In the case of ZSM-5 there are also two peaks with better separation (144). However, in this case the aluminum content is measured only if one counts the ammonia desorbed at the higher temperature. The origin of the lower temperature peak is unclear. While the low temperature peak occurs at about the same temperature for both zeolites. the high temperature peak in the case of ZSM-5 occurs at a higher temperature, consistent with the proposed stronger acidity of ZSM-5. The somewhat arbitrary nature of the counting rules suggest that ammonia TPD is not a reliable method of counting the acid sites involved in catalytic cracking, especially on unknown materials.

The problems with ammonia TPD are related to uncertainty in the assumptions underlying the method. It is assumed that since ammonia is a base the strength of the interaction is related to the strength of the acid, and that the sites measured are related in some simple way to cracking activity or selectivity. These assumptions are not necessarily reliable. Ammonia, besides being a strong base, has a significant polarity and can interact strongly with other polar oxides (145). Zeolites also contain significant polarity as well as acidity, so there is no guarantee that the interaction with the ammonia is simply an acid base interaction. The pore size of the zeolite as well as the polarity of the framework may contribute to the apparent strength of the interaction.

In spite of the difficulties recent work has provided reliable values of about 100 W h o l e (24 Kcdmole) to 120 KJhnole (29 Kcaho le ) for the activation energy of the desorption of ammonia on dealuminated Y zeolites (146). A second recent measurement gave 116 W h o l e (27 Kcaho le ) and 122 KJ/mole (30 KcaVmole) for the desorption activation energy of ammonia on two samples of dealuminated faujasite (147, 148). Both studies contain detailed discussions of the experimental and theoretical difficulties associated with ammonia TPD on zeolites. and saturated hydrocarbons have been obtained on catalysts and on zeolites (149,150).

A recently developed method, isopropyl amine TPD, avoids most of these difficulties and may successfully count acid sites of some certain strength. The method involves the adsorption of isopropyl amine on the solid acid, removal of excess adsorbate by evacuation, and a temperature ramp up to about 500OC using preferably a mass spectrometer as a detector. The desorption of isopropyl amine is first observed at some low temperature, followed by the desorption of cracked products. The number of strong acid sites capable of cracking/deamination is measured by the amount of products, propene and ammonia, desorbed. Results with steamed cracking catalysts have shown that the activity of the steamed catalyst for gas oil cracking is proportional to the number of cracking sites measured by the method (15 1).

Surface migration and reabsorption are possible sources of error. However, the initial removal of the isopropyl amine from the surface and the destruction of the adsorbed species during the cracking reaction should minimize these difficulties. While this method counts the number of sites capable of forming the required carboniudcarbenium ion intermediate, it does not provide an estimate of relative site strength. This method avoids at least in principle some of the objections to the use of ammonia as a TPD probe for acidity as it relates to catalytic cracking.

The kinetics of desorption of a variety of aromatic, olefiiic,

215

Microcalorimetry Both DSC and TPD techniques attempt to measure acidity by measuring the heat of

adsorption of a base on a solid acid. Direct calorimetric methods are probably more accurate for this measurement as well as somewhat more complicated than TPD or DSC methods. A recent application is the measurement of the heat of adsorption of pyridine on a silica alumina catalyst. The results combined with infrared results showed that the strongest adsorption occurred on Lewis sites (97).

TPR (Temperature Programmed Reduction) TPO (Temperature Programmed Oxidation)

TPO or TPR measures the oxidative or reductive stability of a sample. The sample is heated in an oxidative or reductive environment and the rate of disappearance of the oxidant or reductant, e.g. 0 2 or H2, is observed. Cracking catalysts are not intrinsically oxidative or reductive, so the technique has limited usefulness. However, contaminants such as nickel and vanadium have oxidation-reduction cycles. In the case of nickel contaminant, the observation of differences in the ease (temperature) of reduction have been related to the degree of interaction between the nickel and the support (152, 153). The implication is that less easily reduced nickel will produce less hydrogen in the FCCU. Since the TPR experiment, like the other experiments involving temperature programming, measures kinetics, the conjecture is a reasonable one. It has been suggested that the TPR profile of metals on a cracking catalyst is characteristic of the activity of the contaminant metal, and that it may be possible to use this profile as a criterion for laboratory deactivation procedures (154). The lab impregnated and steamed catalyst should have the same TPR profile as the unit deactivated or equilibrium catalyst.

TGA (Thermogravimetric Analysis) In using this technique, the sample is placed in a very sensitive microbalance. The sample

may be exposed to various gases and the weight change noted. The sample may also be heated and the weight loss of the sample is detected. If the desorbing gas is detected, then TPD results can be obtained at the same time. TGA results have been used by Breck and Skeels to show the dehydroxylation of zeolites at about 600OC (1 150OF) (94).

TGA experiments have been used to demonstrate the pick-up and release of sulfur oxides in the development of additives for SOX control in the FCCU. The additive picks up SOX as So3 in the regenerator and subsequently reduces and releases sulfur as HIS on the reactor side of the FCCU. This sequence of events forms the technical basis for the operation of the SOX removal additive. The TGA follows the pick up of SO3 as a weight gain. As hydrogen is added to the catalyst containing the bound sulfate the sulfate is reduced to sulfide and is released as H2S. The reduction and release is observed as a weight loss (155).

DSC (Differential Scanning Calorimetry) DTA (Differential Thermal Analysis)

In differential scanning calorimetry the sample and a reference are heated sufficiently to rapidly increase the temperature of the sample at a rate of between 10 and 100 degrees C per minute. The occurrence of an exothermic or an endothermic phase change in the sample will either release or absorb heat. The amount of heat required during the phase change to maintain the temperature ramp will be different for the sample than for the standard. This difference is measured as a function of temperature. During the phase change there will be either a positive or an inverted peak. This type of experiment is frequently done on cracking catalysts as well as the catalyst components. Zeolite will undergo an exothermic transition to either an amorphous phase or to christolbalite and mullite. Clay, a common inert ingredient of cracking catalysts, is known to go through several changes, first to metakaolin, to a spinel structure, and to mullite, all with well defined phase transitions. The information that one gets is the kinetic parameters of the transition. The temperature is ramped at several rates. As

216

the rate of temperature ramping increases the DSC peak moves to a higher temperature and changes its shape, usually becoming narrower. Kinetic parameters can be obtained either by measuring the change in peak position or in peak shape or both as the temperature ramp rate is changed (138). It is also possible to directly measure the heat released during a transition (156) or the heat capacity of a zeolite or catalytic material (157).

There have been several studies of the stability of various forms of dealuminated Y zeolite (158 - 160). In both cases temperatures of sintering were reported rather than the more meaningful activation energy associated with sintering. A study by Pompe and others used essentially all of the thermal methods including DTA. TPR and TGA as well as XRD to show that vanadium forms a complex with rare earth in the form REV04. The study involved lab Ni, V impregnated commercial catalysts (161).

7. SUMMARY OF CURRENT TRENDS

One area of current interest is the development of high temperature controlled atmosphere in-situ analysis systems. Catalysts are observed at ambient or some other convenient temperature, in an oxidizing or neutral chemical environment, and at ambient or sub-ambient pressures. The FCC catalyst operates under very different high temperature conditions in a strongly reducing environment at pressures of 1-2 atmospheres. There is an increasing interest in observing catalysts under operating conditions. This involves developing in-situ methods of analysis. These methods may involve combining thermal and spectroscopic methods in a single experiment where the acidity or some other property of the active catalyst is spectroscopally observed under higher temperature conditions more representative of the operating environment (103). Other examples of in-situ methods of catalyst characterization, not necessarily related to FCC, have been recently collected (162).

While a discussion of the methods available and in use for the analysis of FCC catalysts tends to deal with individual topics and methods, in practice a variety of methods are typically combined to give as complete a picture as possible of the catalyst. This kind of an approach is emphasized in discussions of the philosophy and guidelines for catalyst testing (163, 164). The interaction between performance evaluation and characterization of FCC catalysts (165) is discussed in another article from the same symposium appropriately entitled

and W y s t Development. An Interactive Appim&, One example of an interactive approach (among many) is the work of Beyerlein and others in characterizing a spent FCC catalyst from an operating unit (59).

Many of these methods are involved in catalyst process and quality control. This is especially true of the diffractive and adsorption methods discussed in sections 2 and 3. X-ray diffractive equipment and nitrogen adsorption equipment are commonly operated in the plant along with the usual methods of elemental analysis. As spectroscopic equipment becomes less expensive and more easily operated, it is also expected to find process and quality control applications.

8. ACKNOWLEDGEMENTS

I would like to thank R. Kumar and N. D. Spencer of W. R. Grace for their helpful suggestions and comments.

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AM-87-69.

’ W. M. Meier and


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CHAPTER 7

MICROACTIVITY EVALUATION OF FCC CATALYSTS IN THE LABORATORY:

PRINCIPLES, APPROACHES AND APPLICATIONS

E.L. Moorehead (l), J.B. McLean (2), and W.A. Cronkright (1)

(1) The M. W. Kellogg Company, Technology Development Center 16200 Park Row, Houston, Texas 77084-5195

(2) Engelhard Corporation, 1800 St. James Place, Suite 501 Houston, Texas 77056

1. INTRODUCTION

The laboratory evaluation of Fluid Catalytic Cracking (FCC) catalysts has evolved into a very common method for measuring performance characteristics of experimental and commercial catalyst samples. While many testing philosophies have been developed over the last twenty years, the most common method employed within virtually every laboratory, makes use of the Microactivity Test Unit, or MAT Unit.

The use of the MAT unit as the primary tool for the laboratory evaluation of FCC catalysts is the subject of this chapter. A brief history of the MAT will be followed by more detailed discussion of how this test is used for assessing performance of catalysts. Essential to this is the need to discuss the pre-treatment of fresh catalysts, i.e. steam deactivation, testing the metals (nickel and vanadium) tolerance of catalysts, and the analytical advances that have been made over the last 20 years.

The Microactivity Test is not the only approach for assessing catalyst performance, but it is the single most common one. Alternative or complimentary approaches for evaluating FCC catalysts make use of larger pilot units; the most common designs being developed b,y ARC0 (1) and Davison (2). As important as this testing protocol is, it is not the subject of this chapter.

The performance of a FCC catalyst involves more than just its catalytic performance. Issues of fluidization, attrition resistance can be of equal or greater importance in some instances. A general review of FCC catalyst evaluation was published by Rawlence and Gosling (3).

224

Finally, it is important to realize that there is no one correct method for evaluating the catalytic performance of FCC catalysts. While a standard ASTM method (4) for MAT testing exists, it is not used to any great extent, although the method is widely referenced. A variety of methods and philosophies have been developed over the years to meet specific objectives. In most cases, the test methods adapted by each laboratory have met their specific objectives. In essence, each laboratory tests differently and each laboratory is right in what they do.

2. HISTORY OF THE MICROACTMTY TEST

The laboratory evaluation of cracking catalysts has a history dating back to the late nineteen forties (5-11). Cracking catalysts at that time were based on amorphous silica-alumina technology. Protocol for the typical test, such as Atlantic's D + L procedure, required large catalyst and oil volumes (200 g catalyst, 50 g oil) with residence times extending up to 12 minutes. The need for the large catalyst and feed volumes was principally dictated by the analytical tools available. Specifically, with the absence of gas chromatographic simulated distillation (GCSD) technology, the, analysis of the cracked oil required batch distillation.

In the early sixties, the introduction of molecular sieve based cracking catalysts resulted in a new era for catalytic cracking technolop The new catalysts were far more active than those they replaced but, the emsting test methods could not accurately differentiate their performance, although commercial performance was clearly superior (12). In particular, the existing test methods did not show the activity advantage for zeolite catalysts. It was generally believed that this was due to the long reaction times at low space velocities. This resulted in high coke yields that masked the intrinsic activity of these catalysts (12).

About this same time, gas chromatographic simulated distillation (GCSD) technology was developed enabling small amounts of cracked oil to be characterized (13-16). The need to differentiate the performance of the high activity zeolite catalyst and the development of the GCSD method for estimating boiling point distributions paved the way for the development of a new test using smaller amounts of both catalyst and oil.

It is generally accepted that the Microactivity Test was first introduced by Davison Chemicals Division in 1967 (12). However, the authors of this paper accurately give credit to Atlantic Richfield as being the initial developers of a test that required small amounts of oil and catalyst samples operating at shorter residence times. In its earliest form, the MAT unit employed 3/32 inch pellets operating at a temperature of 900'F. The catalyst pellet was employed to minimize pressure drop across the catalyst bed. The catalyst charge was 5 g and the oil charge was 1 cc (0.86 g). The test as defined here was able to differentiate the performance between amorphous silica- alumina catalysts and zeolite containing catalysts (12). In particular, activity advantages and coke advantages for the new catalysts were clearly evident. Excellent correlahon of coke yield and conversion was reported.

225

Further improvements in reactor design and product collection were introduced by Gustafson in 1969 (17). These improvements resulted in near isothermal conditions at temperatures of 900'F. Reaction temperatures as high as 950°F were reported with temperature drops during cracking of 35-50'F. Further, the catalysts were evaluated as a powder as opposed to the 3/32 inch pellet to improve heat transfer and reduce mass transfer effects.

The current ASTM referenced MAT procedure (4) has its roots in the work by Gustafson (17) and Ciapetta and Henderson (12). This standard test procedure operates at W'F, uses 4 grams of catalyst and 1.33 grams of oil, giving a catalysttoil ratio of 3. The oil is injected over a 75 second time period. The weight hourly space velocity (WHSV), defined by (3600l(Cat/Oil*Injection Time) is 16. It is this standard test method that is most widely referenced, but not generally practiced as written (18). It should also be noted that the ASTM test is technically applicable only for catalyst activity determinations and does not address measurements of catalyst selectivities. A suitable test for selectivity determinations is being reviewed by ASTM. A tabular summary comparing the MAT procedures of Ciapetta et al, Gustafson, and the ASTM procedure is presented in Table 1. A schematic of a typical MAT unit is presented in Figure 1.

Table 1 Comparative Microactivity Test Conditions

Ciapetta Gustafson ASTM

Catalyst Charge, grams 5 NIA 4 Catalyst Shape 3/32 in microsphere microsphere Oil Charge, grams 0.86 NIA 1.33 CatalysttOil (wh) 5.8 3.5-4.0 3.0 Temperature,oF 900 900-950 900 WHSV, Hr-1 2-16 16-64 16

3. APPLICATIONS OF THE MICROACTM'IY TEST

The MAT is the primary tool in accessing the performance of catalysts Sam led from

The former are commonly referred to as "equilibrium catalysts" or "E-Cats". The importance in testing such equilibrium catalyst samples has been discussed by Upson (19). In particular, the assessment of catalyst activity, independent of the commercial unit, is an important diagnostic tool in determining the factors impacting unit performance. Without separating the catalyst performance from feedstock changes or operating changes it is difficult, if not impossible, to troubleshoot operating problems. In addition to catalyst activity, the importance of coke selectivity for various catalysts has been discussed in great detail. To help quantify the coke selectivity of catalysts, Mott popularized the term "Dynamic Activity"(20). It is intended to relate the catalyst performance as determined by the MAT, to the heat balance of the commercial unit. It has become common to compare coke selectivity between two catalysts on this basis.

commercial operating units as well as experimental and new commercia P samples.

226

Figure 1 Schematic for Typical Mat Unit

MAT PROCESS FLOW SCHEME I L

HEATED OIL DELIVERY ZONE

IL DELIVERY TUBE CALIBRATION

ED THERMOCOUPLE CATALYST BED

3ZONEFURNACE GLASSREACTOR

O'C COOLANT

The testing of new or "fresh" catalyst samples is used to support a number of areas. These include R&D for improved catalyst formulations, comparative analysis of commercial samples, and quality control for catalyst manufacturing. The most common application of testing fresh catalysts is in assessing the relative performance of two or more formulations. Typically, this is in support a program for choosing the "best" candidate from a sampling of three or more vendor offerings. Specific performance goals are generally established rior to this comparative testing. Based

is performed as part of the catalyst selection process that most every refiner goes on a regular frequency. The basis on which the MAT results are interpreted

is thro% hi ly individualized. Some laboratories will compare performance at constant conversion while others prefer constant coke comparisons, Still others are most concerned with hydrothermal stability of the catalyst samples. Each approach has its advantages and disadvantages. Reference to these will be brought out in the discussions to follow.

on the MAT results, a candidate is selected trl at best meets the goals. Most often this

227

The evaluation of fresh catalysts normally includes a deactivation step that precedes the actual MAT evaluation. This deactivation typically involves the steaming of a catalyst sample at temperatures ranging from 1000°F to 1700°F for 2 to 24 hours. The primary objective is to deactivate a fresh catalyst such that its performance in the MAT is representative of what is observed when testing a commercially deactivated sample of the same catalyst. In this way, prediction of commercial performance for new catalysts can be made. An additional deactivation parameter that is widely studied is that induced by the presence of metals in the hydrocarbon feedstock, 'I)lpically nickel and vanadium, these contaminants are relatively non-volatile and when deposited on the catalyst, are not removed during catalyst regeneration. Their presence on the catalyst is known to negatively impact both catalyst stability and product selectivities.

Approaches to the deactivation of fresh catalysts and the impact on catalyst performance were reported by Moorehead et a1 (18). Discussion of this as well as comparative testing philosophies will be addressed in a later section.

An alternative approach to evaluating catalysts, without steam deactivation, is to test commercially deactivated (equilibrium catalysts) samples of the catalysts in the MAT. The principal advantage of this approach comes from eliminating the uncertainties assoclated with laboratory deactivation procedures. For cases where the history of catalyst samples is well known, and operating conditions of the commercial unit(s) is not a variable, this approach has merit. However, it is often very difficult to get equilibrium samples that have seen the same commercial operation. When the objective is to evaluate several catalysts from different vendors, getting samples having equal levels of metal contaminants that have also been exposed to similar severity of catalyst regeneration and unit operations is difficult. Making comparisons on the relative performance of these catalysts having such different histories can be misleading. However, testing of equilibrium catalysts in conjunction with the laboratory evaluation of the fresh catalyst sample is a powerful combination.

Accepting that the results from both approaches have strengths and weaknesses, the combination of the two provides the truest picture of catalyst performance.

A third a plication of the MAT is to assess the reactivity and selectivity patterns of

MAT for this purpose will not be discusse in any detail.

A fourth application is to use the MAT for assessing process variables such as temperature or contact time. Given the mechanical differences between a MAT and a commercial unit, this application is limited.

B different R ydrocarbon feedstocks. As im ortant as this is to the refiner, use of the

228

4. YIELD DETERMINATIONS

4.1. General

Demands on the MAT lab involve more than the simple rating of catalyst activities. At the very least, there is sufficient interest in characterizing the coke and hydrogen producing properties of a catalyst to require collection and analysis of the gas and to determine the carbon on the discharged catalyst. Calculation of a weight balance is another reason for obtaining samples for gas and coke analyses. Most MAT laboratories are capable of obtaining weight balances between 95 and 102% in studies using gas oil feedstocks and many have modified the equipment to achieve this performance with resids. A major interest for the MAT is in obtaining product selectivity data because it is recognized that this inexpensive laboratory test can provide good replication of plant yields if suitable chromatographic technology is used with both the liquid and gaseous products.

4.2. Product Collection

In a typical set up, as shown in Figure 1, the reactor outlet is connected directly to a liquid product receiver immersed in a cooling bath. The outlet of the receiver is connected to a gas holder from which water or saturated brine is displaced. Following injection of the feedstock, the MAT reactor is swept with an inert gas (usually nitrogen) for a period of time sufficient to swee all vapors from the reactor and to

the reaction have been collected in three locations: the coke and a small amount of liquid residue are in the reactor; most of the liquid products are in the receiver; and the gaseous products are in the gas holder.

Many variations in technique among laboratories are found. A single liquid product receiver, constructed as described in ASTM D-3709 and cooled in an ice bath, may not be adequate to stop all of the gasoline when high space velocities'and temperatures are employed. Use of sub-zero cooling baths and multiple receivers is common. Changes 111 the design of the receiver to prevent the bubbling of gas through the product liquid have been reported to help minimize spray losses (21). In all cases where sub-zero cooling is employed, a weathering period, with the liquid warmed to 25°C and nitrogen flowing to the gas collector, is recommended for removal of volatile materials that would otherwise be lost during handling. Followins! the gas sweep the liquid receiver(s) is disconnected, sealed and weighed. The volume of gas roduct and flush nitrogen is equal to the volume of liquid displaced from the

must be employed before removing gas samples for anal sis. Catalyst is removed

catalyst to provide preheat and aid in feed dispersion, a separation is made so that catalyst and dispersion material can be analyzed for carbon individually. Liquid holdup in the bottom of the reactor is typically measured by using a tared cotton swab or filter paper and weighing the swab after wipin In some cases the holdup is rinsed

removal. Usually this material is not analyzed but it's weight is added to amount of 650"F+ material found in the liquid product by analysis. Analysis of material recovered by the solvent technique generally confirms this assignment.

transfer all non-condensed material into the gas R older. At this point the products of

gas ho P der. Temperature and pressure readings are made and some method of mixing

from the cooled reactor after the run. If glass beads or he r ices were placed above the

from the reactor with a volatile solvent and t f e weight determined after solvent

229

4.3. Coke

Carbon on spent catalyst and on dispersion material (if used) is determined by an appropriate combustion method. One common method uses a LECO carbon analyzer. Care must be taken in this analysis since typical catalyst-to-oil ratios of 4 or higher cause multiplication of errors when yields are calculated. In calculating coke yields an allowance for 7 to 10 % hydrogen in the coke is appropriate. This is, however, not part of the ASTM procedure (4).

4.4. Liquid Product Yields

Simulated distillation via ASTM method D 2887 is widely used to determine the boiling range distribution of the liquid receiver contents (liquid product is also termed syncrude). When two traps are used, the contents are combined for analysis. Difficulties occur when sub-zero traps are used and the weathering is insufficient to reduce C4 and lighter material to a few percent since the very sharp peaks produced are not measured correctly. In most cases, however, this technique provides a very precise analysis with a standard deviation better than 0.5 wt % for the determination of any boiling range fraction.

In ranking catalysts and feedstocks, no further correction to this analysis is necessary, but for estimation of absolute yields there is need for use of response factors and the use of assumptions before the data are translated into yields. It has been found by careful calibration with distillation cuts from the FCCU that the response factor for the flame ionization detector is lower with fractions boiling above 650°F than with lower boiling fractions, Interestingly, this is not the case with fractions of virgin material. For accurate yields it is necessary to determine the appropriate response factors.

At the other end of the boiling range, there are unresolved C3 and C4 components dissolved in the collected liquid as well as considerable C5 + material left as vapor in the gas sample. For analysis of the C3 and C4 component yields, it is necessary to estimate or se arately determine the composition of the unresolved "C4 minus" fraction from ti! e simulated distillation to increment the amount of these components found in the gas. Likewise, the amount of "C5 plus" found in the gas must be added to the C5 plus material in the liquid. This, too, requires the estimation or determination of the composition of an unresolved peak to assign the correct response factor.

5. CATALYST DEACTIVATION PROCEDURES-STEAMING

Steaming is used to artificially deactivate a fresh catalyst sample, such that it represents a typical "equilibrium" sample. The choice of steaming conditions determines the physical and chemical characteristics of the "real" catalyst. Therefore, under constant MAT conditions, steaming conditions are responsible for the observed activity and selectivity.

230

Laboratory steaming of fresh FCC catalysts is generally done in the presence of 100 percent steam in a fluidized bed configuration. Catalysts are usually loaded at ambient temperature. In the presence of fluidizing nitrogen the temperature is increased to the desired target. Steam, obtained by vaporization of injected water, is introduced and the nitrogen flow is stopped. After a specified eriod of time, the water injection is stopped and the nitrogen is introduced again an dP the temperature is set back to an ambient or low level. Having reached the desired temperature the catalyst is unloaded and may be screened to remove fines. Alternatively, the catalyst can be introduced into a hot steam environment as opposed to the more gentle temperature ramp identified. The rapid addition of the catalyst to a hot reactor is referred to as a shock steaming.

As simple as this procedure may appear, the methods used to achieve this are varied. There are five approaches commonly used for deactivating a fresh catalyst sample. First, a number of laboratories use a fixed time and vary temperature to achieve a range of deactivated samples which, when evaluated in a MAT unit, will have a range of conversions so that they can make an assessment of catalyst stability and selectivity, Temperature ranges are typically 1300°F to 1600"F, for 3-17 hours. A plot of steaming temperature versus MAT conversion is commonly referred to as a hydrothermal stability curve. Such a plot is one method of measuring the steam stability, or hydrothermal stability, of a catalyst sample. Figure 2 presents typical hydrothermal stability curves for REY and USY catalysts.

Figure 2 Hydrothermal Stability Curve-Constant Time, Variable Temperature.

Typical Hydrothermal Deactivation Curve Constant Time, Variable Temperature

1350 1400 1450 1500 1550 1600

Steaming Temperature, OF

23 1

A second approach for steaming uses a fixed temperature, but the time is varied to generate a second type of hydrothermal stability curve. Temperatures in the range of 1400-1500°F are generally used with times ranging from 5-60 hours. Preferred times, however, tend to be 4-24 hours. The times employed can be tied to either a target conversion or some physical property such as surface area or unit cell size. The deactivated samples are then evaluated in a MAT unit under a standard set of conditions. Figure 3 presents typical hydrothermal stability curves for the same two catalysts from Figure 2.

A third approach involves steaming catalysts for a constant time and temperature, independent of catalyst type. Typical temperatures are in the ran e of 1300-1500°F for 4-17 hours. This method is the preferred protocol as reported 1 y McElhiney (22). This method does not provide for any data on the hydrothermal stability of a catalyst. It does, however, result in equilibration of the unit cell size for octane catalysts.

Figure 3 Hydrothermal Stability Curve-Variable Time, Constant Temperature.

A fourth approach involves a variation from variable temperature/constant time by blending hfferent ratio’s of deactivated samples to represent the inhomogeneity of commercially deactivated samples (23). This method does not appear to be used to any extent at this time.

232

A fifth approach involves the blending of 5% fresh (unsteamed) catalyst to a batch of steamed catalyst (24). Application of this method is limited to the case where fresh catalyst properties are uniquely designed to impact unit performance. If conventional deactivation procedures were to be used for such formulations, the expected commercial benefits would not be observed in the laboratory. As new formulations become commercially available, such non-traditional test methods may have to be developed so as to accurately predict field performance.

The first three approaches are the most commonly practiced. The choice of which method is preferred is highly subjective. The fact that no one method is correct is a measure of the complexity of the task.

For all of these methods, the goal is to produce a catalyst sample that has chemical, physical, and catalytic properties that are indicative of what will be observed in a commercial unit. For a catalyst that has commercial experience, knowing what to expect from the laboratory evaluation makes it easier to determine what steaming method is best. In fact, many steaming methods were developed by varying conditions until the lab deactivated sample provided physical and catalytic properties that matched an equilibrium sample of the same catalyst. However, when testing a new catalyst that has not been used commercially, it is not always clear as to what deactivation procedure is best. For example, if the goal of steaming is to target only the unit cell size, then it might be concluded that one steaming severity is needed. As suggested by McElhiney (22), this would be 1500°F for 5 hours. This approach assumes all catalysts will deactivate to the same extent at these conditions. While the recommended procedure will result in equilibration of the unit cell size (UCS), it does not account for expected changes in MAT activity, zeolite content or total surface area. Figure 4 shows that an equilibrated Unit Cell Size (UCS) for a zero rare earth catalyst (USY) can be obtained at relatively mild steaming conditions; but as presented in Figure 5, the MAT activity and surface areas will continue to change with steaming. As the differences between catalysts become greater, the need to be aware of these other parameters becomes more important. By way of example, Figure 6 shows the effect of steaming severity on zeolitic surface area for an REY and USY catalyst. Also identified are typical values for equilibrium catalysts. What is seen is that the conditions needed to deactivate REY are different than for USY. If the USY is deactivated using the preferred conditions for the REY, then activity and surface areas are not in line with commercial experience. If the reverse is true, then REY is deactivated too severely.

The need to have more than one steaming procedure for different catalysts was recognized by Magee and Blazek (25). These authors stress the need to develop a steaming procedure that accurately reflects the age distribution that is present in commercial equilibrium catalysts. In addition, they correctly point out ”.., that different catalyst types (clay-based, synthetic, semisynthetic, etc.) may deactivate differently in the same commercial unit under identical operating conditions.” The difficulty, of course, is that it is not practical to have a unique steaming procedure for every catalyst. However, it is practical to target steaming severity such that the steamed properties for each type of catalyst are representative of what will be observed commercially.

233

Figure 4 Dependence of Unit Cell Size Equilibration on Steaming Severity

74 3 I I I I

1

- 0 4 8 1 2 1 6

Steaming Time (Hours)

REY @ 1450 F REY @ 1500 F USY @ 1450 F USY @ 1500 F - --&.--- ..... 0 ..... -+-.

Reproduced with permission from Moorehead, E.L.; Margolis, M.J.; and McLean, J. In Characterization And Catalyst Development, An Interactive Approach, Ed. Bradley,S.A, M.J. Gattuso and R.J. Bertolacini; ACS Symposium Series 411,120-134,1989. Copyright 1989 American Chemical Society.

234

Figure 5 Effect of Steaming Severity on MAT Conversion and Zeolitic Surface Area.

Equilibrium ZSA Not Achieved For USY and REY

~ ~~ ~~ " 0 2 4 6 8 10 12 14 16

Steaming Time (Hours at 1450 F) REY Catalyst USY Catalyst + --&--.

Reproduced with permission from Moorehead, E.L.; Margolis, M.J.; and McLean, J. In Characterization And Catalyst Development, An Interactive Approach, Ed. Bradley,S.A., M.J. Gattuso and R.J. Bertolacini; ACS Symposium Series 411,120-134, 1989. Copyright 1989 American Chemical Society.

235

Figure 6 Comparison of Zeolitic Surface Area of Fresh Steamed and Equilibrium Catalysts

MAT Conversion and Zeolitic Surface Area (ZSA)

85 A

E 80 c,

3 c 75 I

0

$ 70 C

6 65

I- 2 60

.-

Continue to Decline for USY Catalyst 125

120

115 E 0)

110 ;3 z 105 4.

v)

100

95

90 55 0 4 8 12 16

Steaming Time (Hours at 1450 F) MAT Conversion ZSA - --&--.

Reproduced with permission from Moorehead, E.L.; Margolis, M.J.; and McLean, J. In Characterization And Catalyst Development, An Interactive Approach, Ed. Bradley,S.A, M.J. Gattuso and RJ. Bertolacini; ACS Symposium Senes 411,120-134,1989. Copyright 1989 American Chemical Society.

236

6. COMPARATIVE TESTING PHILOSOPHIES

In a recent paper, a study was presented that surveyed the fresh catalyst testing philosophies from fifteen companies (18). The objective was to understand the merits of various deactivation and MAT procedures. One conclusion from this study was that each laboratory has a unique testing program, with major differences being practiced in both steam deactivation and MAT methodology, including hardware.

6.1. Steaming Philosophies

Steaming severities were found to range from 1350°F for 17 hours to 1600°F for 4 hours. Some prefer slow heating of the catalyst while others practice "shock" treatment. Some treat all catalysts the same independent of application, while others steam deactivate to a constant conversion or other measurable physical property. Others pre-calcine the sample prior to steaming. A summary of steaming procedures that are generally employed is presented in Table 2.

Table 2 Summary of Steaming Conditions

STEAMING STEAMING PRECALCINATION,'F TEMMF'ERATURE "F TIME, HRS.

None None 1100/1 hr 1100/1 hr

11OO/4 hrs None 1200/3 hrs None 11 12/4 hrs 1 112/3 hrs 1300/1 hr 1OOO/1 hr None 1OOO/1 hr None

*Shock Addition Method

1292-1562 1300- 16W* 1350-1454 1350-1550*

1360-1430 metals 1375

1400 @ 15 psig 1400 & 1500* 1400 & 1500

1418 1425 *

1475 * 1475 1500

1382-1490

1430-1525

5 4 17 4

4.75 17

5 & 10 5 5 15 4

5-80/20% sla 5

6 & 6 4

Adapted with permission from Moorehead, E. L.; Margolis, M. J.; and McLean, J. In Characterization And Catalyst Development, An Interactive Approach, Ed. Bradley, S. A, M. J. Gattuso and R. J. Bertolacini; ACS Symposium Senes 411,120-134, 1989. Copyright 1989 American Chemical Society.

237

All of the approaches have developed over time. The development of the various approaches was based on comparison of laboratory results with testing of equilibrium samples taken from commercial units. Most of these studies were conducted when the most common catalyst type was a high rare earth containing catalyst that may or may not have been calcined during the manufacturing process. As catalyst formulations and market demands have changed, steaming protocols in general have not.

6.2. MAT Philosophies

For the MAT, the same wide range of approaches exist. While an ASTM MAT does exist, no laboratory was found to practice it in total. MAT temperatures vary from 900 to 1000"F, with catalystloil ratios of 3 to 6 and WHSV's of 10 to 40. Some laboratories test all catalysts at constant conversion by either adjusting steaming severity or maintaining steaming severity while varying C/O in the MAT unit. Each laboratory has developed individualized steaming and MAT testing procedures that best suit their needs. Like steaming, where three typical approaches were used there are an equal number for MAT testing:

1) constant temperature and cat/oil--vary conversion by varying steaming severity, 2) constant temperature and vary catloil (injection time held constant), and 3) constant temperature and vary cat/oil (injection time can be varied to maintain

constant space velocity or constant injection time with variable space velocity).

Variations in cat/oil can be obtained by adjusting the oil weight or catalyst weight. Most often it is usually performed by varyin4 the oil weight as opposed to catalyst weight. If the C/O ratio is varied but the injection time is fixed, then the space velocity is changed. Alternatively, the injection time can be varied to maintain a constant space velocity. The former approach is most common. Using a severity relationship described by Wollaston, et. al. that relates severity to catloil and WHSV, the greatest change in reactor severity is obtained with this methodology (26). It should be noted however, that the impact of space velocity in MAT studies is less than that reported by Wollaston. Specifically, the severity factor reported was a function of [(C/O) ** 0.651 * [WHSV * * -0.351. In MAT testing however it has been observed that the severity factor is better described by [(C/O) * * 0.85]* [WHSV * * -0.151 (21).

A summary of MAT procedures used within the industry is presented in Table 3. One unique observation for MAT testing is that testing philosophies in Europe tend to be different from those in the United States. In Europe, MAT testing is characterized by short injection times, less that 25 seconds, giving rise to higher space velocities, usually greater than 30 Hr-1.

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Table 3 Summary of Mat Procedures

MAT Delivery Temperature ‘F Cat/Oil Ratio Time, Sec. WHSV Hr-1

850 900 900 900 915 925 950 950 950 950

950 - 1022 975 985 986

2 3.0

Vary 2.79 3.0

1.875 5-9

2.5 - 5.5 4.0 4.5 4.5

1.5 - 4.5 3.3 6.0

300 75 75 94

N/A 75 35

45 - 75 18 40 40 60 75 20

6 16

Vary 13.7 17 25

11 to 21 13.5-15.5

50 20 > 40

12.8-38.5 14.5 30

Adapted with permission from Moorehead, E. L.; Margolis, M. J.; and McLean, J. In Characterization And Catalyst Development, An Interactive Approach, Ed. Bradley, S. A, M. J. Gattuso and R. J. Bertolacini; ACS Symposium Series 411, 120-134, 1989. Copyright 1989 American Chemical Society.

7. COMPARISON OF STEAMING/MAT PROCEDURES

The results from the comparative steaming/MAT study revealed that only measurement of coke selectivity was method dependent (18). More importantly, it is the effect of steaming, as opposed to the MAT conditions, that impacts coke selectivity. Depending upon the steaming severity, the coke selectivity rankings for two different catalysts (USY and REY) could be reversed as presented in Figure 7.

For clarity,a high coke selectivity catalyst is taken to be one that produces a high level of coke per unit of activity. This is also termed the specific coke. The reciprocal of this is referred to as the dynamic activity (20).

Perhaps not surprising is the fact that the test approach that gave the greatest difference in coke selectivity was that which used the lowest C/O’S and mildest steaming. This results in the greatest difference in unit cell size between a high and low rare earth containing catalyst (see Fi ure 4). This is supported by the work of

zeolite. Rajagopalan and Peters (27) which relate b! coke selectivity to the unit cell size of the

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Figure 7 Dependence of Coke Yield on Steaming Severity

Coke Yields From the MAT Are Dependent Upon Steaming Severity

Steamed Steamed Steamed 1350-1 550 F/4 hrs 1500 F/4 hrs 1400 F/4 hrs

5 Cat/Oil High Cat/Oil Low CatlOil

Variable Temp High Severity Low Severity

Steaming Method

REY Catalyst USY Catalyst

Measured coke selectivities could be reversed if a severe steaming protocol with adjustments made with C/O to vary conversion. Based on Rajagopalan (27), the greatest effects of UCS on coke selectivity are observed within the range of 24.33 and 24.57. Given the small difference in UCS reported for this method of evaluation, it is probably not surprising that the coke selectivities are influenced by not only UCS, but perhaps surface area and C/O in the MAT which were increased to achieve the targeted 70 % conversion.

In another study, McElhiney reported the need to steam fresh catalysts severely to ensure that the UCS of the zeolite had been stabilized (22). Specifically,

the enouP c oice of one steaming procedure was recommended for all catalysts; 1500°F for 5 hours. In cases where an evaluation of similar catalysts is to be conducted, a common steaming procedure like that recommended by McElhiney may be very appropriate. For cases where a variety of catalyst types are to be evaluated, however, this approach may be too generalized.

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It is generally observed commercially that octane catalysts reduce the observed regenerator temperature of an FCCU. Many attribute this to the lower coke selectivity of an octane catalyst. The results from the comparative study show that to see this in the laboratory, the use of a steaming procedure that maximizes the difference in UCS between an octane catalyst and a gasoline catalyst is preferred (18). This tends to be a milder steaming then recommended by McElhiney. The more severe steaming method recommended tends to show that the octane catalyst has high coke selectivity (makes more coke/conversion) than a gasoline catalyst. Commercial experience also reveals that the equilibrium activity of a gasoline catalyst is generally higher than an octane catalyst by up to 10 MAT numbers (28). Table 4 presents comparative selectivity data for a gasoline (REY) catalyst and an octane (USY) catalyst for the case where the "equilibrium" activity of REY is greater than USY.

Table 4 Typical Equilibrium Comparison High Activity REY vs. Low Activity USY

Weinht % Yields @ 70% Conv. REY USY

Dry Gas LJG Gasoline LCO Bottoms Coke

Catalyst Activity Catfoil

1.46 11.77 52.89 18.34 11.66 3.89

76.9 2.9

1.81 13.56 50.89 18.62 11.38 3.73

63.1 7.1

Adapted with permission from Moorehead, E. L.; Margolis, M. J.; and McLean, J. In Characterization And Catalyst Development, An Interactive Approach, Ed. Bradley, S. A, M. J. Gattuso and R. J. Bertolacini; ACS Symposium Series 411, 120-134, 1989. Copyright 1989 American Chemical Society.

Using this approach, the coke selectivity for the USY catalyst is slightly less than REY, as expected from commercial experience. The selectivities for dry gas to bottoms are also in good agreement with commercial experience. Another interpretation for the observed reduction in regenerator temperature of the commercial unit is that the impact of delta coke (coke/Cat/Oil) is greater than is shown by the difference in intrinsic coke selectivity of the catalysts (24). This falls out of the heat balance of the FCC unit which demands that as catalyst circulation increases (increased cat/oil at constant feed rate), the regenerator temperature must decrease (29-31). This is true only when the reactor temperature is held constant. At constant conversion, when the only change is catalyst, reactor temperatures will be nearly equivalent in a commercial unit. Testing catalysts in the MAT, on the basis of constant delta coke has been proposed by Mauleon et al. (28). When similar catalyst

24 1

types are being evaluated, the test philosophy will have less impact on the relative performance then when different catalyst types are being evaluated. This has implication when formulations are developed that require modified test procedures, or when a test method that was developed for gasoline catalysts is now to be used for octane catalysts. As a result, it is important that when making assessments of the catalytic performance of FCC catalysts, the benefits and debits of various methods be reviewed.

8. METALS TESTING

Most catalyst evaluations performed in the United States exclude the impact that contaminant metals (nickel and vanadium) have on catalyst stability and product selectivities. In large part this results from most United States operations being focused on gas oil cracking as opposed to residual operations. Testing with metals is considerably more complicated and expensive and it is not considered to be justified for when the objective is to simulate a low metals commercial operation. For gas oil cracking, contaminant metals are low enough to be virtually ignored. However, not all commercial cracking operations are immune from the impact of metal contaminants. Equilibrium catalysts having metals levels as high as 10,OOO wppm while rare, are possible. More common, levels are 1,OOO to 5,000 wppm. In Europe and the Far East, metals testing is quite standard. Simulation of the metals effects in the laboratory is growing area of concern within many domestic laboratories.

The MAT procedure employed for such testing is generally equivalent to that em loyed for metals free evaluations. The most notable problem that needs to be adgessed is that the feedstocks used for many metal tolerance studies tend to produce much higher yields of coke. The impact this has on material balance can be significant. Also, with a measurable fraction of the feed boiling above standard reactor temperatures, problems of feed vaporization are more pronounced. The need to run higher reactor temperatures to ensure preferred complete vaporization is one factor that has led to higher MAT temperatures in Europe. The recently developed "Microscale Simulation Test" or MST, by AKZO is an example of the test improvements being developed for heavier feedstocks (32). As important as MAT conditions are for this application, the more significant difference in testing arises from the deactivation procedure that is selected.

Procedures for incorporating nickel and vanadium onto a sample of FCC catalysts have been reported as early as 1957 (33). At this time the application was targeted for amorphous silica alumina cracking catalysts. Typically the methods employed involved introduction of a fuel oil or gas oil feed containing nickel and vanadium napthenates to a fluidized bed of fresh catalysts at elevated temperatures. Refinements in this method involved the additional cyclic aging of the metal laden catalysts with oil and air, sequentially to simulate the cracking-regeneration steps. Results from this method compared well with that found for equilibrium catalysts. The major problem was the time that it took for preparing each catalyst to be tested.

242

A simpler, much faster method for preparing metal contaminated catalysts was introduced by B. R. Mitchell in 1980 (34). Now referred to as the "Mitchell Method", this procedure involves the impregnation, using incipient wetness, of nickel and vanadium napthenates on to a calcined catalyst which was then steamed at 1350°F for times as long as 10 hours. Today this is probably the most common method employed given its relative ease.

A subtle, but important variation in this approach gaining in popularity is to evaluate the nickel and vanadium tolerance of catalysts independently. Nickel or vanadium is impregnated onto the catalyst often at more than one level, so as to provide a measure of catalyst erformance as a function of contaminant level. The advantage

loss in catalyst activity associated with moderate to high levels of vanadium. Given that vanadium has more of an impact on conversion, as opposed to selectivity, interpretation of the results is more straightforward. This methodolgy is more common in Europe than in the United States.

of this approach is t K at evaluation of nickel contamination is not complicated by the

The importance of steaming temperature used for deactivating metal impregnated catalyst samples was discussed by Speronello (35). The generally recommended procedure called for a steaming temperature that was approximately 100-150'F higher than the commercially observed regenerator temperature, for typically 4 hours.

As popular as the Mitchell Method is, it has long been recognized to have some key debits. The most striking is that the reactivity of the metals (nickel and vanadium) are greater than what is observed with equilibrium catalysts having equivalent nickel and vanadium levels. It has become somewhat common for laboratories to develop a general correlation that for a given equilibrium metals level, the equivalent level from pore volume impregnation is lower to achieve similar metal effects on the catalyst. This is true for both activity and selectivity determinations. Impregnation of an FCC catalyst to a metals level of 1/3 that observed commercially is common.

A second weakness in this method is that the distribution or profile of the metals within a catalyst microsphere is not the same as what is observed commercially (36). This is most pronounced for nickel, where it has been reported that commercially nickel is deposited on the outer shell of the microsphere (37). The nickel profile resulting from the Mitchell Method is uniformly deposited throughout the microsphere. Vanadium profiles are generally more uniform owing to the relative high mobility of vanadium in the presence of steam at elevated temperatures (38).

Improved test methods aimed at overcoming the debits of the Mitchell Method have been under development for some time. A common theme that they all have is to introduce the metals onto the catalyst in such a manner so as to simulate the commercial realities. The common approach is to crack a feedstock that is enriched with nickel and vanadium napthenates in the presence of the test catalyst. This is followed by a coke bum and perhaps a mild steaming. This is typically done in a self contained test unit, commonly referred to as a fixed fluid bed. This crackinghegeneration cycle is repeated any number of times until the desired metals loading on the catalyst is achieved. It is common that 20-100 cycles, each taking upwards of one hour, are required for each catalyst.

243

In a recent paper, Gemtsen et al. presented a comparison of Mitchell Method deactivation to that of a cyclic aging using an automated fixed fluid bed test unit (39). The reported advantages for the cyclic aging are seen in the metals profile (primarily nickel) through a catalyst microsphere. The impact that such metals distribution has on measured catalyst performance revealed that relative rankings between different catalyst formulations may be reversed, depending upon the deactivation method selected (40). Addition of fresh catalyst during the cyclic deactivation is another feature that aids in simulating the age distribution present in commercial environments.

As good as this cyclic deactivation methodology is, the trade off between accuracy of the data and the time required to prepare each sample will be an important factor that each laboratory will have to access. The development of an improved Mitchell Method type procedure, resulting in the enrichment of nickel on the surface of the microsphere, should represent an attractive compromise.

9. ANALYTICALOPTIONS

There have been a number of important advancements made in the field of gas chromatography during the past fifteen years that are having an impact on the quality and quantity of data that can be derived from MAT evaluations. Key among these is the ability to quantitatively characterize the molecular composition of gasoline. As significant as this is, it is critical to point out that the reproducibility from these advanced analytical techniques may be greater than the ability to reproduce the same gasoline from repetitive MAT runs. As as result, variability in the analysis of MAT syncrudes is largely not the problem of the technique but rather the preparation of the sample to be analyzed. A key objective for improved evaluation for FCC catalysts is to reduce the variability in syncrude composition from replicate runs.

9.1. GC-Octane

Improvements in chromatographic techniques have had a major impact on the the data that are now available from MAT studies. Multi-dimensional chromatography, using both column switching and temperature programming has been successful in separating the gasoline from heavier products to avoid fouling the capillary column (41). Chromatographic techniques are routinely employed to estimate the octane number of the gasoline produced in the MAT unit. Through use of a capillary column it is possible to resolve many of the individual components of the gasoline fraction. Most commonly these are lumped into 31 component groups defined by Anderson, Sharkey and Walsh and each group is assigned its own ”blending octane number” for use in a linear blending model (42).

Using the published linear regression co-efficients, the correlation of results with FCC gasolines having known octane values is not accurate, even after correcting for changes due to use of a different chromatographic column. Considerable independent calibration has been applied by many users. In particular, the same method is used to predict both MON and RON although only the latter were attempted by the method originators. Excellent repeatability is obtained with the GC

244

method so that statistical "noise" is not a problem in discerning differences between samples. Over a wide range of samples there remain inaccuracies that are inherent in the assumption of linear blending and in the variation among unidentified components in the "in-between" groups and may result in a bias between GC and engine values greater than 1 octane number. With this caveat, however, the method has great value in identifying differences between catalysts and feedstocks and is a significant adjunct to the MAT operation.

9.2. PIONA and PIANO

With the advent of improved analytical tools for characterizing hydrocarbons, two new chromatographic techniques have been developed for the molecular characterization of gasolines. The acronyms PIONA and PIANO refer to two different approaches to determining n-Paraffins, I-paraffins, Olefins, Naphthenes and Aromatics in gasoline by gas chromatography. For applications to samples produced in the MAT unit a precolumn is used to separate higher boiling materials before the gasoline peak is analyzed.

Elaborate multi-dimensional analysis is the cornerstone of the PIONA method with traps and multiple columns supplying the required separation of aromatics, olefins and saturates (43). Two ovens are needed to provided an isothermal environment for the group separations while also using temperature programming for separation of paraffins and naphthenes by carbon number. High resolution capillary chromatography is the basis of the PIANO method, which attempts to identify all individual components in the gasoline. The software library and control functions used in this method are extensive and the chromatograph must be precisely programed to obtain correct identifications. In most cases all but 2% of an FCC gasoline can be measured as individual components, then "lumped" in any manner selected by the user. Both methods have distinct advantages and disadvantages and each has staunch advocates. PIONA is fast, offers direct quantitation of each hydrocarbon type by carbon number and is not troubled by co-elution of components. Disadvantages admitted include its considerable expense and the maintenance problems associated with the use of switching valves and traps (44). A particular disadvantage, when applied to MAT products is the very limited life of the olefin trap with materials containing major amounts of olefins. Recent information from users, however, suggests that this problem may have been resolved (40).

The PIANO method has been continuously improved in recent years by use of 100 meter capillary columns and multiple step temperature programming. With the present state of the art all olefins in typical FCC gasoline are individually quantitated but there remain some uncertainties in separation and identification with the higher boiling isoparaffins and naphthenes, leaving a typical uncertainty in classification of 2 wt%. The equipment is more rugged and less expensive than the equipment used with PIONA A benefit of this technique is the availability of information about individual components, making it possible to trace effects of catalysts and feedstocks at the molecular level. Co-elution has not been eliminated, with a superposition of toluene and trimethylpentane cited as a significant problem in finished gasolines (45).

245

In examining the molecular composition of gasolines produced in the MAT versus pilot plant or commercial unit, it has been observed that significant differences exist (32, 46). In particular the olefin content of the MAT gasolines is significantly lower than what is observed commercially (46).

For the most part, traditional testing approaches focus on yields, as opposed to product quality. Future improvements in MAT evaluations may well be tied to selecting conditions that yield a gasoline that has nearly equivalent chemical composition to that observed commercially. This may be difficult however, given the trade-offs in composition as a function of operating conditions that have been reported (46).

10. MAT VERSUS COMMERCIAL UNIT

Despite the remarkable agreement that can be obtained between a MAT and a commercial FCC unit, there are some fundamental differences which prevent even the most highly developed MAT test from duplicating a commercial operation. Some of these are discussed in detail below. As a result, no MAT test will ever be a "mini- FCC unit". But MAT testing is quite useful in generating relative activity, selectivity, and product quality information which, when properly interpreted and used, can accurately predict commercial behavior. The accurate interpretation is a formidable challenge, and differences of o inion among researchers and organizations exist as to

due to the test scale and quantities of products available, while others are more fundamental.

10.1. Reactor Design

A standard MAT reactor uses a fixed bed of catalyst with a crossflow of feed (Figure 1). It may be either cylindrical (ASTM design) or annular AKZO design (32). Versions have been designed to operate with a fluidized catalyst bed. Regardless of specific design, it is quite different from a commercial FCC riser, which operates in a dilute phase, plug flow mode.

In a commercial riser, fresh feed and regenerated (active) catalyst are brought together in the bottom of the riser, with spent catalyst separated from cracked products at the riser outlet. Contact times are the same for both oil and catalyst (if slip is neglected), and the temperature profile depends on many factors such as feed preheat, re enerator temperature, catalyst type (heat of cracking effects), etc.

reactor temperature and vaporization oE the feed. Steam is used for dilution and to provide mixing and transport of the catalyst and oil. The cat/oil weight ratio is essentially the same at all points in the riser. As the catalyst deactivates due to coke deposition, it contacts only partially or fully cracked products in the upper portion of the riser. Generally all catalyst particles see the same oil exposure, and uniform deactivation results (except for the effects of age and activity distribution). The unit operates in a continuous, steady state. Riser contact times of 1 to 5 seconds are typical, again for both catalyst and oil. Pressure drops in the riser itself are quite low due to the dilute phase system.

the ultimate capability and uti P ity of MAT scale testing. Some of the deficiencies are

Generally t t e hot regenerated catalyst is used to accomplish the final heating to

246

A MAT reactor has many different features. Rather than a plug flow of oil and catalyst together, it uses a crossflow configuration. The actual instantaneous cat/oil weight ratio contained together at any point in the run is very high, on the order of several hundred depending on specific unit design and conditions. The operation is non-steady state, so that the results obtained represent time-averaged values. Overall cat/oil ratios in the same range of commercial operation, typically 3-10, are obtained by extending the feed delivery time so that the oil inventory in the reactor turns over many times through the duration of the test. Thus the contact times for catalyst and oil are quite different. As the catalyst deactivates due to coke deposition, it is continually exposed to more fresh feed. There is typically a coke deposition profile through the bedi with the catalyst at the top coking to a greater extent. The MAT is designed to operate with an isothermal temperature profile, although some temperature drop typically is noted due to incomplete feed preheat or endothermic cracking effects. Steam is normally used, although purge nitrogen may be used during feed injection and is always used in post-reaction stripping. The deadman assembly is designed to fully preheat, vaporize, and deliver the feed at reaction temperature. Pressure drops can be significant, particularly for the annular bed design (longer linear distance through the bed) at high space velocities. Fluidized bed designs may improve on the coke deposition profile and pressure drop effects, but otherwise have the same characteristics as the fixed bed designs. Each of these noted differences can have an impact on the comparative results, as discussed individually below.

10.2. Contact Time

The typically designated contact time (e.g., 75 seconds for the ASTM test) refers to the feed injection time, or the catalyst contact time with oil. What is not generally understood is that the oil vapor residence time is much shorter, less than 1 second for the ASTM test. Many authors have written about the inaccuracies due to "long contact times" in MAT testing (46), and numerous "short contact time" tests have been developed (32). While the shorter time tests may indeed have some advantages for certain selectivity predictions, it is not because they come closer to matching commercial oil contact times. On the contrary, they are farther away, as illustrated by the comparison in Table 5. The reason why shorter time tests change selectivities has more to do with the time-averaging effects in MAT testing than with matching commercial conditions (46).

Table 5 Comparison of MAT Contact Times

Catalyst Contact Times, sec. 48 15 ~~

Reactor Temperature,"F Feed Injection Time, sec. Catalyst, gms. Oil, gms. WHSV Cat/Oil (instantaneous) Cat/Oil (overall) Oil Residence time, sec.

910 980 48 15 6 4 1.2 1

15 370

60 385

5 4 0.7 0.2

247

10.3. Cat/Oil Ratio

Calculated instantaneous cat/oil ratios for the two sets of MAT conditions compared are estimated in Table 5 . These values will change somewhat depending on unit design and experimental conditions, but it is clear that the order of magnitude is quite different from a commercial riser. Overall cat/oil ratios of 3 to 10 are obtained for the time averaged run. The generally close agreement between overall cat/oil and feed conversion for MAT and commercial units is quite remarkable given that they get there in such different fashions.

10.4. Time Averaging Effects

As noted above, measured MAT results are by nature, time averaged for the non- steady state system. Since most selectivity and product property attributes are nonlinear with conversion, the averaged results will therefore be different from the steady state results obtained using the same feed and catalyst at nominally the same conditions of temperature and cat/oil in a continuous riser unit. The early sta es of

hydrogen transfer rates with the very active catalyst. As the catalyst deactrvates, both conversions and selectivities change. In the end of the MAT run, conversions are very low due to the high level of coke on catalyst. Thus, while conversion may average to the desired level for commercial comparison, the predominance of cracked products formed early in the run accentuates selectivity and product quality effects. This is why MAT products tend to be less olefinic than comparable riser unit products (46). The effect of higher space velocity (shorter injection time) in the MAT is to accelerate the coking rate and lessen the mpact of overcracking on time averaged results, which gives better correspondence to certain commercial selectivity parameters (22).

10.5. Feed Preheat and Vaporization

In a commercial unit the feed is heated and vaporized by contact with hot catalyst, while a MAT is designed to fully preheat and vaporize gas oil feeds before catalyst contacting. This can lead to an undesired thermal crackmg effect. This problem is accentuated with resid containing feeds and higher reactor temperatures. Both have become more common with attempts to duplicate commercial feeds and conditions. Coke and gas yields with resid feeds are generally higher in MAT testing (Table 6).

the MAT run feature the highest conversions and lead to overcracking an t high

Table 6 Resid Testing in MAT and Pilot Plant Unit Eauilibrium Catalyst with 4OOO uum total metals

Wt % Yields MAT Pilot Riser

Conversion Coke H2 Gasoline

81.4 79.2 12.0 8.1 0.6 0.3

47.9 54.9

248

10.6. Temperature and Pressure Drop

The usually quoted (furnace) temperature in a MAT is actually the highest temperature experienced, and drops of up to 100°F may be experienced due to incomplete feed vaporization and endothermic heat of reaction. Improved deadman and reactor design may help minimize the temperature drop (17,32). The usually quoted temperature in a commercial unit is the riser outlet, which is the lowest temperature experienced. As with cat/oil contacting, it isn’t possible to truly simulate the commercial operation. Most MAT designs have been developed to approximate isothermal operation as closely as possible.

The pressure drop for an ASTM type MAT is quite low, usually less than 2 PSI (4). Higher space velocity and longer annular beds can increase this up to 7-10 PSI (40). At these levels coke selectivity differences are exaggerated, and misleading effects due to particle size and pore size distributions may result.

Despite all the above differences between a MAT unit and the commercial FCC unit, the performance of this small laboratory apparatus is remarkable. Because of its cost and ease of operation, the application of this test is continually being expanded to yield more information on the perforamnce of individual catalysts and the quality of the products produced.

It PREDICTING COMMERCIAL PERFORMANCE FROM MAT RESULTS

As mentioned previously, the ASTM MAT procedure is applicable for providing a relative activity ranking of commercial equilibrium catalysts. The parallel development of laboratory steam deactivation procedures extended the applicability of the test to activity and stability rankings of fresh catalysts. As instrumental analytical test methods were developed which improved precision and allowed more detailed information to be derived from small quantities of products, the expectations of the MAT scale test have also grown. It is quite typical today for detailed predictions of commercial selectivities and product properties to be derived from a MAT study. The desire to ”match commercial performance” has led to the vast diversity of testing conditions employed by different R&D organizations.

The agreement with commercial results is amazing when one considers the fundamental differences between this fixed-bed reactor and a modem FCCU and between the analysis of a time integrated product mixture and a continuous product stream. Data in Table 7 provides an example of such a comparison using samples of hydrotreated residual feed and of equilibrium catalyst that were taken when the commercial results were obtained.

249

Table 7 Comparison of MAT and Commercial Results

MAT RESULTS COMMERCIAL RESULTS

430" F conversion% wt Gas product % wt Total C3 % wt Total Q % wt

650'F + %wt Coke % wt

C5-430 F % wt 430-650 F % wt

79.8 3.3 6.4 10.3 49.1 14.3 5.9 10.7

79.8 4.7 5.6 9.4 48.5 14.2 5.7 11.9

In order to estimate commercial 'elds from MAT data it is necessary to perform

neither the goal or the practice in most laboratories. Nonetheless, yield responses to key variables such as temperature and catalyst-to-oil ratio and to changes in catalyst formulation are mirrored quite successfully by the MAT unit.

As discussed previously, the two fundamental applications of this technique are screening catalysts and screening feedstocks. Occasionally, there are investigations in which operating conditions are screened. In all of these investigations the goal is to provide a table of comparative data where everything other than the parameter under investigation is constant and the differences in yields reflect the effect of this parameter alone. Typically a series of runs are made at different conversions and data are interpolated to provide the desired comparison. For the most part, comparisons are made at a constant conversion giving a direct read out on selectivity differences. Another popular method for making comparisons is at constant coke yield to reflect the heat balance limitations of the FCCU. Table 8 and 9 provide an example of both methods from a comparison of three equilibrium catalysts.

calibrations with FCCU produce d" materials, as described later, but this is usually

250

Table 8 Comparison Of Catalysts at Constant CONVERSION (75%) After Adjusting to The Same Metal Levels (3,000 ppm nickel equivalent)

Yield, wt% Cat A Cat B Cat C

H2 CH4 c2H4 c2H4 C3H6 C3H8 C a 8 ISOBUTANE N-BUTANE C5-430 F 430-650 F 650-800 O F

800 + O F

COKE

0.5 0.7 0.7 0.6 4.8 0.8 3.3 3.0 0.9

52.9 15.5 6.6 3.0 6.7

0.5 0.9 0.7 0.7 4.3 0.5 2.5 2.5 0.7

53.3 16.7 6.1 2.2 8.4

0.3 0.8 0.7 0.6 4.7 0.5 3.8 3.0 1.0

52.2 15.7 6.3 3.0 7.4

Table 9 Comparison Of Catalysts at Constant COKE YIELD (7%) After Adjusting to The Same Metal Levels (3,000 ppm nickel equivalent)

Yield, wt% Cat A Cat B Cat C

H2 CH4 c2H4 C2H6 C3H6 C3H8 C a 8 ISOBUTANE N-BUT- C5-430 " F 430-650 "F 650-800°F 800+"F COKE Conversion

0.4 0.8 0.8 0.7 5.0 1.0 3.5 3.2 1.0

53.3 14.7 6.0 2.6 7.0

76.8

0.4 0.8 0.6 0.6 4.1 0.6 2.5 2.1 0.7

52.6 18.1 7.2 2.7 7.0

72.0

0.3 0.8 0.7 0.6 4.8 0.5 3.8 3.0 1 .o

52.4 15.8 6.3 3.0 7.0

74.8

Neither constant conversion nor constant coke comparisons are truly representative of commercial performance differences expected however. Conversion is a dependent variable in commercial operation and will change depending on the unit

251

limits and constraints. Constant coke comparisons can be misleading since coke selectivity has been identified as the single most method dependent parameter to be determined in the MAT (18). Also, the deactivation levels required to achieve constant conversion or coke yields will generally vary for catalysts having different activities. One method of dealing with making the catalyst comparisons as accurate as possible is to use a heat balanced commercial projection model which uses MAT results as input. mica1 input parameters include hydrothermal stability data, which are translated into a relationship between catalyst addition rate and equilibrium MAT activity, along with selectivities, GC octane res onse, and metal tolerance parameters. The computer model generally compares res up ts for a base catalyst which generated a known commercial response to a candidate replacement catalyst. Heat balance effects and unit constraints are imposed which are not reflected in the raw MAT data. An example of such a comparison is shown in Tables 11 and 12. A base catalyst was compared to a candidate new catalyst in the MAT. The results were used to generate a prediction of commercial performance for the new catalyst relative to the base catalyst. The MAT data show the new catalyst to have improved coke and dry gas selectivity, slightly oorer bottoms reduction and better metals tolerance over the

projection, which was run assuming a constant wet gas constraint, the improved coke selectivity lowered the regenerator temperature and led to a higher cat/ol ratio. This was combined with a higher reactor temperature, achievable due to improved gas selectivity and a reduced catalyst make-up rate. The latter results from needing a lower equilibrium activity because of the higher cat/oil. The net result was higher conversion, higher gasoline yield and octane, lower bottoms make, and a more favorable split between LPG and dry gas. All at constant wet gas yield. Not all of these effects are immediately obvious from the raw MAT data, indicating the usefulness of combining these two tools.

base. Activity, sta ! ility and Octane potential were the same. In the commercial

Table 10 Comparative MAT Results for Two Catalysts

B W New catalyst Catalyst

MATActivity B W Same Hydrothermal Stability B W Same Metals Factor (coke/gas) B W 0.7 X Base

Yields. wt% (Metals Free) Ca 70 wt% Conversion

H2 CdC2 CdC4 Gasoline LCO Bottoms Coke GC-RON

0.06 1.4 12.3 52.6 18.9 11.1 3.7 90.5

0.03 1.3 12.6 52.8 18.3 11.7 3.3 90.5

252

Table 11 Translation of MAT Results To Projected Commercial Performance

Base New Catalyst Catalyst Actual Projected

Catalyst Make-up (TPD) E-Cat MAT E-Cat Ni E-Cat V

Reactor Temperature Regenerator Temperature CatJOil Conversion, Vol.%

Yields, Dry Gas (C2-), Wt.% Total Wet Gas, MSCFM Total LPG, Vol.% Gasoline, Vol.% LCO, Vol.% Bottoms, Vol.% Coke, Wt.%

Engine RON

2.2 67

1500 3000

Base Base

5.7 70.1

4.9 10.2 27.1 52.7 20.1 9.8 5.8

91.6

2.0 66

1650 3300

Base + 7 Base-15

6.3 71.6

4.7 10.2 29.2 53.2 19.0 9.4 5.9

92.0

12. SUMMARY

The Microactivity Test is the predominant tool used in assessing the catalytic performance of FCC catalysts. Since its inception in 1967, the methodology, apparatus, and applications have undergone many refinements. Where once this test was a semi-quantitative tool for assessing catalyst activity, it is now routinely used to predict commercial yields and product qualities of the syncrude.

Improvements in the testing of new catalysts will continue. The move to increased reaction temperatures and reduced contact times is likely to expand to nearly all laboratories. In short order, a "standard" MAT procedure will have reactor temperatures between 950 and 1000'F, with WHSV's greater than 30 hr-1. Directionally, this represents an incremental improvement in the ability to predict commercial performance.

Similarly, improvements in the laboratory procedures for deactivating fresh catalysts will also evolve. The ability to predict performance of new formulations that are

253

designed to have improved metals tolerance for example, will demand better methods for catalyst deactivation.

Finally, as the era of reformulated gasoline expands the ability to characterize the molecular composition of the product fuels produced in the FCC unit will gain importance. Development of new formulations that control gasoline composition, not just yield, will depend heavily on new analytical procedures.

Use of the MAT to screen all of these formulations will continue to be a critical component for the laboratory evaluation of FCC catalysts.

13.

1

lb 2a

2b 3 4

5 6 7 8 9 10 11 12 13

14 15

16

17 18

19 20 21 22 23

24

REFERENCES

Wachtel, S. J., et al; Preprint, Division of Petroleum Chemistry Amer. Chem. Society 16 (3) Sept. 1971. Humes, W. H. Chemical Engineering Progress, February 1983. Creighton, J. E.; Young, G. W. The Catalysis Society; Eighth North American Meeting, May 1983. Young, G. W.; Davison Catalagram, 80,1990. Rawlence, D. J.; Gosling, K. Applied Catalysis, 43 (1988) 213-237. ASTM Standard D 3907-87, "Standard Method for Testing Fluid Cracking Catalysts by Microactivity Test". Alexander, J. Proc. API 27, I11 51-56, 1947. Birkhimer, E. R., et a1 Ibid 27 111,80-89,1947. McReynolds, H. Ibid 27 111,78-83, 1947. Shankland, R. V. and Schmitkon, G. E., Ibid 27 111,57-77, 1947. Conn, M. E. and Connally, G. C., Ind. Eng. Chem. 39, 1138-1143,1947. Grote, H. W. and Olsen, C. R., Oil and Gas Journal 46 (28) 332,1947. Ivey, F. E. Jr., and Veltman, P. L., Pet Ref 31 (6), 93-10, 1952. Ciapetta, F. G.; Henderson D. S. Oil and Gas Journal 65 (42) 88 (1967). Eggertsen, E. T.; Groennings, S.; and Holst, J. J. Analytical Chemistry 32, 904 (1960). Barras, R. C.; Boyle, J. F. Oil and Gas Journal July 30, 1962. Green, L E.; Schmauch, L. J.; and Worman, J. C., Analytical Chemistry 36,1512- 16 (1964). Petrocelli, J. A; Puzniak, T. J.; and Clark, R. O., Analytical Chemistry, 36, 1008- 11 (1964). Gustafson, W. R. ACS Div. Petrol. Chem. 14 (3) B56-67, 1969. Moorehead, E. L.; Margolis, M. J.; and McLean, J. In Characterization And Catalyst Development, An Interactive Approach, Bradley, S. A et a1 Ed.; ACS Symposium Series 411,120-134,1989. Upson, L L Hydrocarbon Processing, 11,253-258 1981. Mott, R. W. Oil and Gas Journal 1987 85, January 26,73-77. Engelhard Corporation, private communication, (1990). McElhiney, G. Oil and Gas Journal 1988 86, February 8,3538. Keyworth, D. A; Turner, W. J.; Reid, T. A Oil and Gas Journal 1988 86, March

McLean, J.; E. L. Moorehead; Hydrocarbon Processing, February 1991, pg 41. 14,65-68.

254

25

26

27 28 29 30 31

32

33

34 35 36 37 38

39

40 41

42

43 44

45

46

47

Magee, J. S.; Blazek, J. J. In Zeolite Chemistry and Catalysis, Rabo J. A. Ed.; ACS Monograph 171 615 1976. Wollaston, E. G.; Halfin, W. J.; Ford, W. D.; D’Souza, G. J. Hydrocarbon Processin 1975 54 (19), 93.

Mauleon, J. L.; Courcelle, J. C. Oil and Gas Journal 1985 83, October 21,64-70. Upson, L., 3rd Katalistiks Fluid Cat Cracking Symposium, Amsterdam, 1982. Pierce, E., Hydrocarbon Processing, February 1983, pg 39-42. Yen, L. C., Wrench, R. E. and Kuo, C. M., Oil and Gas Journal Sept. 16, 1985 pp

O’Connor, P.; Hartkamp, M. B., In Characterization And Catalyst Development, Bradley, S. A et a1 Ed.; ACS Monog raph 411,135,1989. Conner, J. E., Rothrock, J. J., Birkhimer, E. R. and h u m , L. N., Ind Eng. Chem. 49 276 (1957). Mitchell, B. R., Ind. Eng. Chem. Prod. Res. Dev., 1980,19 pg 209-213. Speronello, B. K. and Reagan, W. J., Oil and Gas, Jan 1984 pg 139-143. Masselli, J. A, Peters, A. W., Catal. Rev. Sci. Eng. 1984 26,525. Nishimura, Y., Ogata, M., Ida, T., Takakura, K. J. Japan Pet. Inst. 1983 26 344. Letta, D. P., Kugler, E. L., In Characterization And Catalyst Development, Bradley, S. A et a1 Ed.; ACS Monograph 411,354, 1989. Gerritsen, L. A, Wijngaards, H. N. J., Verwoert, J., and O’Connor, P., AKZO Catalysts Symposium 1991, Scheveningen, The Netherlands (obtainable from AKZO Chemie). Engelhard Corporation, private communication, (1991). Cronkright, W.A.; Butler, M.M. and Harter, D.A., Ketjen Catalysts Symposium ’86, Scheveningen, The Netherlands (obtainable from AKZO Chemie). Anderson, P.C.; Sharkey, J.M. and Walsh, R.P., J. Inst. of Petrol. 58, (1972), pp

P. Van Arkel, et al, J. Chromatographic Science 25, 141-148 1987. Kosal, N; Bhaei, A. and Ali, M.A, ”Determination of Hydrocarbon Types in Naphthas, Gasolines and Kerosenes: A Review and Comparative Study of Different Analytical Procedures”, FUEL, 69, pp 1012-1019 (1990) Yatsu, C.A and Keyworth, D.A., Oil and Gas Journal March 26, 1990, pp 64-73 1990 Margolis, M. J. and McLean, J. Symposium on Advanced FCC I1 1991 AIChE Annual Meeting, Los hgeles, November 1991. Biswas, J., I. E. Maxwell, Applied Catalysis 63 (1990) 197-258.

Rajagopa f an, K., Peters, A W. J. Catalysis 1987 106,410.

87-92.

83-93

255

Bottoms

Light Cycle Oil (WO)

Gasoline

LPG

Dry Gas

Coke

MAT Conversion

Kinetic Activity

Dynamic Activity

Cat/Oil (C/O)

Injection Time

WHSV

Delta Coke

Deadman

TERMS AND DEFINITIONS

Definition

Feed or Product Oil with boiling point > 650 F

Feed or Product Oil with boiling point > 430

Product Oil with boiling point > 60

C3 and C4 Alkanes/Olefins

Hydrogen, Cl-C2 Hydrocarbons

Carbon deposited on catalyst

lOO-(LCO +Bottoms)

Conversion/( 10-Conversion)

Kinetic Activity/Coke Yield

Weight ratio Catalyst to Oil used in test

Time (seconds) required to inject oil into MAT

Weight Hourly Space Velocity 36OO/(cIO*Inj. Time)

Coke Yield/(Cat/Oil Ratio)

Metal Insert in MAT used to vaporize Oil prior to contacting catalyst


257

CHAPTER 8

REALISTIC ASSESSMENT OF FCC CATALYST PERFORMANCE IN THE LABORATORY

GEORGE W. YOUNG

W. R. Grace & Co.-Conn., Davison Chemical Division Washington Research Center, 7379 Route 32, Columbia, Maryland 21044

1. INTRODUCTION

It is an axiom of catalyst development that the evaluation methods used to define catalyst performance usually determine the direction of catalyst improvements. In some cases, it may be laboratory testing that determines if a catalyst is judged to be a commercial success, since appropriate commercial data may not be useful or available. This places a heavy burden on the laboratories responsible for catalyst or process performance assessments. Inappropriate test methods will lead to non-optimal choices for catalysts or process conditions, resulting in lost opportunities at best and, at worst, misdirection of research and development efforts.

The guiding principal for practical catalyst test methods should be to achieve as realistic a simulation of all aspects of the commercial process as possible. This contrasts sharply with the philosophy of finding testing techniques that will maximize differences between catalyst performance. All too often, these techniques have no commercial relevance and in some cases, can give grossly exaggerated or even contradictory information. Examples of these are the widely used Microactivity (MAT) test (ASTM D3907 or D5512) and some of the methods designed to illustrate metal (Ni, V) tolerance properties of catalysts. Awareness of the differences between the laboratory and the commercial environment is crucial, and continuing efforts must be made to develop tests that predict accurate outcomes in the real world.

It is tempting to try to review the evolution of FCC catalyst evaluation techniques and the sometimes less than noble reasons for the evolution of particular techniques. However, this chapter will concentrate on describing techniques which the author believes can provide reasonable commercial simulation and realistic assessment of the expected performance of FCC catalysts. Almost all major and many of the smaller oil companies, together with catalyst manufacturers and some academic research centers, are involved in the business of evaluating the performance of either the FCC catalyst or the FCC unit with laboratory techniques. In FCC catalysis, it is rare for any two laboratories to perform their testing in the same way, and energetic debates are often encountered as each laboratory staunchly defends its techniques and rejects results obtained any other way. Many of these differences are cosmetic, but they provide political flexibility for rejection of

258

outsider information. On the other hand, some lab methods are completely erroneous, yielding contradictory rankings to those observed commercially. There is still much that needs to be done, not only to close the gap between the laboratory results and real world performance, but also in the education of users and originators of catalyst evaluation data.

2. FCC CATALYST EVALUATION STRATEGIES

Several common types of FCC catalyst studies that are performed in commercial laboratories are summarized in Table 1, together with the typical experimental strategies employed. From the perspective of the catalyst manufacturer, the most important of these types of studies are: 1) catalyst research and development for performance improvements; 2) catalyst selection for a specific commercial FCCU where the operational objectives are known. From the perspective of the catalyst user, the most important studies may be: 1) feed studies; 2) FCCU process optimization; 3) catalyst selection; 4) catalyst additive performance evaluation.

Table 1 FCC catalyst evaluation strategies

l%uUm& GQdS Pretreatment Reactor

Catalyst Screening

Catalyst R and D

Catalyst Selection

Catalyst Additives

Feed Studies

FCCU Optimize

Catalyst optimize

Database of Relative Performances

Performance MAT/Riser/Other Improvements

improve FCCU Operation

Octane, SOX, etc. Riser/MAT/Other

FCCU Impact

FCCU Constraint

Verify Effect of Catalyst Change

Steam MAT/Riser Pilot Plant

Steam

Steam Riser/MAT

Steam Additive

Blend with E-Cat or Steamed Catalyst

E-Cat/Steam Riser/MAT

E-Cat/Steam Riser

E-Cats Rise r/MAT

259

Catalyst obtained from a commercial FCCU is termed “equilibrium catalyst” (E-Cat), and in the best (and rarest) case, may be composed of only a single grade of catalyst. Usually, the circulating catalyst inventory is contaminated with remains of previous catalyst grades, spalled refractory, equilibrium catalyst purchased to either control contaminant metal levels or for start-up after a shut down. Even the same grade of catalyst obtained as an equilibrium catalyst from different commercial FCCU’s will have been subjected to different environments (regenerator temperatures; steam concentration; contaminant metals such as sodium, vanadium and nickel; fresh catalyst replacement, etc.). Consequently, they are unlikely to give identical catalytic performance. Equilibrium catalysts are also characterized by having an “age distribution” (1) which has in turn led some workers to propose the use of mixtures of differently steamed catalysts in an attempt to simulate the equilibrium condition (2, 3).

Because of the uncertainties about their true nature, it is strongly recommended that equilibrium catalysts (especially from different commercial FCCU’s) should only be used when the catalyst performance itself is not the primary focus of the investigation. An exception to this general rule is the situation in which the equilibrium catalysts have been carefully selected to ensure homogeneity of type and comparability of contaminant metals levels, although the ability to blend catalysts of the same type from different units will increase the flexibility of this approach. Nevertheless, equilibrium catalysts should not be used in studies designed to select a particular catalyst grade for a specific FCCU. The probability of obtaining the specific catalysts of interest and with properties that might be expected to occur from use in the target FCCU is virtually zero. In addition, use of only equilibrium catalysts eliminates the possibility of evaluating new or specially customized catalysts. On the other hand, equilibrium catalyst from the target FCCU is essential for determining an appropriate pretreatment procedure for the fresh catalysts considered for its use.

3. A CATALYST EVALUATION ORDEAL!

Table 2 shows the outcome from a typical FCC catalyst evaluation study performed using steamed catalysts and the Microactivity (MAT, e.g., variations of ASTM 03907 and D5512-91). These data were obtained using catalysts that had been steamed for 4 hours at 816°C in 100% steam at atmospheric pressure in a fluidized bed reactor; prior to the introduction of the steam, the catalysts had been slowly heated to 816°C over a period of 3 hours under a nitrogen flow. Each catalyst was tested in the MAT at four different catalyst/oil (C/O) ratios, but at constant catalyst contact time (time on stream) at a nominally constant reactor temperature (527°C). This method produces, for each catalyst, a range of conversions and product selectivities which can be plotted against conversion, coke yield, or C/O, and catalysts can be compared to one another by interpolating the results. Usually, catalyst performance is compared at constant conversion or constant coke yield and typically, the catalyst yielding the most gasoline with the highest octane or some other desirable product slate is declared to be the winner. It is not uncommon for a catalyst purchase decision to be based upon such results, although more than likely a yield estimate, which predicts heat balanced yields for the commercial FCCU, will be determined using the MAT data as a guide to rival catalyst selectivities.

260

For the particular catalysts shown in Table 2, the following conclusions could be reached.

Table 2 Pre-treatment study; steamed 4 hr. 81 6°C; constant conversion comparison

Catalvst A B c D

Conversion 65 c/o 3.9

Weia ht YQ H2 0.05 c1 +c2 2.4

Total C3’s 5.8 c3= 4.7

c4= iC4 Total C4’s

Gasoline LCO Bottoms

Coke

RON MON

5.5 3.1 9.4

45.0 21.5 13.5

2.4

90.2 80.5

Gasol ine ComDos ition P 5.0

I A N 0

35.9 31.3 8.3 19.5

65 5.0

0.06 2.4

4.8 5.7

6.3 2.7 9.7

45.0 22.3 12.7

2.1

91 .o 80.5

5.0 31.7 29.8 8.0 25.5

65 4.5

0.1 0 2.4

4.8 5.7

6.5 2.7 9.8

44.0 22.7 12.3

3.0

91.5 80.5

4.7 30.8 28.7 8.3 27.5

65 3.9

0.05 2.3

4.7 5.7

5.7 3.1 9.5

45.0 21.5 13.5

2.4

90.5 80.5

4.7 35.9 31.3 8.3 19.8

1. Catalysts A and D are the most active catalysts, since they achieved the standard conversion (65 wt% ff) at the lowest C/O ratio. This may be important for a unit that is circulation limited.

2. Catalyst C makes twice as much hydrogen as catalyst A, 8, and D. Since hydrogen has low molecular weight compared to the other light gases, this could present a major problem if the gas compressor for the commercial FCCU is already at capacity.

261

3. All catalysts appear to yield the same amount of C1, Cp’s, and C3’s .

4. Catalysts B and C yield higher levels of butenes, lower levels of butanes, and overall, higher amounts of C4’s than catalyst A or D. This could be desirable for a unit with downstream alkylation capacity, or may be a problem if a unit has a gas compressor limitation.

5. Catalyst C produces less gasoline than the other three.

6. Catalysts B and C make more light cycle oil (LCO) than catalysts A or D, with C showing the best LCO selectivity. Catalyst B offers the highest liquid product selectivity (G+D).

7. Catalyst C has the highest (worst) coke yield and Catalyst B has the best coke selectivity. Usually, lower coke yield will permit a reduction of regenerator temperature, resulting in an increase in catalyst circulation rate and increased conversion.

8. All catalysts produce a gasoline with the same motor octane (determined by a gas chromatographic method known as G-CON@ (4), but research octane ranks the catalyst; C > B > D > A.

9. Catalysts B and C produce a highly olefinic gasoline.

Table 3 shows the results from identical MAT testing of these same four catalysts in which a “minor” modification of the catalyst pretreatment was introduced. The catalysts were steamed at 816°C as before, but the recommendation of Moorehead, et. al. (2) was followed in that 5% of fresh, unsteamed catalyst was intimately blended with the steamed catalyst. As can be seen from a comparison of the two sets of data, there are considerable changes in the ranking of catalysts, as well as the absolute values of the yields. Catalysts A, B, and C are approximately equal in activity; Catalyst A makes the least dry gas; Catalyst B, and especially C, make higher yields of propylene; Catalysts B and D make less gasoline than A, although C still shows the least gasoline; differences in coke selectivity are virtually e Ii mi nated.

Thus, by making what appears to be a modest change in the pretreatment of the catalyst, an entirely new set of conclusions is obtained. As a result, a different catalyst would be declared the winner in a catalyst selection study. That catalyst rankings can change depending on the method of steaming has been pointed out by many workers (e.g., 5, 6). However, this case history illustrates a real problem facing those involved in the art of catalyst evaluation and leaves unanswered the question of which technique, if either, is the more accurate predictor of what will happen in the commercial context. Catalyst pretreatment methods and strategies are discussed elsewhere in this chapter in more detail, together with some discussion on the pros and cons of some of the techniques encountered by the author.

The pretreatment problem is just one of the dilemmas facing the laboratory worker who has been delegated the responsibility of recommending a new catalyst. Table 4 shows how three catalysts were ranked by Sapre (7) for activity and coke selectivity when tested in different reactor types (FFB, MAT, Riser), each using

262

different feeds. (It is common practice for the different types of pilot plants operating in a laboratory to utilize different, “standard” feedstocks). The general conclusion from Sapre’s work is that the choice of the reactor can determine the outcome of the catalyst rankings, and without sophisticated analysis of the results using reaction engineering models, simple laboratory reactors can give incorrect predictions. These are not new conclusions (e.g., Nace and Weekman, a), but as catalyst performance tests are made more elaborate through the use of greatly enhanced analytical techniques, they are often ignored!

Table 3 Pre-treatment: 5% fresh/95% steamed (4 hr. 81 6°C) 65w% conversion comparison

Catalvst A B c D

Conversion 65 c / o 3.5

0.06 !f!huba H2 c1 +c2 2.4

c3= 4.1 Total C3’s 5.6

c4= iC4 Total (24%


5.0 3.5 9.4

44.1 21.6 13.4

Coke 3.4

RON MON

90.2 80.5

Gasoline ComDosition P 4.9 I 35.2

A 31.3 N 8.6 0 20.0

65 4.2

0.07 2.6

4.5 6.0

5.7 3.2 9.8

43.2 22.2 12.8

3.3

91.5 80.9

4.7 32.0 31.3 8.0 24.0

65 65 3.5 3.3

0.1 1 0.06 2.7 2.6

4.8 4.3 6.2 5.9

5.9 5.3 3.3 3.5 10.0 9.8

42.4 43.2 22.2 21.6 12.8 13.4

3.6 3.4

91.6 91.2 80.9 80.9

4.7 4.6 32.0 33.9 31.3 31.3 8.0 8.0 24.0 22.2

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4. CATALYST PRETREATMENT PRACTICE

4.1. Catalyst Steamlng In use, an FCC catalyst undergoes several types of deactivation. In the

reactor, carbonaceous residues deposit rapidly and significantly poison the catalyst and alter its selectivities. This rapid poisoning is the distinctive feature of FCC catalysis, and has led to the evolution of unique reactor configurations to manage this problem (9).

Table 4 Catalyst ranking by various reactors (From Sapre, et. al., ACS Meeting, August 1990, Washington, D. C.)

A Conversion. "/9 Fixed Fluid B d w m

Catalyst A (REY) 63 66 70 B (USY) 65 62 64 C (RE-USY) 68 67 67

Selectivu . .

B (Coke yield/Kinetic Conversion)

Catalyst A (REY) B (USY) C (RE-USY)

0.55 1.33 1.77 0.38 0.99 1.20 0.41 1.07 1.33

The carbon is removed by high temperature oxidation in the regenerator, and it is here that the catalyst undergoes three other forms of deactivation. In the regenerator, the catalyst experiences high temperatures and steam is always present. Therefore the zeolite in the FCC catalyst undergoes both dehydroxylation and dealumination, causing the zeolite unit cell to be reduced. Furthermore, the zeolite can undergo crystal destruction or sintering. The non-zeolitic portion of the FCC catalyst, which may also contain active reaction sites, can also undergo sintering, as observed by a loss of surface area and modification of the catalyst's pore size distribution. These effects alter the activity and selectivity of the cracking catalyst. In zeolite catalysts, Moscou and Mone (10) demonstrated that even high temperature calcination, in the absence of steam, will drastically alter the catalyst activity and selectivity. They also showed that when the catalyst is calcined or steam treated, performance actually improves (lower coke and higher gasoline). Equilibrium catalysts (E-Cat) exhibit much better selectivities than fresh, unsteamed catalysts and consequently, pre-steaming zeolite catalysts has been an accepted practice to mimic their properties. There are as many methods of steaming catalysts as there are laboratories, and a few of these have been reviewed by a

264

number of people (1 1, 12). The most common techniques currently in use can be classified as follows: a) high temperature (800-820°C) for short times (2-6 hours); b) low temperatures (730-770°C) for long times (up to 24 hours); c) blending mildly steamed or fresh catalyst with more severely steamed catalyst. Some laboratories steam catalysts to constant activity by varying temperature (at constant steaming time) or vary time at constant temperature. Other laboratories choose to compare catalyst performance after steaming all catalysts under the same conditions.

Steaming to constant activity is based on the premise that a catalyst formulated to be more active (higher zeolite content, higher rare earth, etc.) will either undergo a more severe deactivation or fresh catalyst addition rate will be lowered so that the “E-Cat” activity will remain constant. !n practice, the FCCU E-Cat activity may be at a different level upon change of catalyst type and, unless there are special circumstances, the refiner will usually maintain catalyst addition rate at the previous level. From the evaluation perspective, there are many laboratory pilot plants that have little flexibility to permit them to obtain a yield curve (i.e., operating over a broad range of testing severities to obtain a range of conversions and product yields). In this case, the simplest way to compare catalysts is to use the pretreatment process to make all catalysts have comparable activity. This technique, required because of poorly designed or outdated pilot plant equipment, completely ignores the considerable benefits that might be derived from formulating a catalyst to have a high activity.

A radical example of such a conflict would be a refiner changing from a REY catalyst to one containing a low rare earth level USY for improved coke selectivity. The MAT activity of the catalysts (at the same steaming conditions and probably as equilibrium catalysts) will probably be considerably different, but the FCC unit will respond by altering catalyst circulation rate as a result of a drop in regenerator temperature. A laboratory that pretreats to constant activity will give a milder steaming to the USY catalyst, which may result in too high a unit cell, too much hydrogen transfer activity, etc. This in turn will most likely alter the intrinsic selectivities of the catalyst and hence, the lab evaluation process will have introduced a potentially significant bias.

Another complication arises because the catalyst activity is assessed by the MAT test and, as pointed out earlier, the ranking of catalyst activity can be different from the ranking observed in a riser reactor. Most catalyst suppliers report “Fresh Activity” for their various grades of catalysts. These are MAT values determined after the catalyst has been subjected to a mild steaming (e.g., 6 hours at 760°C). These “fresh” MAT values generally have no relationship to the MAT activity of the equilibrium catalyst, which is highly dependent on the catalyst management practices and the severity of the regeneration environment at the specific FCCU. It is also possible for catalysts with different “fresh activities” to equilibrate in the reverse activity order if the catalysts have different hydrothermal stabilities (e.g., NaY vs. USY, 5).

Another objection to testing catalysts of significantly different activity arises because in the subsequent selectivity tests, catalysts will be compared either at different reaction conditions or at different conversions. In MAT testing, there is a widespread belief that selectivities obtained at low C/O are more favorable (higher gasoline selectivity, lower coke, lower hydrogen) than those obtained at high C/O, hence the most active catalysts can appear to have an advantage. This weakness may be partially overcome by using more realistic deactivation procedures (many of

265

the methods practiced are generally not severe enough to mimic the commercial deactivation and short contact time, high temperature adiabatic reactor designs).

Many laboratories find it most convenient and practical to steam catalysts overnight, using steaming times of 12 to 24 hours. To maintain reasonable catalyst activity, the temperature is usually in the range of 730-790°C. At least one commercial laboratory (13) that is involved in providing FCC catalyst evaluations has advocated this approach claiming that long time, low temperature steaming gives a “more accurate” deactivation of the matrix than does the high temperature, short time steaming that is used by many labs. While some FCCU units may provide a very ‘severe deactivation to the catalyst components, and hence a long time steaming may be required, generally catalyst properties are “point functions”, having characteristics that can be achieved by a variety of methods of different time- temperature-steam pressure combinations. Table 5 illustrates a comparison among several different catalyst types after steaming at either 24 hours at 773.8”C or 4 hours at 816°C and atmospheric pressure with 100% steam. Also, Table 5 shows a catalyst’s properties after it was steamed for 40 hours at 732.2”C, 24 hours at 760°C, and 4 hours at 816°C. Each of these steaming conditions produced a catalyst with equivalent properties.

The more important aspect of catalyst pretreatment is to realistically simulate either a particular FCCU or a “typical” FCCU. It is the author’s experience that a reasonable match of equilibrium catalyst activity and physical properties can be achieved using a 4 to 5 hour steaming at 816% (14). If a particular FCCU is being targeted, then a protocol such as described by Patrose and Young (15) or in ASTM D4463-91 can be followed. This protocol suggests that a deactivation curve (e.g., vary steaming time at constant temperature) be obtained. Then, using the equilibrium catalyst properties as a guide, select the appropriate steaming time to simulate the commercial deactivation severity.

4.2. Complications with Nickel and Vanadium An additional deactivating mechanism for FCC catalysts involves the deposition

of contaminant metals from the feeds. Of these, the worst are usually nickel and vanadium, both of which can have a dehydrogenation role. Vanadium is especially detrimental, because it can accelerate the destruction of the zeolite. For the lab practitioner, metals add a further complication.

The most commonly used method is some variant on the techniques described by Mitchell that involves a wet impregnation using metal naphthenates followed by steaming (1 6). Unfortunately, these techniques do not accurately assess the effect of metals in equilibrium catalysts and even worse, can give contradictory information. The major weaknesses are that the vanadium deactivation is much too severe and the dehydrogenation effect of lab impregnated nickel and vanadium is greatly exaggerated. Zeolite destruction is principally influenced by steam partial pressure, as described by Wormsbecher (17)’ and most of the commonly used lab methods involving metals deposition use close to 100% steam. Catalysts that have been metallated by the Mitchell approach generally produce much higher coke and hydrogen levels compared to equilibrium catalysts with the same metals level. Therefore, most workers attempt to overcome this effect by using a fraction (1/3-1/5) of the equilibrium metals levels on lab treated catalysts, as suggested by Cimballo, et. al. (18).

Other limitations of traditional impregnation methods are: metal passivators, such as antimony and bismuth, do not always show their commercially well

266

established performance improvements; metal passivators (especially for vanadium) can be developed which show excellent performance in the laboratory, but which perform poorly in a commercial FCCU. Examples of this are barium titanate and some of the basic oxides, such as MgO. Generally, these so-called V traps work well with high steam pressure that gives rise to high vanadic acid formation and vanadium mobility resulting in high trapping efficiencies. Also, most of these lab protocols are conducted in the presence of pure steam or with small concentrations of air, but rarely in the presence of sulfur oxide gases. The so-called V traps are more selective to sulfate formation than to vanadate formation, hence the selective pick-up of V can be essentially eliminated by S competition. The unwary catalyst tester will arrive at the wrong conclusion concerning so-called V stability, while the unwary catalyst developer will invent materials that perform well in lab testing, but which will show no Performance enhancement in the commercial FCCU.

Table 5 Comparison of different steaming conditions

Liakksm 4 hr.: 81 6°C 24 hrs.: 773.8"C Microactivity, w% 71 67 Unit Cell Size, A 24.25 24.24 Zeolite Surface Area, m2/g 79 85 Matrix Surface Area m2/g 48 52

l2mkLEl Microactivity, w% 66 65 Unit Cell Size, A 24.25 24.26 Zeolite Surface Area, m2/g 140 136 Matrix Surface Area m2/g 43 36 - Microactivity, w% 73 71 Unit Cell Size, A 24.39 24.39 Zeolite Surface Area, m*/g 56 51 Matrix Surface Area m*/g 63 70

74 hrs.. 760 C 40 hrs.: 730°2; . o i a a u Q l Microactivity, w% 70 70 70 Unit Cell Size, A 24.26 24.26 24.24 Zeolite Surface Area , m2/g 134 138 137 Matrix Surface Area m*/g 49 50 48

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To overcome these problems, many laboratories are returning to using some form of cyclic metal deposition. In this rediscovered technology, a catalyst is subjected to repeated cycles of cracking with a metal doped feedstock followed by a regeneration. An alternative procedure involves subjecting a metal impregnated catalyst to repeated cracking and regeneration or reduction and oxidation cycles, as described in many of the Phillips patents relating to their antimony passivation development. Several papers relating to cyclic aging techniques were presented at the AlChE symposium of FCC technology in the 1991 Annual meeting in Los Angeles (19a).

The catalyst pretreatment step is a crucial one, and the use of two approaches are recommended. For general catalyst screening comparisons, a single, clearly defined steaming treatment, giving properties, activities, and selectivities generally representative of low metal equilibrium catalysts is suggested (e.g., 4 hours at 816°C and 100% steam in a fluidized bed reactor). While there has been much discussion concerning the use of blending differently steamed catalysts (HIKE, 3), the author has found this to be largely a cosmetic effect that offers no real improvement in assessing relative catalyst performance. Furthermore, the blending of fresh catalyst into the steamed catalyst is definitely not recommended, as it appears to disguise coke selectivity differences and is a gross exaggeration of the age distribution. Table 6 shows analyses of the same type of catalyst (Catalyst D from the case history in Table 2) deactivated using: i) a homogeneous steaming (4 hours at 816OC); ii) blending 5% fresh catalyst; iii) low metals equilibrium catalyst. The MAT yields (Table 7) show good agreement between the equilibrium and homogeneously steamed catalyst. The blended catalyst shows significantly higher coke and lower gasoline, and generally does not simulate the equilibrium catalyst performance.

Table 6 Lab pre-treatment vs. equilibrium catalyst properties

L Steamed 5% Fresh

4 hr.. 816°C 95 % 4 Hr. 816°C Eau ilibrium

Al2O3, wt% 30.0 30.0 29.9 RE2O3, wt% 0.89 0.89 1.07 Na20, wt% 0.39 0.39 0.4

93 400

Zeolite Surface Area, m2/g 143 147 131 Matrix Surface Area, m2/g 34 35 30 Unit Celt Size, A 24.24 24.24 24.24

Ni, ppm v, PPm

268

Table 7 Comparison of selectivities

w

4 hr.. 816°C 95% 4 hr.. 81 6OC Steamed 5% Fresh

Conversion, wt% c / o

c4= iCA T&al C ~ S


Coke

GC RON GC MON

65 3.9

0.05 2.3

4.7 5.7

5.7 3.1 9.5

45.0 21.5 13.5

2.4

90.5 80.5

65 3.3

0.06 2.6

4.3 5.9

5.3 3.5 9.8

43.2 21.6 13.4

3.4

91.2 80.9

65 3.8

0.04 2.2

4.6 5.7

5.3 3.6 9.7

45.0 21.4 13.6

2.3

90.3 80.7

The second technique, which is commonly used when selecting a catalyst for a specific commercial FCCU, attempts to match the deactivation severity to the properties of the base equilibrium catalyst. This can generally be done using time at constant temperature, and is especially practical i f a multiple reactor steaming system is used. Varying temperature can also be used, but with caution. Temperatures above 830°C or below 760°C should be avoided. At high temperatures, unrealistic zeolite sintering rates can be encountered, and at low temperatures, it may be impossible to match the equilibrium unit cell size, especially for highly exchanged (rare earths) zeolites. Once the specific deactivation severity which provides a good match of steamed and equilibrium catalyst properties (MAT activity, zeolite unit cell size, zeolite and matrix surface areas) is determined, that deactivation protocol is then applied to all contending catalysts. Catalyst activity should be adjusted, if appropriate, by modifying the catalyst formulation (Le., selecting a different grade level), not by the severity of the lab deactivation.

269

In the case where contaminant metals are important, the use of cyclic metal impregnation and aging will become the standard practice for laboratory studies of the future. Until these cyclic techniques have been well established, great care should be exercised when employing traditional methods for preparing metal treated catalysts. One approach is to perform separate testing to assess the relative impact of nickel and vanadium, in addition to metal-free selectivity testing. Another is to use 1/3 to 1/5 of the equilibrium metal. However, the results should be reviewed care-fully, as there are "metals traps" which give excellent results in traditional laboratory testing programs, but which do not show performance improvements commercially.

5. PILOT PLANTS FOR FCCU EVALUATION

The key to any significant advance in FCC technology, whether it be in process development or in catalyst improvement, is the ability to accurately assess performance in the laboratory and have a successful bridge to the real world (the scale-up problem). In FCC work, evaluation of catalyst selectivity (and activity) is typically done using a variant of the Microactivity test (MAT) described in the ASTM 03907 or D5154-91 procedures. MAT is a small scale test employing a fixed (packed) bed of approximately 4 to 6 grams of catalyst, the results of which are difficult to directly scale up to a commercial, short residence time riser FCCU. As pointed out previously, MAT results can give totally contradictory rankings, and hence may require sophisticated reaction engineering models to predict even approximate FCCU performance. Several workers have recently described attempts to overcome the inherent limitations of the MAT (19, 20). Suggestions have included increasing reaction temperature and using annular catalyst beds. However, the MAT will always be an unsteady-state system and can never hope to simulate the lean phase reaction conditions nor the heat balanced regeneration- cracking cycle that is the FCCU. Because a casual inspection of MAT results may not provide any meaningful information, more pertinent data are frequently obtained in larger pilot units designed to mimic commercial FCCU's.

There have been numerous articles written about different pilot unit designs (21-30), and the types of pilot plant most favored are those which have riser reactors. The most desirable configuration is one that also includes continuous regeneration of the coked catalyst, but to simplify the design and operability, so- called "sling-shot" or once-through reactor systems have been popular. These type of pilot plants do not provide continuous regeneration of the spent catalyst, and therefore can require large inventories of catalyst, making it difficult to examine fresh catalysts (as opposed to equilibrium catalyst from the commercial FCCU). Furthermore, they are limited in the type of studies that can be accomplished. Often, these pilot plants will not reach steady state by the time data are being collected. Many of the larger pilot plants process 1 to 5 bbl/d of feed, and may require headroom of 40' or more to accommodate the riser length needed to give adequate catalyst and vapor residence times. By today's standards, many of these are dinosaurs requiring high levels of staffing, maintenance, feed, and catalyst. Some have been specifically designed for residuum processing and have modern control systems; however, the larger the pilot plant, the greater the cost to operate and maintain it.

270

Early in the 1970’s, ARCO published details (see references 21, 22) of a small circulating (i.e., continuous regeneration) pilot plant (LAB FCC) that employed a moving fluid bed reactor that was claimed to mimic bed type FCCU’s. This technology was offered for license to the industry and because of its small size, became a popular pilot unit in many companies conducting FCCU studies. ARCO continued to provide improvements to the technology (e.g., introducing a five-stage reactor to decrease backmixing) in an attempt to better simulate riser cracking. Later, two types of folded riser were offered, along with pressurized operation and some computer control (22). With the closing of the ARCO research laboratory in Harvey, Illinois in 1986, no further technical development work has been reported by ARCO for their LAB FCC. Among the major weaknesses of this technology are: limited ability to process modern heavy oil and resid feeds; inability to closely simulate the modern commercial riser reactor with high temperature regeneration: operating uncertainties with the folded riser design, especially at typical commercial pressures; inability to simulate the interrelation of the process variables; operating as a controlled isothermal reactor rather than the adiabatic reactor of the commercial process. In the mid 19703, Davison designed a pilot plant to closely resemble the commercial slide-valve FCC unit. The original circulating pilot plant was commissioned in 1979, and was subsequently described in a 1983 article (29) in which its performance was compared to that of a fixed fluid bed reactor and MAT. The original Davison design had several limitations and was difficult to operate with manual pneumatic controls and the use of a partitioned box furnace heating system.

In 1986, a major redesign was undertaken. The result was the Davison Circulating Riser (DCR) pilot unit that featured complete computer control (31). In the original design of the DCR, the reactor was heated in five different zones, each one being independently controlled so that the reactor temperature profile could provide a forced adiabatic profile that simulated the commercial one. Alternatively, the DCR could be operated isothermally for kinetic studies. Subsequent improvements have permitted both true adiabatic (32) and fully heat balanced operation.

6. GENERAL PROCESS DESCRIPTION OF THE DCR

The DCR is a small scale (less than 12 feet total height) FCCU pilot unit that features a vertical, lean phase riser reactor, which can be operated under isothermal, adiabatic, or pseudo-heat balanced conditions. The proprietary reactor design is unique, allowing for operation with catalysts of widely varying particle properties (density, shape, size) with minimum effect on reactor hold-up. This means that the unit can be used for fresh, steamed, or equilibrium FCC catalysts, and does not require recalibration to define contact time. Table 8 provides a summary of the normal operating ranges of the main process variables for the DCR.

271

Table 8 DCR operating ranges

System Pressure, Bar (abs) Riser Temperature, "C Regenerator Temperature, "C Stripper Temperature, "C Stabilizer Temperature, "C Feed Preheat Temperature, "C Feedrate, glh Catalyst Charge, g Catalyst Circulation Rate, g/h Feed Types

2 - 2.8 510 - 540 690 - 750 480 - 540

-34 120 - 400

400 - 1500

4500 - 7500 3000

VGO, CGO, Resid, ATB (up to 5.3 w% Conradson Carbon)

PROD UNIT PRES LINE PSIC 2.8

15:10:22 M A C PROCESSING OVERRUNS: 8 CYCLES SKIPPED 15:11:34 M A C PROCESSING OVERRUNS: 4 CYCLES SKIPPED

MNL FEEDAllFE B 0 B I R ON

Enter data: D ISP : Ha inu iew ENABLE

Figure 1. Process computer display of DCR operational status

272

Figure 1 shows the schematic of the DCR that is normally presented on the controlling computer's display screen. Feed is transferred from one of the twin storage tanks into one of the feed weigh cells (both of which are directly linked to the process control computer). The dual feed supply permits a rapid switch from one feed type to another while the system is operating. A metering pump precisely controls the feed rate as feed is pumped from the load cell through a preheater to the injection nozzle. Either nitrogen or steam can be used as a feed dispersant, and is injected through a separate preheaterhaporizer. Dispersed oil passes into the proprietary vertical riser and mixes with hot catalyst returning from the regenerator.

The reactor is equipped with both surface and adiabatic heaters. The adiabatic heaters are controlled to eliminate heat transfer across the reactor wall. Alternatively, the reactor surface heaters, located in five independently controlled zones, can provide a pseudo-adiabatic or an isothermal temperature profile in the riser. Reactor temperatures in excess of 595OC can be achieved, and any individual zone can be controlled within +l0C during the reaction. Internal thermocouples independently monitor the actual catalyst-vapor mix along the riser. Pressure taps across the riser monitor the pressure differential (DP), which is directly proportional to the catalyst hold-up in the reactor, and which can be used to determine the actual WHSV. Oil, catalyst, and dispersant pass from the riser into the stripping disengager.

Products exit the stripper through a refrigerated stabilizer column to a control valve which maintains unit pressure at the desired level (usually 2.7 bar). Spent catalyst drops into the vertical stripper standpipe, forming a dense phase fluidized bed which slowly moves down to the stripper slide valve. This slide valve controls catalyst flow from the stripper to the regenerator and uses the DP of this transfer line as its controlling set point. For non-adiabatic operations, adjusting this set point becomes the primary way to adjust catalyst circulating rate.

Part of the stripper-regenerator spent catalyst transfer line consists of a jacketed heat exchanger (hot catalyst in the tube and cold air in the jacket). The heat transfer across this exchanger provides a precise and reliable method for the direct, continuous display of catalyst circulation rate (and by calculation, also of

In the regenerator (also a dense phase fluidized bed), spent catalyst is burned clean with mass flow controlled air. Excess air and combustion products exit the regenerator through control valves, are cooled and then continuously analyzed for oxygen, carbon dioxide, and carbon monoxide, and then flow metered before being batch collected for subsequent GC analyses. The system also has the capability to continuously analyze for SOX and NOx. The oxygen analyzer provides the set point to the mass flow controller for the regenerator air rate. Although normally operated in full combustion with 4% 0 2 in the flue gas, the DCR is capable of operating in partial combustion when a continuous CO analyzer is added to the flue gas analysis train.

Regenerated catalyst passes down the regenerator, through the slide valve, into the catalyst return line, which is independently heated. For non-adiabatic operations, the slide valve uses the regenerator DP as set point. The regenerator is also provided with multiple zone heating for complete independent control over regenerator temperature (and profile). In routine adiabatic operations, the regenerator temperature is usually in the range 70O-73O0C, but for specific commercial simulation, the commercial dense bed temperature is used.

C/O).

273

The gaseous hydrocarbon products exit from the stripper and pass through a primary stabilizer column that provides a product cut between C4 and C5. The liquid products (C5 and heavier) are condensed and batch collected for subsequent distillation to provide gasoline, LCO (221 - 343"C), and 343"C+ bottoms fractions. The gaseous products are metered and usually batch collected for analysis by GC. Batch collection of the products was chosen to minimize turnover time in the stabilizer and relieve the need for continuous fractionation.

The DCR is fully instrumented with computer control and monitoring, and includes back-up, fail safe control, and data logging features. The system process control software is Intellution's (33A) FIX, operating in a DOS compatible PC (386 or 486 computer) with a high resolution color monitor protected with an uninterruptable power supply (UPS). The system provides process control and monitoring, safety management, alarm functions, data acquisition and trend analyses, and real time display of system parameters. The computer maintains historical data records (temperatures, DP, etc.), which can be examined as time trend plots.

A typical mass balance run (at 1000g/hr feed) lasts one hour, and follows a line-out period of perhaps a couple of hours. Operating on two shifts, it is routinely possible to obtain four or five mass balanced runs under different operating conditions (i.e., C/O ratios). Mass balance closure is normally greater than 97%, and averages close to 99.

6.1. Isothermal Operation A detailed description of the isothermal mode of control of the DCR was

provided in a 1988 NPRA paper (31). Briefly, the catalyst circulation is controlled using the pressure drop in the transfer line from the stripper to the regenerator as the set-point for the stripper slide valve. The regenerator slide valve uses the pressure drop in the regenerator as its set-point, thus ensuring a constant bed height in the regenerator. Feed and catalyst temperature are kept constant. This mode offers the greatest flexibility, and a wide range of C/O ratios is achievable.

6.2 Adiabatic Operation To be able to provide a more realistic simulation of the commercial FCCU, the

DCR reactor was equipped with adiabatic heaters and the catalyst circulation control method altered. The use of adiabatic heating is well known, and its use for FCCU pilot plants has been described by others (25, 27). The technique involves including a second set of heating elements on the outside of the reactor insulation layer and maintaining a zero temperature differential to the reactor wall. It is worthwhile to note that not even commercial FCCU's are truly adiabatic, since there can be considerable heat loss by axial conduction and by radiation. Therefore, with the adiabatic heating arrangement installed on the DCR, it is feasible to investigate the influence of heat loss (or gain) by adjusting the power to the adiabatic heating elements and adjusting the delta T.

In the adiabatic mode, the process control strategy is the same as for many commercial FCCU's. The riser outlet temperature is used as the control set point for the regenerator slide valve, which directly controls the catalyst circulation rate. Control of the other slide valve is done using the delta P in the transfer line from the stripper to the regenerator as set point. Using this control scheme, the feed temperature becomes the primary means for altering the catalyst circulation rate and permits a method of determining conversion-yield relationships at constant riser

274

outlet temperature and over the range of conversions that are of interest to the commercial operation.

This operating mode not only provides a stable process, but has been frequently used to successfully match commercial operations, not only in yields and conversion, but also in key process variables like catalyst to oil ratio (C/O) when operating at the commercial feed and regenerator temperatures.

Table 9 compares, at constant conversion, the results from testing the same catalyst adiabatically and isothermally. The adiabatic mode produces higher C2-, propylene and butenes, with lower isobutane, gasoline, and light cycle oil. Gasoline octane was higher in adiabatic testing.

Table 9 Comparison of isothermal and adiabatic modes

Conversion

H2 c2-

c3= C3 total

c4= iC4 C4 total

CS+ gasoline RON MON

LCO HCO

Coke

78

.64 1.7

4.8 5.5

6.4 2.7 9.7

52.9 92.9 80.9

13.6 8.4

6.7

78 78

.40 2.1

4.6 5.7

5.3 3.8 9.9

51.5 93.1 81.2

12.7 9.3

7.6

.47 2.4

5.3 6.1

7.3 2.2

10.0

51.3 94.5 81.6

12.0 10.0

6.6

78

.34 3.0

5.6 6.7

6.6 2.8

10.1

50.1 94.1 81.6

11.3 10.7

6.9

Test C o w Riser Exit Temp, "C 527 527 52 1 52 1 Catalyst Temp, "C 538 538 746 746 c/o 5.6 4.6 5.2 4.1 Feed Temp, "C 31 6 31 6 260 371

. .

275

6.3 Heat Balance Operation While the adiabatic mode provides the essential features of commercial FCCU

control and better insight into process variable effects, its weakness is that all operations are usually done at constant regenerator temperature. It is a simple matter to estimate the regenerator temperature using a heat balance calculation procedure (34) along with the flue gas analysis. In the DCR, there is continuous analysis of the flue gas, hence the heat balanced regenerator temperature can be continually calculated. The result of the calculation can then be used to update the regenerator temperature set-point. Care must be taken not to over-control. For example, in this mode when the feed temperature was changed from 149°C to 371°C the regenerator temperature changed from 712°C to 732°C. Table 10 lists the process variable values from a yield comparison study of a single catalyst performed in both the adiabatic and heat balanced modes. Only minor differences in the selectivities (constant conversion comparison) were obtained between modes.

Table 10 Comparison of adiabatic and heat balanced pilot plant [constant 70 w% conversion]

Riser Top, "C Feed Temperature, "C. Regen Temperature, "C c / o Yields. wYQ H2 SCF/BBL C2- c3=/c3 C3 total c4=/c4 C4 total C5 gasoline LCO HCO Coke

Adiabatic

504 269 704 6.6

125 2.8 0.69 5.8

0.584 8.2 47.6 17 13 5.1

Heat Balanced

504 220 71 3

6.6

125 3.1 0.69 5.8 0.576 7.7

47.6 17 13 5.1

6.4. Laboratory Operations of a Pilot Plant Each of the various operating modes that are possible using a flexible pilot

plant such as the DCR may cause some variation in the approach to evaluating catalyst performance. A typical technique involves obtaining mass balanced yields at different operating conditions and generating a yield curve (product yield vs. conversion or coke). Preliminary analysis of the results can highlight intrinsic differences in catalyst activity and selectivity, but will not directly suggest how different catalysts will perform in the commercial FCCU. An exception to this

276

evaluation technique comes when testing in the heat balance mode. For a direct indication of how a new catalyst or feedstock will respond, the pilot plant is initially set up to mimic the commercial FCCU heat balance operation (same riser, feed, and regenerator temperatures, system pressure), which should produce the same catalyst circulation, conversion, and yields. Feed rate to the pilot plant may need to be adjusted to match the commercial residence time. Using these conditions as the base case, introducing the new feed or catalyst, and allowing the system to re- equilibrate will give a direct indication of the changes to be expected in the FCCU.

Most pilot plant studies can be done by generating a yield curve. In an adiabatic pilot plant, this is usually accomplished by changing feed temperature over the commercial range 150-370°C while maintaining riser and regenerator temperature constant. Typical yield responses are shown in Figures 2A-C. For a well defined curve, a minimum of four conditions should be obtained. In a small pilot plant, this can easily be achieved in a single, two-shift day (1 6 hours).

Yields of C,'s, C;s and LCO Yields of Dry Gas and Coke

i # > 9 ; 10

5 1 60 64 68 72

CONVERSION, V%

Figure 2a

Yield of Gasoline

50 1 60 64 68 72

CONVERSION, V%

Figure 2c

4 j t

64 60 72 CONVERSION, V%

Figure 2b

Figure 2. Reproducibility of DCR testing; one catalyst, multiple tests over four weeks

277

The key to efficient laboratory operation of a pilot plant is in the length of time required for a mass balanced test and the amount of time required to reach steady state following a process change. The latter depends on the design and control of the pilot plant. Pilot plants incorporating on-line fractionation systems add an order of magnitude to the complexity of the operations, and generally there is little need for such a system. Its main use is to be able to conduct continuous recycle studies. However, a pilot plant with batch product collection can easily perform such studies using pre-blended feedstocks and offers a much more cost effective way of operating.

The length. of time required for a mass balance run is usually dictated by the stability of the process variables and the amount of product required for subsequent analyses. The amount of product is usually dictated by the quantity of gasoline required for octane rating on the CFR knock engine. For conventionally operated engines where both MON and RON are being determined, approximately 1 liter of gasoline is required. This in turn requires that 2 to 4 liters of liquid product is needed from the pilot unit and hence, mass balance runs for the small pilot plants may have to be 2 to 4 hours. In addition, working with these large amounts of liquid products requires the use of large capacity distillation equipment.

Figure 3. Davison Laboratory Automated Distillation (LAD) Unit

278

To overcome these limitations, the knock engine can be equipped with a microcarburetor and the falling fuel method of octane rating can be employed (35). This technique has been in use in our laboratories for over twelve years and, based on Davison Research's participation in the ASTM D2 committee's Atlantic region round robin testing program, the results are indistinguishable from the conventional method. This technique requires only 30 cc of gasoline for an octane rating and hence, distilling 200 to 250 cc of liquid product provides ample sample for all the analyses of the gasoline and cycle oils. Figure 3 shows one of Davison's specially designed automated laboratory distillation units (LAD) that offers multiple plate (approximately five) distillation and automated collection of gasoline, light cycle, and heavy cycle oils. With having such a low sample requirement for complete product analyses, the normal mass balance period is one hour. Following the mass balance period, the feed temperature is changed and while the catalyst circulation and temperatures equilibrate quickly, it takes longer for the stabilizer column (primary separation of C4- from Cg+ liquids) to re-equilibrate. Normally, this takes about 11/2 to 2 hours. Figure 5 shows plots of some of the key process variables over a typical two-shift day.

TE49: FEED TElOl :

TEMPERATURE

990- R I S E R

T E

P E R A T U R E

CATALYST CIRCULATION

940- RATE

RISER BOTTOM TEMPERATURE ( 1000°F+70): TE105 CURSOR DISPLAY

Figure 4. Principal DCR variables response to feed temperature changes

279

While we continue to perform actual distillations and traditional property measurement methods, the improvements in GC techniques are permitting simulated distillation and capillary analyses of gasoline to be used as excellent substitutes for the older time and labor consuming methods. Thus, some of the same features that make the MAT test so attractive can also be applied to pilot plant products to provide very rapid and cost effective determination of the yields. The ASTM procedure for simulated distillation, 02887, can be used to provide product yield determinations within 1 to 2 hours. Furthermore, once the GC analysis is available, yields based on any cut points for the liquid products are also immediately available. Using the combination of a prefractionator equipped capillary GC analysis and a licensable gasoline property model such as G-CON@ (4) can provide octanes and much more (Table 11) in short time and with negligible sample preparation time. Table 12 shows yields determined entirely by GC techniques compared with those determined using the traditional methods. The GC-based information can be available within a few hours at negligible operating cost, whereas the traditional analyses may take several days (or weeks) and at high cost. It is our experience that the GC analysis methods provide better precision than the classical methods. Furthermore, with the current interest in composition of gasoline, techniques such as G-CON@ are proving to be much more valuable than a simple CFR engine octane measurement.

Table 11 Program G-CON@ Uses identity checked peak data to calculate:

Wt.% and Vol.% yields of individual peaks and the PIANO hydrocarbon groups

RON and MON octanes Reid Vapor Pressure Specific gravity and Molecular weight Bromine Number and Aniline Point All results on de-butanized, de-pentanized or de-hexanized basis, if desired Adds separate gas phase yield data such as from debutanized overhead Fractionates sample at user specified cut temperatures in OC or O F

Outputs results to printable files and spreadsheets such as LOTUS or EXCEL Calculates for 1 sample at a time or from a list of many samples

6.5 Pilot Plant Reproducibility Figures 2A-C show the type of reproducibility that can be expected from a pilot

plant. These results are from the same catalyst tested one day each week, over a four week period. To consistently achieve this level of repeatability, it is essential to have a regular preventative maintenance program coupled with regularly scheduled calibrations of thermocouples, flowmeters, gas chromatographs, oxygen, and C02

280

meters. If these steps are not taken, pilot plant or analytical drift can make comparisons of day to day results worthless. It is essential that a standard catalyst be regularly tested under standardized operating conditions. Additionally, the use of statistical process control analysis techniques can greatly improve the operator’s awareness of the pilot plant’s condition.

Table 12 Comparison of pilot plant yields determined using true distillation or simulated distillation

66.0 66.0 79.3 80.3 Conversion, w%

Yields. w% H2 c1 +c2

c3= C3 total

c4= iC4 C4 total

0.55 1.9

0.55 1.9

0.63 1.6

0.63 1.6

2.9 3.4

3.0 3.5

4.9 5.6

5.0 5.7

4.8 1.3 6.3

4.9 1.3 6.6

6.8 3.0

10.4

7.1 3.1

10.8

Gasoline 47.7 47.5 52.2 52.6

LCO HCO

16.0 18.0

16.4 17.6

12.7 8.0

11.4 8.3

Coke 4.7 4.7 7.8 7.8

6.6 Commercial Simulation with an FCC Pilot Plant In the paper that first described the DCR (31), a comparison was given

between the commercial FCCU at Marathon’s Robinson refinery and the DCR operating in the isothermal mode. To be able to obtain the yield agreement, numerous test runs had to be made in the DCR to be able to fine tune some of its operating parameters (e.g., the catalyst temperature returning to the reactor mix zone). The resulting simulation of the commercial FCCU was quite good. In a subsequent paper that described the adiabatic operation of the DCR (32), it was demonstrated that the DCR, with the same basic operating conditions as used in the commercial FCCU’s, not only showed yields that were a close match of the commercial ones, but DCR catalyst/oil ratio, now a dependent variable, was similar to the commercial values (Table 13). Two of the licensees of the DCR technology have also reported good success at simulating commercial FCCU’s, as shown in

28 1

Tables 14, 15. In one of these, the DCR produced the same process conditions, conversion, gas, LCO, and gasoline octanes. The coke yield from the DCR was lower than commercial, and gasoline slightly higher. In the other comparison, the DCR C/O was slightly higher; gas yield and gasoline yields were comparable, with similar octanes. LCO yield from the DCR was 2V% lower, and coke yield was slightly lower on the DCR.

Table 13 Commercial simulation with the DCR (From Reference 32)

~ ~

Feedstock A A B B

. . 0-

Riser Outlet Temperature, "C 531 53 1 51 9 51 9 Catalyst Temperature, "C 733 733 71 3 71 3 Oil Preheat Temperature, "C 309 309 246 246 System Pressure, Bar 3.4 2.05 2.7 2.7

Conversion c/o Hz c1 +cz Total C3 Total C4

Gasoline RON MON LCO Bottoms

Coke

73.7 5.5 0.04 2.8

10.5 15.1

59.5 92.4 81 .O 16.5 9.8

4.3

75.6 5.3 0.1 3 3.3

12.9 16.4

59.6 95.2 81.6 14.8 9.7

3.9

69.8 70.1 6.2 6.2

NA 0.29 4.6 2.5 4.8 5.1 8.4 7.2

47.6 48.6 NA 93.8 NA 81.4 19.7 20.3 10.5 9.6

4.4# 6.4

# Data does not heat balance. With partial combustion (3/1 CO2/CO ratio) coke yield should be 6.1w0/0 to heat balance.

282

Table 14 FCCU - DCR correlation data from licensee 1

Feed Fccu m

Riser Temperature, "C Cat/Oil Conversion, wt% Yields fwt%) Gas Gas0 Ii ne RON MON LCO HCO Coke

505 4.3 57.6

14.44

90.0 78.9 24.95 17.46 4.75

38.4

505 4.2 57.7

13.76 40.35 90.0 78.4 24.61 17.66 3.6

Table 15 FCCU - DCR correlation data from licensee 2

Fccu DCR

Operati na Condition Riser Temperature, "C c / o Feedstock Conversion, w% Yield (wt%l Fuel Gas LPG Light Gasoline (C5-150°C) RON MON Heavy Gasoline (1 50-1 85°C) Naphtha (1 85-260°C)

HCO (340+C) Coke

LCO (260-340°C)

51 5

FCCU (1 7/1/90) 5.9

66.2

2.3

31.1 93.1

6.4 12.7 13.3 20.4 4.5

8.7

78.3

51 5 6.6

67.2

2.2 9.2

31.4 93.3 79.4 7.2

13.1 11.3 21.4 3.9

283

Table 16 Physicakhemical analyses

Catalyst A Catalyst A Catalyst B Equilibrium Steamed Steamed

2 hr. 816OC 2 hr. 816°C

A1203, wt.% Re203, wt.% Na, wt.% Fe, wt.% v, PPm Ni, ppm

J33U2 0-20 p 0-40 p 0-80 p 0-1 05 p 0-149 p

ABD, glcc BET surface area, m2/g Zeolite, surface area, mag Matrix surface area, mag Unit Cell size, A X-Ray peak intensity

Il!wLwa Conversion, wt.% H2 Yield, wt.% 0.102 Coke, wt.%

33.6 1.65 0.39 0.52

71 0 454

0 3

56 83 99

76

0.86 142 108 34 24.29 56

68 0.075 3.3

34.0 1.38 0.22 0.42

70 28

0 7

46 70 97

84

0.86 133 96 37 24.32 57

69 0.061 3.2

32.3 0.72 0.14 0.37

60 23

0 6

50 75 97

80

0.80 167 128 39 24.27 70

70

2.5

7. STUDIES WITH THE DCR

7.1 Catalyst Selection Study for a Commercial FCCU It was suggested earlier in this chapter that the selection of a catalyst for a

commercial FCCU should be conducted with fresh catalysts, but using its equilibrium catalyst as a guide to the deactivation conditions. In a DCR study of this type, the equilibrium catalyst can also be tested in the pilot plant using the commercial feedstock to establish the validity of the deactivation protocol of the fresh base catalyst. Additionally, if the commercial yields and operating conditions are available, then the pilot plant's ability to simulate the commercial unit can also be assessed.

284

Table 16 shows analyses of equilibrium and fresh, deactivated catalysts in which the lab catalyst was deactivated for 2 hours at 816°C in 100% steam. There is generally good agreement in the key properties, and the metals (Ni, V) levels on the equilibrium are moderately low. The proposed catalyst, chosen to improve octane and coke selectivity, was also steamed at the same conditions. All three catalysts were tested in the DCR at 532°C riser outlet and 716°C regenerator temperatures varying feed preheat from 149 to 371°C The resulting yield curves were interpolated to provide the constant conversion and constant coke comparisons shown in Table 17.

Table 17 Catalyst constant conversion comparison - vol. % Simulated distillation with distilled product specific gravities (Yields were interpolated and mass balanced to 100°/o).

Catalyst ~~~

Catalyst A Catalyst A Catalyst B Equilibrium Fresh Fresh

2 hr. 816°C 2 hr. 816°C

c / o Conversion, vol% H2, wt% c 1 + C2'S, wt% c3=, vol% nC3, vol% Total Cg's, vol% c4=, VOlYO iC4, vol% nC4, vol% Total C4, vol%

Gasoline, 221 "C, vol% API Gravity RON MON Aniline Point, OF Bromine number

API Aniline Point, OF Bottoms, 64OoF, vol% API Gravity Viscosity @ 122°F

Coke, wt%

LCO, 221 -343"CF, VOI%

8.7

0.09 2.4 9.4 1.8

11.2 11.0 5.5 1 .o

17.5

63.0 57.3 94.7 82.1 63 72

78

12.6 17.0 16 9.4 4.0

4.6

65

7.2

0.06 2.4 9.5 1.8

11.3 11.1 6.0 1.2

18.3

62.7 58.0 94.4 82.1 64 70

12.2 17.0 16 9.8 5.0

78

58

4.4

7.2

0.04 2.5

10.1 1.9

12.0 11.5 5.7 1.1

18.3

62.7 58.3 95.3 82.1 62 77

12.3 17.3 16 9.7 4.0

4.0

78

58

285

These results show that there is good agreement between the equilibrium and lab treated base catalyst performances, and that the proposed catalyst offers improved catalyst selectivities.

Table 18 Comparison of catalyst testing in different reactor types Yields at 70% conversion (from reference 29)

Test Method 5!2 EEE MBI

ODeratina Conditions c / o 8.1 5.5 2.8 WHSV 49 29 17

Product Catalyst Coke : WYO H2 : w% c1 +c2 : WYO Total C3's : v% Total C4's : v% Total C3+C4 : v%

CS+Gasoline : v% Gravity : OAPI Aniline Pt. : O F

RON + 0 MON + 0

Light Cycle Oil : v% Gravity : OAPl Aniline Pt. : O F

Heavy Cycle Oil , v% Gravity : OAPI Aniline Pt. : O F

p 1 Q 2 3.8 3.8 0.04 0.08 1.7 1.7 9.9 9.8 14.9 15.7 24.8 25.5

55.5 55.5 59.5 57.0 76 84 88.0 88.0 78.3 78.3

20.5 23.5 23.0 23.0 89 89 9.5 6.5 8.0 2.7

127 87

Ql 5.7 0.05 1.5 10.4 12.9 23.3

54.5 61 .O 81 91.5 80.9

19.5 23.0 83 10.5 6.0

123

a2 6.7 0.10 1.5 10.3 11.7 23.0

54.5 61 .O 81 91.5 80.9

21 .o 23.0 83 9.5 5.0

123

u p 2 3.0 4.4 0.08 0.19 1.8 1.8 8.1 7.3 12.6 12.2 20.7 19.5

58.0 58.0 58.5* 5 8 5

21.0 22.0 23.0* 23.0*

9.0 8.0 9.5* 9.v

_ _ --

-- --

These values obtained by correlation, not measured.

7.2 Comparison of Microactivity and Riser Testing The selection of the pretreatment conditions of the catalysts is a very important

aspect of assessing realistic catalyst performance, however the method of testing the catalytic performance is also crucial. The common types of laboratory reactors in use for FCC catalyst performance testing are the Microactivity test (MAT), the fixed fluid bed reactor, and the riser reactor. A comparison of the performance of

286

these reactor types was reported several years ago by Creighton and Young (29), who concluded at that time that while there were differences in absolute yields for the three types of reactor, they agreed on the major selectivity ranking for two lab steamed catalysts uncontaminated with nickel or vanadium. The catalysts were of similar type (same zeolite type and catalyst family), but differed only in their active matrix content, which influenced the bottoms cracking and coke selectivity.

Close examination of the results of that study (Table 18) show that different reactor types could give different ranking of the catalysts for some of the selectivities. Comparing the coke selectivities of the catalysts D1 and 02, the MAT shows that D2 produces almost 50% higher coke than D1; the riser pilot plant, on the other hand, shows both catalysts producing the same coke yield. Based on this MAT result, and because of the great importance of coke yield on commercial operability, catalyst D2 would probably be inappropriately rejected. That paper suggested that the reason for the difference in coke selectivity measured by MAT between the catalysts was attributable to the longer contact time in the MAT, and suggested that D2 had a higher coking rate than D1. That comparison also showed different catalyst rankings for C4's; in the riser, catalyst D2 produced more, in the MAT, D l produced more. The riser showed larger LCO differences than the MAT, but with the same ranking.

A recent study has compared two steamed (with Ni and V) catalysts in the MAT and the Davison Circulating Riser (DCR) operating in the isothermal mode. In this work (Table 19), while the ranking for coke selectivity was not changed, the DCR showed greater coke selectivity differences between the catalysts than were observed in the MAT. In addition, the hydrogen yields were lower in DCR testing, and the DCR showed differences between the catalyst's selectivity for LPG olefins. Selectivity differences were noted for the gasoline and cycle oils in both the MAT and the DCR, although the DCR showed a greater gasoline difference. Absolute yields were similar, except as noted above and in the cycle oils where the DCR produced higher bottoms (HCO) and less light cycle oil (LCO). Analysis of the gasolines by G-CON@ indicated similar RON, but higher MON from the MAT testing.

In another study, again in which the catalysts had been pretreated with Ni and V, a reversal of catalyst ranking by the DCR and MAT was observed. Table 20 shows the MAT results interpolated to constant conversion showing Catalyst 1 making less coke, more gasoline, and the same LCO as Catalyst 2. DCR testing (isothermal mode) of these same catalysts, using the same feed, resulted in lower coke, higher gasoline, and higher LCO for Catalyst 2. Another DCR comparison, this time performed adiabatically, produced comparable rankings at higher conversion to the isothermal study and again contradicted the MAT. Other MAT studies were done with these same catalysts, but using a different feedstock. Those studies did not alter the MAT rankings for coke, gasoline, or LCO selectivities. Ashland Oil, in a poster paper at the 1988 American Chemical Society meeting in Los Angeles (36), has reported a similar reversal of catalyst ranking between MAT testing and Riser tests of metallated catalysts.

287

Table 19 Comparison of DCR and MAT: case study 1 Constant conversion comparison, wt.%

GaU!Sm hnATm

Catalyst G w Y s M MATm

Conversion, wt% c / o

Yield wt%

H2 c1 +c2

c3= Total C3’s

c4= iC4 Total C4’s

Gasoline G-CON@ RON G-CON@ MON

LCO Bottoms

Coke

70 70 4.5 8.0

1.1 0.56 2.7 1.95

3.8 3.9 5.0 4.9

3.9 4.6 3.0 3.0 7.8 8.2

44.4 43.7 92.5 92.4 82.7 80.9

19.0 16.6 11.0 13.4

9.0 10.7

70 70 5.1 11.6

1.2 0.66 2.8 1.95

3.8 4.2 5.0 5.1

3.9 5.2 2.9 3.3 7.7 9.0

43.1 40.3 92.8 93.1 83.0 81.3

18.5 16.0 11.5 14.0

10.2 13.0

These reversal of trends are not only catalyst dependent. In an earlier paper (31), the DCR was used, in the isothermal mode, to assess the impact of a proposed feed change at a mid-west refinery. The results of that work indicated that the main impact of the “winter” feed was a difference in the LCO selectivity. Subsequent to reporting those results, a MAT study using the same equilibrium catalyst and feeds was carried out. The results, summarized in Table 21, indicate some major differences, especially in coke yield (Figure 5). The MAT data showed the LCO selectivity differences, but also suggested there would be a significant impact on the gasoline. Consistent with other studies comparing the DCR with MAT data, the MAT produces much higher hydrogen and lower olefin/saturate ratios.

288

Table 20 Comparison of DCR and MAT: case study 2

Yield. - MBI DCR &W.atk

Catalyst Conversion, wt%

Yiel- H2 c2-

c3= C3 total

c4= iC4 C4 total

Cg+ gasoline

LCO HCO

Coke

2 75 75

- 1 -

0.62 0.39 1.5 2.0

4.1 4.2 4.6 5.2

6.1 5.2 2.3 3.2 8.8 9.1

53.2 51.3

14.6 13.3 10.4 11.2

6.3 7.1

1 75 -

0.72 2.3

4.6 5.8

5.4 3.8 10.0

48.3

15.5 9.5

7.7

1 2 - 75 78

0.57 0.47 2.1 2.4

4.1 5.4 5.0 6.2

4.4 7.4 3.8 2.2 8.8 10.1

52.1 52.1

16.2 12.0 8.8 10.0

6.4 6.7

2 78

0.34 3.0

5.6 6.7

6.7 2.8 10.2

50.7

11.3 10.7

7.0

6 .’ I I I

: -0-MMATFeedl

: d’k BCRFeedf

9 4: A DCR Feed2

9 5: -0-MATFeed2 0 -

0 -

’45 $5 65 15 d5 CONVERSION, W%FF

Figure 5. Feeds influence catalyst coke selectivity ranking in the DCR and MAT

289

Numerous other examples of this reversal in catalyst ranking have been encountered and in each case, the results by each test method have been reproducible, indicating that there are real cause and effect reasons for the observations. Reaction engineering models have demonstrated that differences in the poisoning rates of the catalysts and time averaging phenomena produce such reversals. Also, there may be an influence of capillary condensation of high boiling feed components, especially in the MAT environment, which is characterized by low temperatures, low pressures, poor feed vaporization, and uncertain contact times.

The vast improvements in analytical techniques available for small sample amounts has propelled the MAT test far beyond it’s original purpose as an activity screening tool. With the potential for misinterpretation, it is dangerous to base important decisions such as catalyst selection or catalyst development only on MAT results.

Table 21 Comparison of DCR and MAT: case study 3

Yields. wt% - MAT DCR

Feed l&uLmd Wlnter l2a!&@d Wlnter

Conversion 77.7 77.7 77.7 77.7

H2 c2-

0.21 2.3

0.23 2.6

0.09 3.0

0.09 3.1

c3= C3 total

5.8 7.4

5.7 7.3

5.8 7.5

5.8 7.4

c4= iC4 C4 total

5.9 5.2

12.9

5.9 5.2

12.9

6.9 3.4

11.3

6.8 3.4

11.3

C5+ gasoline 51.6 50.0 51.8 51.7

LCO HCO

14.6 7.7

14.0 8.3

13.2 9.1

12.2 10.1

Coke 3.4 4.5 4.0 4.0

For the most accurate assessment of catalyst performance as it relates to the commercial environment, a short contact time riser reactor should be used. This becomes more important when dealing with metallated catalysts (nickel and vanadium) that have strong dehydrogenation activity (i-e., rapid coking rates) and with feedstocks that have significant amounts of high boiling components and

290

carbon forming residues. In these cases, selection of the reactor type becomes critical, and long contact time, unsteady state reactor designs operating at temperatures significantly below the mix-point temperature of a commercial FCCU can give completely misleading results that are almost impossible to translate to the com mercial ope rat ion.

The short contact time riser reactor is an essential piece of equipment for those involved in serious FCC research, development, or process optimization. The most effective laboratory pilot plants have the capability for adiabatic operation and direct process simulation, such that the lab results are in close agreement with the commercial results. It is also possible to use small scale riser pilot plants such as the DCR that retain the advantages of commercial simulation, but which require only modest amounts of catalyst (<3000 gms) and feed. Recent advances in computer control systems for laboratory systems permit the pilot plants to operate in the same manner as the commercial FCCU and even permit the pilot plant to operate in heat balance. Advances in analytical techniques also permit the low cost, rapid GC methods to be applied to pilot plant operations. These will significantly reduce the cost of pilot plant testing.

The MAT will continue to have widespread appeal, since it is inexpensive to set up and operate, but its results should be used with caution, even when it is used as a screening tool.

8. SUMMARY AND CONCLUSIONS

To undertake realistic laboratory assessment of FCC catalysts or FCC performance, a short contact time, steady state reactor that mimics the commercial environment should be considered as absolutely essential. It should be operated as closely to the commercial conditions as practical, and be periodically cross-checked with the commercial data. Catalyst screening should be done using laboratory pretreated materials. If contaminant metals are expected to be significant on the equilibrium catalyst, then some form of cyclic aging pretreatment needs to be performed such that the laboratory treated base catalyst performs similarly to the equilibrium catalyst. If metals are not an issue, an appropriate steam deactivation procedure will be adequate.

If only MAT results are available, then at the very least, heat and mass balances using base case commercial data along with any other relevant commercial experience should be performed.

9. ACKNOWLEDGMENTS

The author is indebted to many of his colleagues in Grace-Davison with whom he has spent many hours in discussions and heated debate over all aspects of FCC catalysis and laboratory evaluation techniques. In particular, Dr. Gordon Weatherbee, Wilson Suarez, Larry Langan, and Don Chin deserve special mention. Helpful suggestions and comments on the manuscript were received from Tom Habib, Jr. and Ray Mott. The author also acknowledges W. R. Grace and Co.- Conn's Davison Division for permission to publish this work.

29 1

10.

1

2

3

4

5 6

7

8 9

10 11 12

13 14 15

16 17

18

19

19a

20

21 22

23

24 25

26

27

REFERENCES

Beyerlein, R. A., e l al., ACS Symposium Preprints, Div of Petrol Chem., Vol. 35 (4), p. 694, 1990. Moorehead, E. L., and McLean, J. B. Hydrocarbon Processing, February 1991,

Keyworth, D. A., Turner, W. J., and Reid, T. A., Oil and Gas Journal March 14,

Licensed by W. R. Grace & Co.-Conn., Davison Chemical Division, Baltimore MD, U.S.A. Chen, N. Y., et al., Ind. Eng. Chem. Prod. Res. Dev., 16 (3), 247, 1977. Chester, A. W., and Stover, W. A., Ind. Eng. Chem. Prod. Res. Dev., 16 (4), 285,1977. Sapre, A. V., and Leib, T. M., ACS Symposium Preprints, Div. of Petrol Chem., Vol. 35 (4), p. 719, 1990. Weekrnan, V. W., and Nace, D. M., AlChE Journal 16 (3), 397, 1970. Avidan, A. A., Edwards, M., and Owen, H., Oil and Gas Journal, January 8,

Moscou, L., and Mone, R., J. Catal., 30, pp. 417-22, 1973. Magee, J. S., and Blazek, J. J., ACS Monograph 171, pp. 615-679, 1976. Moorehead, E. L., Margolis, M. J., and McLean, J. B., ACS Symposium Preprints, Div. of Petrol Chem., Vol. 33 (4), p. 575, 1988. R.P.S. FCC Technology Course, R.P.S. McElhiney, G., Oil and Gas Journal, February 8, p. 35-38, 1988. Patrose, B., and Young, G. W., Paper presented at the 10th National Catalysis Conference of the Venezuelan Catalysis SOC. Los Teques, Venezuela, May

Mitchell, 8. R., Ind. Eng. Chem. Prod. Res. Dev., 19 (2), 209, 1980. Wormsbecher, R. F., Peters, A. W., and Maselli, J. M., J. Catal., 100, pp. 130- 137, 1986. Cimbalo, R. N., Foster, R. L., and Wactel, S. J., Oil and Gas Journal, May 15, 1972. Mauleon, J. L., and Courcelle, J. C., Oil and Gas Journal, October 21, pp. 64- 70, 1985. AlChE Symposium Series (in press 1992) Editors Benslay, R., Chuang, K. C., and Young, G. W. O'Connor, P. and Hartkamp, M. B., ACS Symposium Preprints, Div. of Petrol Chem., Vol. 33 (4), p. 656, 1988. Humes, W. H., Chem. Eng. Prog., February 1983, pp. 51 -54. Wagner, M. C., Humes, W. H., and Magnabosco, L. M., Plant/Operations Progress, 3 (4), 222, 1984. Martinez, N. P., et al., Paper presented at Ketjen Conference on FCC, May 1986. Kraemer, D. W., and deLasa, H. I., Ind. Eng. Chem. Res. 27, pp. 2002,1988. Schlossman, M., et al., Paper presented at Katalistiks FCC symposium, Venice, Italy, May 1986. Corella, J., Fernandez, A., and Vidal, J. M., Ind. Eng. Chem. Prod. Res. Dev., 25, pp. 554, 1986. Haunschild, W. M., Chessmore, D. O., and Spars, B. G., Paper presented at 79th AlChE Mtg. March, 1975, Houston, TX.

pp. 41 -45.

pp. 65-68, 1988.

1990, pp. 33.

3-5, 1989.

292

28 Rice, T., Carpenter, J. K., and Ackerman, C. D., I and EC, 46 (a), 1558, 1954. 29 Creighton, J. E., and Young, G. W., Paper presented at 8th N. Am. Cat. SOC.,

Philadelphia, May 1983. 30 Stripinis, V. J., ACS Symposium Preprints, Div. of Petrol Chem., Vol. 35 (4),

31 Young, G. W., Weatherbee, G. D., and Davey, S. W., Paper AM-88-52, NPRA Annual Meeting, San Antonio, TX, March 1988.

32 Young, G. W. and Weatherbee, G. D., Paper presented at the AlChE Annual Meeting, San Francisco, November 1989.

33 Zeton, Inc., 41 29 Harvester Road Burlington, Ontario L7L 5M3, Canada. 33a Intellution, Inc., 35 Perwal Street, Westwood, MA, U.S.A. 34 Pierce, E., Davison Catalagram No. 59, 1980 Davison Chemical Division of

W. R. Grace & Co.-Conn. 35 ASTM Proposed Methods of Test for Knock Characteristics of Motor Fuels,

2nd. Edition, Am. SOC. for Testing and Materials, Philadelphia, Pa, 19103, 1970.

36 Mitchell, M. M., and Moore, H. F., ACS Symposium Preprints, Div. of Petrol Chem., Vol. 33 (4), p. 547, 1988.

37 Hettinger, W. P., et al., Oil and Gas Journal, April 9, pp. 102- 11 1, 1984. 38 Xytel Corporation, 801 Business Center Drive, Mt. Prospect, IL, U.S.A.

p. 795, 1990.


293

CHAPTER 9

RESIDUAL FEED CRACKING CATALYSTS

MAURICE M. MITCHELL, JR.*, JAMES F. HOFFMAN, and HOWARD F. MOORE

Research and Development Department Ashland Petroleum Company

P.O. Box 391 Ashland, Kentucky 41 114

1. INTRODUCTION

The addition of residuum to FCC feedstocks dates to at least the 1950's. At that time its purpose was to add heat to the unit when required or as a means of disposal. In the 1970's two trends spurred the consideration of more extensive resid cracking and the development of processes to accomplish this more efficiently. One of these trends was the rapidly rising cost of crude oil and lagging product prices which were thinning margins severely. The second trend was that the average crude oil processed was becoming heavier; forecasts [l] have shown the world's crude reserves to be about 2:l in favor of heavy vs. light crudes. The consequence of these trends was that residuum would have to be processed to higher value added transportation fuels. Several processes were available for feedstocks of this type: hydrocracking, coking, solvent deasphalting and, of course, fluid catalytic cracking. It was viewed by several organizations that adapting fluid catalytic cracking to resid processing was the most efficient alternative. The others involved either high capital intensity, poor product quality or difficulty in disposing of the final rejected products (coke or pitch). In 1982 approximately 37% of the industry was practicing resid cracking [2]; by 1987 the number of resid cracking operations had increased to 39.8% of US refiners and 52.4% of Canadian refiners. Although a precise resid survey is not available, an NPRA report of FCC feedstocks reported that US refiners were processing resid in one form or another at 1520% of the total cracker feedstock in 1990 [3].

*Current address: Ohio University Southern Campus, Ironton, OH 45638

294

It is well known that residual fractions contain contaminant metals -- vanadium, nickel, some iron and, of course the sodium that inevitably escapes the desalters. These metals, particularly vanadium and nickel, lead to catalyst deactivation and catalyze undesirable non-selective reactions. Further, catalysts that were successfully employed for gas oil have little or no bottoms cracking capability. By this is meant that traditional FCC catalysts tend to make coke and gas instead of the desired transportation fuels from the most refractory hydrocarbons in the residual fractions, asphaltenes and polar compounds. That, coupled with the ravages of vanadium and nickel, created the challenge for the development of better catalyst systems for resid cracking. This challenge required the development of selective catalysts that were capable of handling the rigors of higher severity operation. Various traps and passivators, either incorporated with or mixed into the active portions, were required to negate the effects of the contaminant metals. This and the subsequent chapter are devoted to describing the characteristics and evaluation of these catalysts and additive systems for the successful selective catalytic cracking of feedstocks containing significant quantities of crude oil residuum.

2. RESID CRACKING PROCESSES

Addition of resid to FCC units has historically been tracked by equilibrium catalyst metals content 141. Traditionally a t low levels, these metals began to increase mildly after the Arab oil embargo followed by a 300%+ increase after the Iranian revolution. Refiners now routinely add residual feeds when economics dictate and where unit constraints allow. Many existing units are also being revamped to allow increased levels of residue addition. Of particular interest, however, are the four FCC systems available as grassroots resid units. Major competitors in this area include M.W. Kellogg, UOP, Stone and Webster, and Shell. A recent survey 151 estimates that almost 800,000 BPD of grassroots FCC units specifically designed for residual feeds will be on stream by 1994.

M.W. Kellogg offers the Heavy Oil Cracking (HOC) Process. This process was the first specific design for residue feeds, with the initial unit coming on line in 1961 at Phillips Petroleum's Borger refinery. original HOC units were side-by-side, but the latest designs are stacked. HOC units have been constructed with both internal (regenerator bed steam coils) and external catalyst coolers [61 . They are often designed to include antimony passivation, and use plug valve catalyst flow control. Several have been installed with residuum hydrodesulfurization (HDS) units for pretreatment, and HDS is recommended for residues with carbon and metals contents above about 10% and 30 ppm, respectively [7] . Particular attention is paid to decreased oil partial pressure, reducing catalyst contact time while improving contacting efficiency, and increasing operating temperature.

The

295

UOP offers the RCCB process, [81 developed in partnership with Ashland Oil, Inc. The first RCC unit was a modified FCC unit a t Louisville, KY, in the late 1970's, followed by units in Catlettsburg, Norway, and (under construction) Indonesia. The key elements of the process include two-stage regeneration, external catalyst coolers, lift gas catalyst contacting, and a vented riser to limit oiVcatalyst contact time. Originally designed for Arabian Light atmospheric residue, the Catlettsburg unit has successfully processed feeds as heavy as 15"API with Ramsbottom Carbon content of 8.5 wt% and total metals nearing 70 ppm nickel plus vanadium. Most RCC units operate with rare earth Y (REY) or rare earth exchanged ultrastable Y catalysts.

Stone and Webster offers a residue system developed in conjunction with Total Petroleum [91. The first of these units came on stream in Ardmore, OK, in 1981. Key elements of the process include advanced feed injection, two-stage regeneration, and careful design of the catalyst circulation equipment. Study of feed effects emphasized the importance of maintaining ring-opening rates significantly higher than dehydrogenation rates. As a result, Total appears to run relatively high regenerated catalyst temperatures using ultrastable hydrogen Y (USHY) catalysts [lo]. These units often use external second stage regenerator cyclones.

Shell began operation of their own proprietary residue design about 1988 in England, and are reported [5] to have three units on line'totaling about 120 MBPD as of 1992. The Shell system [ l l l boasts a proprietary lift pot and feed injection system, a compact reactor design, a staged catalyst stripper, a unique two-stage (single vessel) regenerator, catalyst coolers, and continuous catalyst additiodwithdrawal. Shell's designs are side-by-side units, and often include regenerator flue gas power recovery equipment. The latest units are designed for 6-7% Conradson Carbon and 20 ppm nickel plus vanadium with 5% and 10 ppm operations, respectively, common. Gasoline yields of 55 vol% on a 5% carbon feedstock are quoted.

All of these processes emphasize the challenges encountered when cracking residue. Increased heat loads, higher thermal instability of the feeds, increased contaminant metals, and decreased hydrogen availability dictate the catalyst requirements, including:

High thermal and hydrothermal stability. Enhanced metals tolerance. High gasoline selectivities. Minimum coke yield. Catalytic (as opposed to thermal) bottoms cracking.

As described above, there are several operating philosophies in use when charging resids. As a result, one specific catalyst may not (and probably will not) be optimum for all users. A residue catalyst must be tailored not only to the feedstock, but also the specific unit design and operating philosophy utilized. This is particularly emphasized by the opposing catalyst selection of Ashland and Total noted above. In addition, many refiners charge residue to conventional FCC units when unit constraints and economics dictate.

296

Catalyst selection in this case must consider both the conventional and the residual feeds.

3. CATALYST DESIGN AND COMPOSITION

The design of fluid catalysts for the selective cracking of resid- containing feeds must take into account the selective but active cracking of the 850"F+ material, the subsequent cracking of this material into transportation fuels, the avoidance at both stages of non-selective products (coke and gas), all in a hostile environment of contaminating metals which are deleterious to the catalyst and which in themselves catalyze non-selective reactions. It is necessary to have a balance between zeolite and matrix activities to not produce too much material in the vacuum gas oil (VGO) range relative to the amount of active zeolite present to selectively crack that material into transportation fuels. Too little matrix activity relative to the zeolite forces higher severity and as a result higher delta coke and gas. Thus the design requirement is to have the zeolite and matrix activity kinetically balanced and to retain that balance as much as possible as the catalyst deactivates. It is also necessary that there be accessibility to the active parts of the catalyst and this necessitates certain pore volume and pore size distribution characteristics. These concepts were developed by W. P. Hettinger, Jr. and his group at Ashland Petroleum Research and Development during the late 1970's and early 1980's [12,13]. There was equally important work being conducted elsewhere, in particular a t Filtrol (now Akzo), W. R. Grace and Engelhard. During the mid-1980's to date, much more has been learned about the details of these catalyst characteristics required for selective bottoms cracking and these are described below. Sufficient data have been assembled on the characteristics of many catalysts and their cracking behaviors such that it is possible to model and predict behavior of untested new catalysts based on physical and chemical characteristics.

3.1 Zeolite

Virtually all cracking catalysts contain, as the main active component, some form of Y zeolite. Recent advances in zeolite technology, along with the elimination of tetraalkyl lead compounds in gasoline, have resulted in a shift from rare earth and calcined rare earth exchanged Y zeolites to either hydrothermal (USY) or chemical dealumination (e.g. DY, LZ-210, etc.) of Y zeolites. The newer forms of Y zeolite generally exhibit a higher degree of hydrothermal stability, contain stronger acid sites, and produce higher octane gasoline with lower delta coke. However, with USY's the gasoline yields have been found to be lower than with conventional Y zeolites [lo].

In resid processing, octane is generally not a concern because of the high aromatic and olefin contents of gasoline produced with these feedstocks

297

[14]. Zeolite stability is a concern in resid cracking because of the high concentrations of vanadium in the feedstocks and the higher coke yields. Because of the importance of zeolite stability and the desire to maximize the yield of transportation fuels, a compromise in the optimum type of zeolite is needed for resid cracking. As will be shown in Section 3.6, a zeolite with a unit cell of 24.60A is desired for the optimal economic performance of resid cracking catalysts. Other reports 1101 suggest that USY catalysts are preferred because of their ability to produce low delta cokes. However, one advantage of using rare earth exchanged Y zeolites in resid cracking catalysts is the wider operating window, an example of which is shown in Figure 1. These results were obtained on a unit which was processing resid with a Conradson carbon greater than six. The operator of a typical resid unit which employs a REY catalyst has greater operating flexibility a t the gasoline knee without going into the overcracking regime.

Figure 1. Typical gasoline versus conversion curve for REY and USY catalysts ested in a pilot plant unit.

Gasoline vs. Conversion

52

50

G - 48 E3 $ 46 8

a,

2 44

42

40

1 I REY,

50 60 70 80 90 Vol.% Conversion

3.2 Matrix

Various forms of silica and alumina, along with kaolin clay, account for most of the matrix components used in cracking catalyst production today. Silica and clay are inert and for most of the early years of zeolitic cracking,

298

silica based catalysts were preferred. The zeolite provided essentially all the activity for the catalyst. However, some forms of alumina have been found to be active for catalytic precracking of large molecules [ E l . A schematic of precracking is shown in Figure 2, where the side chain of the large resid molecule is stripped from the aromatic moiety. The side chain is often further cracked on the matrix then enters the supercage of the zeolite where it undergoes selective catalytic cracking to produce transportation fuels.

Figure 2.

Initially the matrix in cracking catalyst was formulated to perform two functions: 1) act as a diluent for the active component, generally Y or USY zeolite, and 2) bind the different components which make up a catalyst and improve the mechanical strength of the microsphere [16,17]. More recently certain matrices have been found to have a third function: bottoms cracking activity. The matrix of a resid catalyst should have the following properties: acidity of sufficient strength to crack side chains, proper and selective pore size distribution, large pore volume, metals tolerance and metal immobilization capability, along with the typical properties and characteristics of a matrix in a VGO cracking catalyst [18]. An indication of the selectivity of bottoms cracking can be shown by evaluating the slurry (decant) oil yields, andor the light cycle oil to slurry oil ratio. Most gas oil molecules are of a size which fits through the large pore of Y zeolite (7.5A). However the molecules present in resid are too large to enter the large pore of the zeolite (see Figure 2). The presence of acidity in the matrix gives rise to selective catalytic cracking of the large resid molecules so that the fragments can selectively be

299

cracked within the supercage of the zeolite [19]. Selective bottoms cracking is defined as the conversion of resid into transportation fuels. Thermal cracking and non-selective bottoms cracking are distinguished by the production of larger quantities of dry gas and coke.

An important function of the matrix is to act as a sink for the contaminant metals which are in relatively high concentrations in residuum (20-500 ppm Ni+V). The control of the effects of metals is discussed in detail in the next chapter. However the use of rare earths in the matrix have been found to be effective in controlling the deleterious effect of vanadium; in particular, high lanthanum (La) to cerium (Ce) ratios have been found to be effective in controlling zeolite destruction. The lanthanum has been found to be most effective when precipitated onto the matrix as opposed to either impregnation or ion exchange [ 171.

3.3 ZeoliteMatrix Interaction

Experience with the Ashland RCC process has shown that the catalysts which give the best performance are those in which the zeolite and the matrix deactivate at about the same rate [17]. It is important that both components remain near the optimum relative activity so that the selective precracking activity of matrix remains in balance with the zeolitic activity. The gradient separation technique [20] has been used to characterize a number of catalysts in which the matrix and zeolite deactivate together are the better performers in the commercial process. In those catalysts in which the matrix deactivates rapidly or the zeolite is extremely stable, non-selective bottoms cracking occurs. The result is higher amounts of coke and gas produced in the unit.

3.4 Physical Properties

Matrix Acidity While there is no accepted or completely satisfactory method of

specifically measuring matrix acidity, there have been a variety of methods examined. One of the most common methods used for acidity is the adsorption of a base onto the acid site and measuring the amount of base by either thermal gravimetric analysis (TGA) or calorimetry. The major challenge in developing a method for matrix acidity is to be able to distinguish between the acid sites on the zeolite and those on the matrix. One proposed method for measuring matrix acidity has been to use a base of sufficient size which will not enter the supercage of the zeolite. A number of experiments have been done with tri-dodecyclamine with mixed results. The results of some of these experiments are shown in Table 1. It is known that the acidities of the matrices in the catalysts are significantly different based on the physical properties of the components used in their preparation.

300

Table 1. Acidity Results of Representative Catalysts

Catalyst Total Matrix Slurry Oil Yield*, Acidity (meq/g) Acidity (meq/g) wt.%

M 0.34 0.13 13.45 D 0.42 0.15 12.60 J 0.63 0.10 N/A A 0.36 0.00 15.20

* at 75% Conversion

Total acidity was measured in this set of experiments by n-butyl amine calorimetric titrations. Catalysts M & D, which produced low slurry oil yields in pilot plant tests (see Table 1) have been used in the RCC process. These two catalysts were also found to selectively crack residuum in the catalyst selection program which Ashland uses. The data in Table 1 suggest that some acidity is required in the matrix for selective bottoms cracking. The matrices of catalyst J & A are identical formulations. The small amount of matrix acidity found in Catalyst J is the result of a large amount of zeolite being present in the catalyst. The large probe molecule most likely adsorbed on the exterior surface of the zeolite. These results suggest that some amount of matrix acidity is required for bottoms cracking. Typically acidity of the matrix is increased by the addition of an active alumina during the preparation of the catalyst. Those catalysts with a matrix which consists of silica and clay have been found to exhibit matrix acidities similar to that of Catalyst A.

Unfortunately, the tridodecyl amine titration does not predict bottoms cracking selectivity for all catalyst systems, so this set of experiments is given as an example of how matrix acidity can be used to identify catalysts which can selectively crack resid molecules. The development of a method which correlates with bottoms cracking activity and selectivity is an area of research which is currently being vigorously addressed.

Pore Volume Catalysts that are effective for hydroprocessing heavy feedstocks must

have an adequate pore system to accommodate the large resid molecules. It has been shown that hydroprocessing catalysts must have diameter pores to allow asphaltenes access to the active sites [211. For this same reason, resid cracking catalysts with an adequate matrix pore structure should be more selective for bottoms cracking.

The importance of an optimal pore structure can be seen in Figure 3. Catalyst A is an excellent catalyst for cracking VGO, but has been shown to exhibit poor bottoms cracking. Catalysts B, C, and D have all been used successfully in the Ashland RCC process. All three catalysts crack bottoms selectively, but catalysts C and D produce more of the higher valued products gasoline and light cycle oil (LCO). These data suggest that resid cracking catalysts require a bimodal distribution of pore size, with maxima at pore

301

diameters of 400A and between 800-2000A. Those catalysts which produce the greatest amount of transportation fuels have the largest fraction of the pore size near lOOOA. Figure 4 compares slurry oil yield with virgin catalyst pore volume between 100-1000A, clearly showing the importance of having a significant amount of the total pore volume in this region. A minimum in slurry oil yield is obtained at pore volumes in this region greater than 0.15cdg.

z- 15 .- t- a t

Figure 3. Mercury Pore Volume

Catalyst

A c

B 0 c c

D I

100 1000 10000

Pore Diameter, Angetrcmns

Figure 4. Slurry Oil vs. Pore Volume

. . . . . . . . . .. . . 0 4 I

0 0.1 0 . 2 0 . 3 0 . 4 0 . 5

Virgin Catalyst ~ o o - ~ o o o A Pore Volume, cc/g

302

42

Maximum gasoline yield is also found in the same pore volume region that the minimum slurry oil yield was found (Figure 5).

. ~~ .

Figure 5. Gasoline v8. Pore Volume

56 T

54 I rn

r n . rn

rn . rn rn .

Total water pore volume of a catalyst has also been shown to relate to the performance of the catalyst. Similar curves are found when water pore volumes of virgin catalysts are correlated with slurry oil and gasoline yields. The minimum slurry oil yield and maximum gasoline yields occur at an equilibrium catalyst pore volume of 0.35 cc/g. A correlation has also been found between the virgin and equilibrium water pore volumes for most catalyst systems. The equilibrium pore volume has been found to be approximately 5/8 of the virgin pore volume (Figure 6). These data are from a combination of catalysts used in a commercial resid cracking unit and a large scale pilot or demonstration unit.

The fact that slurry oil and gasoline yield are found to correlate with pore volume supports the hypothesis that a significant pore volume is required to selectively crack the large molecules in a reduced crude feedstock. When the equilibrium catalyst pore volume is significantly less than 0.30cc/g, there is not enough volume within the catalyst particles to vaporize and selectively react the feedstock. With lower pore volume, catalysts produce higher coke and gas yields, and poorer transportation fuels selectivity.

303

0 . 5 ~

a, 5 0 . 4 5 -

2- 0 . 4 ~ ~

g> 0 . 3 5 .

-

m

ns 0 . 3 - .~ & 0 . 2 5 Ld

0 . 2

F i g u r e 6 .

Virgin vs. Equilibrium Catalyst - Water Pore Volume

** , I* , , C.’.

, + m * ,m- , ,

~.

, ~ , 3 **

RCCDU - - ~ Regression RCCCU - Actual Actual

3.5 COdCO Ratio

The cracking of residuum produces higher levels of coke on spent catalyst. Since catalytic cracking units operate in a heat balanced regime, i.e., the heat for the cracking reaction is supplied by the regeneration of the catalyst, the amount of heat produced in the regenerator affects the amount of feed that can be processed through the cracking unit. The amount of heat produced in the regenerator is affected by the carbon on the catalyst and the resulting CO2/CO ratio of the flue gas. High CO2/CO ratios can result, especially in units without feed preheat control, in reduced feed throughputs because of temperature limitation in the regenerator. Because the oxidation of carbon to carbon dioxide produces 3.5 times the heat from the oxidation of carbon to carbon monoxide, the control of ratio of the combustion products can improve the utilization of the cracking unit and increase the amount of feed processed.

A recent report [22] has shown that the CO2/CO ratio of the flue gas is dependent on the catalyst type and the level of contaminant metals, particularly nickel, on the catalyst. In this report a test was described which was able to predict which catalysts will produce high CO2/CO ratios. The major conclusion from this investigation was that the nickel catalyzes the combustion of carbon directly to C 0 2 , and the CO2/CO ratio is not controlled by an afterburning effect, as exhibited when platinum is used. More recent work (private communication) has shown that the nickel sites which are responsible for producing hydrogen and coke on the reactor side of a FCC unit are the same as those producing C 0 2 in the regenerator.

304

3.6 Inferences To Optimum Residuum Catalyst Composition Via Statistical Modeling Of Pilot Plant Results

An extensive data base of pilot plant performance results for 17 different catalysts (both in-situ and incorporation) from five vendors has been accumulated by the method described in Section 5. Statistical modeling of these results as a function of virgin catalyst properties has been used to provide insight into optimum resid catalyst properties. While the actual cracking reactions are a product of the equilibrium (rather than virgin) catalyst mixtures, we have modeled virgin catalyst properties so that performance of new catalysts can be predicted prior to andor as justification for experimental work. These modeling results can be used to prequalify catalyst candidates for full scale evaluations.

Detailed physical and chemical analyses utilized for this study are delineated in Table 2. Purely mechanical data such as attrition were judged to be unlikely to contribute to performance and have not been used. In order to further reduce the data set to a workable number of variables, only those characteristics a priori deemed important have been included. Since optimum (minimum and/or maximum) behavior can be expected from many (if not all) of these variables, squared terms were allowed for the major variables surface area, pore volume, matrix surface area, zeolite intensity, unit cell size, and rare earth content.

Table 2. Virgin Catalyst Properties for Modeling Efforts

Proaertv Surface Area, mVg Pore Volume, cdg Pore Volume Distribution, cdg

c60A 60-80 80-100 100-200 200-400 400-1000 1000-6000

Zeolite Index, % Unit Cell Size, A Alumina, % Rare Earth Oxides, % Magnesium Oxide, % Soda, ppm

Mean 202 0.35

0.05 0.02 0.01 0.03 0.04 0.07 0.08 12.2 24.63 40.0 1.81 2.24 4117

Minimum 157 0.25

0.01 0.00 0.00 0.00 0.01 0.03 0.00 7.4

24.53 29.9 0.44 0.00 1800

Maximum 252 0.50

0.09 0.04 0.04 0.11 0.13 0.17 0.15 19.3 24.72 49.3 3.84 10.7 6800

Dependent variables, Table 3, were defined as the pilot plant yield structure and total product value at (or near) maximum gasoline production when processing Arabian Light reduced crude. For RCC catalysts, this gasoline knee generally occurs at about 70 volume percent conversion, with some as high as 75 percent. Evaluation of these yield structures has

305

demonstrated that comparison of catalysts a t constant conversion causes fewer problems than allowing the conversion to float; further, total product value for a given catalyst does not appear to vary significantly over this range. As a result, all yield structures were selected at 70 volume percent conversion for consistency. Value Index is defined as total product value referenced to good (Value Index = 100) and poor (Value Index = 50) RCC catalysts.

Table 3. Major Yield Components Used in the RCC Model Yield

%&i Dry Gas, FOE % Wet Gas, Vol % Gasoline, Vol % Cycle Oil, Vol %

Coke, Wt % Value Index

Slurry, Vol %

-- Mean 3.5 20.4 49.2 15.1 14.9 13.2 75

Minimum Maximum 2.5 5.3 16.3 25.0 42.1 53.6 12.6 16.6 13.4 17.4 10.9 15.7 38 102

Statistical modeling was performed by SAS [23] for each dependent variable. Note that neither the second of the LCO/slurry pair (since conversion is fixed) nor value index are truly independent and therefore, represent a statistical weakness in the analysis. Models were developed using the stepwise MAXR technique, which sequentially develops the l-term model with the best correlation coefficient, then the best 2-term correlation, then 3- term, etc., up to (n-1) equation parameters where n equals the number of data points. This type of analysis is particularly informative in demonstrating the relative importance of each independent variable on the dependent variable of interest. The major weaknesses in the analysis are the limited number of data points (less than the total number of independent variables) and the obvious, but often overlooked, fact that these are not truly independent variables due to the constraints inherent in catalyst manufacturing techniques. In order to be statistically valid, this technique requires all variables to be independent. However, these results can be quite useful as long as these potential pitfalls are recognized.

Overall results are summarized in Table 4 and discussed below: Major factors are arbitrarily defined as the first 3-4 variables included in the model; minor factors are all other trends noted in the 16-variable (maximum) model. The simple inflection point is defined as the first time the linear and quadratic terms for a variable occur in the model with opposite signs. Finally, a + refers to a modeled increase in the dependent value with an increase in the independent variable.

w 0 Q)

TABLE 4. RELATIVE IMPORTANCE OF INDEPENDENT VARIABLES FROM DETAILED MODEL RESULTS

DN Gas W M i n o r

Surface Area Water Pore Volume

<60 60-80 80-100 100-200 200-400

400-1000 1000-6OOO

Zeolite Intensity Unit Cell Size + Alumina Rare Earth Oxides

Na20

Simplest Inflection Points:

MgO

Unit Cell Size 24.61 N Water Pore Volume >0.5 N Rare Earth Oxide 1.35 N Surface Area 214 X Zeolite Intensity

Wet Gas - M i n o r

+

+ +

+

+

+

24.62 N >0.5 N 1.88 N 216 N 18.8 X

+

+

+

24.59 X 0.40 X >3.0 X 250 N 9.4 N

+

+

+

+ + +

24.62 N 0.42 X

182 X 6.1 N

Value Index - M i n o r

+ + +

+

+

+ + +

24.66 X 24.60 X 0.38 X 0.18 N 2.29 N 4.5 x

176 X 21.4N

NOTE: N denotes yield minimum, X denotes yield maximum.

307

Surface Area On a simplistic basis, surface area tended to increase light products (dry

gas, wet gas, gasoline) at the expense of LCO and coke; it was a major positive factor in defining the value index. However, the predicted maximum value index occurred at only 176 mVg. Subjectively, the 176 mz/g optimum appears low, so present targets are recommended to be in the 180-200 m21g range.

Pore Volume As expected, total pore volume was consistently important in the

results. Generally increasing total pore volume improved all yield components except coke; coke tended to increase with total pore volume. Pore volumes above 0.4 were consistently best, although a maximum coke yield was predicted at 0.38.

Pore volume distribution results were intriguing and, to some degree, surprising. The modeled results preferred a bimodal distribution, possibly representing zeolite and matrix components. Maximum pore volume was called for in the c6OA and 200-4OOA ranges, and minimum elsewhere. This area certainly deserves more detailed study.

Zeolite Intensity Zeolite content was consistently important and positive to most yield

structures (as might be expected, it showed no impact on LCO production). However, conflicting objectives made optimization difficult - zeolite levels a t optimum gasoline and coke yields tended to also correlate with maximum wet gas. Even though the final results called for higher zeolite contents, about 15% zeolite intensity is suggested to minimize wet gas yields. This optimum will be very sensitive to the relative economics of MTBE feed, alkylation feed, and gasoline.

Unit Cell Size Unit cell size was unique in that, when allowed, the model always

wanted to produce curvature. For many of the other variables (in particular pore volume, surface area, and zeolite intensity), curvature was minor and/or occurred only late in the modeling sequence. The results are also fairly clear - most yields were optimized at a virgin unit cell size of 24.59-24.62. The only point of concern was an LCO minimum at 24.62, although as pointed out above, zeolite rarely if ever had any significant impact on LCO production.

Alumina Alumina content was clearly the single most important variable in this

modeling effort - it correlated with decreasing coke and wet gas plus increasing LCO, gasoline, and value index. The correlation between LCO and alumina was the best individual correlation obtained. We are confident that

308

this analysis confirms the need for a robust, active matrix when cracking resid.

Rare Earth High levels of rare earth were consistently called for by the model, since

rare earth showed a positive correlation with all yield components. Based on value index, 4.5% rare earth was predicted to be optimum. This high level was caused by a very strong response of gasoline in these calculations; maximum gasoline was predicted at 11% rare earth. Optimum levels for dry gas, wet gas, and LCO would be in the 1-2% range. Addition of higher rare earth content catalysts to the data base would be beneficial.

Magnesia Magnesium oxide was included in some formulations as a vanadium

trap. This analysis confirms the validity of this approach, and all yield elements are positive for increasing MgO content. However, MgO was not a major correlative factor in any of the yields, so its inclusion may represent only a minor factor.

Soda The trends for soda confirmed conventional wisdom. Minimum soda

content correlated with all improved yield elements as well as value index; soda should be minimized as much as possible within economic constraints. As with MgO, Na20 was not a major correlative factor with any of the yields.

Optimum Catalyst Properties The optimum virgin catalyst properties derived are shown in Table 5.

While it is doubtful that these precise values are necessary, they do demonstrate directions for future catalyst formulation.

Table 5. Derived Optimum Virgin Catalyst Properties

Surface Area Total Pore Volume

<60 60-80 80-100 100-200 200-400 400-1000 1000+

Zeolite Intensity Unit Cell Size Alumina Rare Earth Oxides

180-200 m2/g

Maximize 0.40-0.50 Cdg

Minimize Minimize Maximize Minimize

15% 24.59-24.62

>45% 3-4%

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Computer Models Because limitations in the data set restrict the utility of detailed

models, simple models have been developed for predictive purposes based on trends observed in the detailed analysis. Predictive variables were limited to surface area, pore volume, zeolite intensity, unit cell size, alumina, rare earth, and soda. Unit cell size was the only variable allowed curvature. Limitations in this analysis include:

Statistical correlation does not imply cause and effect. Cracking catalyst properties are not independent, and this internal correlation will bias the derived models. These models should not be used for extrapolation; their use should be limited to the stated range of experience. These models are based on global data, not detailed manufacturing information. For instance, total alumina is used without regard t o its location or activity. Catalysts from five different vendors have been used, without regard to manufacturing technique. There must be differences between, for example, the effect of total rare earth on an in situ formulation when compared to incorporation catalysts.

Each model has been verified by plotting predicted versus actual yields for gasoline and LCO, as shown by example in Figures 7 and 8, respectively. The scatter appears reasonable for modeling of very complex systems with such a limited variable set. Most plots are close to parity, although all show some deviation. The best model appears to be for LCO, with gasoline and coke showing more scatter than would be desired. These results suggest that additional catalyst properties are needed for better yield predictions.

Conclusions Statistical modeling of pilot RCC performance as a function of virgin

catalyst properties has been accomplished and validated. While several limitations apply to this analysis, and further model development is necessary, trends for future attention have been identified and a usable model for screening of new catalyst developments prepared. Optimum RCC catalyst properties were defined as moderate surface area and zeolite content, high pore volume with a bimodal distribution, intermediate unit cell size, and high alumina and rare earth contents. The analysis also confirmed that soda should be minimized and that magnesia addition can be beneficial in specific cases. Additional evaluations are recommended to improve these models, particularly by including catalysts which expand the existing range of catalyst properties. Development of matrix surface area and acidity for the existing catalyst data base are recommended, as well as more sophisticated parameters

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such as acidity distribution to improve model performance. Detailed study of pore size distribution effects is also an area deserving extensive further study.

Figure 7.

Prediction of Catalyst Performance

From Virgin Catalyst Properties

Predicted Gasoline, Vol % 54 I

42 ' , I I I , I 42 44 46 48 50 52 54

Observed Gasoline, Vol%

Figure 8.

Prediction of Catalyst Performance

From Virgin Catalyst Properties

Observed LCO, Vol %

3.7 Bottoms Cracking Additives

The addition of a non-zeolitic active matrix material to the catalyst formulation is an accepted method to improve catalyst performance when

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processing resid. There are many refiners however who do not process resid at all times and the amount and quantity of resid may also vary significantly day to day. Furthermore, some commercially available or advanced development catalysts have been marginal performers with regard to selective bottoms cracking. As a result it has been postulated that a separate matrix particle formulated for addition independently from the bulk catalyst injection would be beneficial. It was felt this route would allow for independent balancing of zeolite and matrix requirements as feeds, product needs, and operating practices change. There was a question about whether residence time limitations might prevent the required zeolite cracking of the heavy fractions produced from the matrix cracking of residue molecules from these separate particles.

A series of studies [241 utilizing fixed fluidized bed testing showed this concept to be feasible, and subsequent 1251, commercial trials have confirmed these laboratory results. The laboratory study showed the response to additive concentration to be non-linear, Figure 9, and the optimum concentration of additive in the total catalyst addition was about 10-20%. Subsequently, it has been shown [25] that by increasing the activity function of the bottoms cracking additive a 10% concentration of additive is sufficient. This improvement should minimize the activity dilution that was observed even for cracking catalysts containing significant amounts of zeolite. Some zeolite has been found to be necessary in the bottoms cracking additive to achieve optimum performance, but the zeolite concentration is significantly less than that on the cracking catalyst to which it is added.

Figure 9. Bottoms Cracking Additive Study --

Response to Additive level

Figure 10 demonstrates the improvement in gasoline yield of a purely VGO cracking catalyst. Its ability to crack bottoms to desired products approached the performance of a catalyst that was qualified and extensively

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52 5 0 - 48-

used commercially in an RCC process unit. Figure 11 further emphasizes the importance of this finding -- more selective cracking to gasoline has been attained concurrent with a reduction in coke yields. Physical and chemical analysis of these additives andor relation of these properties with additive performance have shown that zeolite content, surface area, and silica alumina ratio are the three most important factors affecting the efficiency of the additive.

- . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . .

. . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . .

. . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . .

I I

Figure 10. Bottoms Cracking Additive Effect

at the 20% Blend Level

Gasoline Yield, WPh 64

. . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . VGQ Catalyst Plus Additiue.

/. . . . . . . . .VGQCi

Figure 11. Bottoms Cracking Additive Effect

at the 20% Blend Level

6 1 . . . . . . . . . . . . . . . . . . . . .

70 75 80 85

>430'F Conversion, Wt%

-/- - - - - I Coke, Wt%

7

85 2 70 75 80

>430'F Conversion, Wt%

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By utilizing bottoms cracking additives both the refiner and the catalyst manufacturer have additional flexibility to meet product and operating requirements and the demands imposed by feedstock variations or changes.

4. Metals Management

The major metal contaminants accumulated on the catalyst are vanadium and nickel with some iron and sodium. FCC operations with about 10% resid added to the feedstock equilibrate metals to between 1000 and 3000 ppm nickel plus vanadium, while full reduced crude operations generally reach 7-10,000 ppm. The detailed effects of these metals on cracking operations are described in Chapter 10. It is important to note here, however, that vanadium has a strong destructive tendency toward the zeolite component. Nickel causes unwanted dehydrogenation reactions producing hydrogen and coke. Sodium is a general catalyst poison causing neutralization of acid sites and destruction of zeolite, but with a properly operating desalter the damage is much less than that done by vanadium. Iron porphyrins do exist in many crudes but for the most part iron in the feedstock is tramp or corrosion product iron. Its damage is usually no more severe than that of sodium. Iron is suspected, however, to contribute some dehydrogenation activity. These metals cause degradation of product slate value and increased operating costs to the refiner mainly because of increased fresh catalyst addition required to maintain unit activity. There are several process approaches to minimizing the effect of metals:

Addition of passivators - These are chemicals which react with sDecific metals to form compounds that are inactive. Metals traps - These are materials that are either incorporated in the catalyst formulation or are added as an additional component of the catalyst mixture whose characteristics are compound formation with specific metal contaminant. Use of a flushing catalyst - This is usually a low-metals equilibrium catalyst . Chemical demetallation - This is usually accomplished by total withdrawal of some of the equilibrium catalyst, chemically removing the metals and recycle of the catalyst. Physical processes such as magnetic separation in which only the highest metal fractions are removed from the equilibrium catalyst and the lower metals fraction is recycled to the cracking unit.

Metals passivation and traps are also discussed in depth in Chapter 10. The use of flushing catalysts is a straightforward dilution effect, and usually is a low-metals equilibrium catalyst or a very inexpensive simple FCC catalyst. The disadvantages of this technique are the cost and logistics of injecting the flushing catalyst and disposing of additional amounts of rejected equilibrium

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catalyst. The technical disadvantages are that the removal of equilibrium catalyst from the FCC unit entails removing a significant fraction of the active catalyst along with the flushing catalyst, plus these "flush" catalysts rarely have adequate matrix activity to crack residue. It is practiced commercially, however, and certainly involves a minimum investment which in some circumstances may offset the increased operating cost incurred.

4.1 Chemical Separation

As outlined above, chemical separation of the metals from cracking catalysts involves removal of some of the equilibrium catalyst, processing it through various schemes for metals removal, and recycling of a demetallated catalyst. Phillips Petroleum has a patent [261 on a process currently under development for deactivating spent metal-contaminated cracking catalyst by treating the catalyst with ammonium nitrate, then with a suitable compound, preferably ammonium fluoride, then with an antimony compound. According to the data presented in the patent, improvements of about up to 5% in conversion, 3-5% in gasoline yield, a reduction of 2-4% in light cycle oil, and 2- 3% heavy cycle oil were achieved along with a reduction of about 113 of the hydrogen make that was observed with the untreated catalyst.

The DEMET process now owned by Coastal Catalyst Technology, Inc., involves an initial calcination followed by sulfidation and chlorination. The chlorinated product is flushed to remove unused chlorine and cooled and washed. The demetallized catalyst is returned to the cracking unit and the filtrate is processed for the conversion of chlorides to metal hydroxide powders which are then available for sale presumably as high grade metal ores. This process is applicable to both on-site and off-site equilibrium FCC catalyst processing. In 1991 it was reported [27,281 that two extensive commercial runs had been completed, one of four months and one of nine months duration. Because of design constraints the original DEMET unit was restricted to a maximum of 92% nickel removal and 60% vanadium removal with catalyst containing several thousand ppm total metals. It is claimed that recent process improvements in the laboratory now allow 99.9% nickel, 80% vanadium, 80% sodium, and 90% iron and copper removals to be achieved. On site demetallization was also demonstrated 1291 at the Coastal Derby refinery in Wichita, Kansas. In a recent paper by F. J. Elvin [291 DEMET has proven to be a commercially viable method of reducing fresh catalyst requirements in residue cracking operations. In that same paper fresh catalyst savings of $0.20-$2.00/barrel were claimed. Yield improvement from lower catalyst metals of $0.10 to $0.50 per barrel were seen. Yield improvement from replacement of fresh catalysts in the inventory by demetallized catalysts is worth an additional $O.OS/barrel a t constant catalyst metals level. Increased residue processing capabilities from 8% 1000+"F residue to 24% residue as the percent of demetallized catalysts in the FCC inventory increase from 0 to 40%

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are claimed as an additional economic justification for use of the DEMET process. Passivating additives and their cost including royalty payments can be minimized if not eliminated. Clearly this process may hold great advantage for certain refiners.

4.2 Physical Separation

Chemical demetallization processes are designed to demetallate whole equilibrium catalysts and thus process a full spectrum of age and metals levels. The physical separation processes on the other hand are designed to remove the most highly metallated fraction, which will clearly be the most deactivated as well. Magnetic separation has been developed at both Ashland Petroleum Company and Nippon Oil Company Ltd. The Nippon Oil process [30,31] which operates in a carousel device was constructed and successfully operated for over a year in a 5000/bbl. per day FCC unit. It clearly demonstrated the feasibility of magnetic separation as an FCC unit catalyst inventory enhancement process.

The Ashland Petroleum process known as MagnaCatTM [32,33,34,35] in its current state of technology uses a belt and roller system (Figure 12). The spent catalyst from the regenerator is cooled and fed to a belt of the magnetic separation unit. The catalyst on the belt passes over a rare earth permanent magnet where the most magnetically susceptible catalyst is bound on the belt by magnetic forces. As the roller rotates the least magnetic catalyst is thrown away from the belt and the magnetic catalyst is retained until it passes the magnet, where it falls into a collection hopper. The process operates onsite on a slip stream of the equilibrium catalyst, and has been demonstrated to reduce hydrogen yield, maintain wet gas and gasoline, and reduce coke make. The following conclusions were drawn in a recent paper [351:

rn Equilibrium catalyst activity is raised about 2 MAT numbers a t constant catalyst addition rates. Significant reduction in hydrogen yields and H2ICH4 ratios are consistently obtained.

rn Reduction in delta coke allows either an increase in FCC feed or in incremental vacuum bottoms processing.

rn Improved catalyst fluidization characteristics may result from the use of the MagnaCat Process.

rn The process can be operated in different modes or a combination: rn Increase equilibrium activitylselectivity at constant catalyst

addition. rn Reduce fresh catalyst addition and spent catalyst disposal costs

a t constant equilibrium activity. rn A $0.20-0.45hbl. economic advantage is estimated for use of the

MagnaCat process for FCC applications.

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The extent of resid addition to the FCC feed would dictate the optimum operating mode in any given application.

Figure 12.

When processing reduced crude the utilization of magnetic separation can have a much greater effect on unit profitability than even the quoted numbers for FCC with vacuum bottoms addition in the feedstock. The advantage of this process is that only the oldest most heavily metallated catalyst is removed while the recycled catalyst contains the still active and selective fresher catalyst.

Whichever of the techniques, chemical or physical demetallation, there are clear advantages in certain circumstances, especially the extensive use of feeds containing large amounts of resid. The consequences of this practice on future catalyst design are not as yet well defined. However, the amount and type of metal traps, the advantages of using more stable zeolites and the

317

maintenance of matrix aciditylactivity will be involved in the determination of optimum catalyst design in the face of one or more of these processes.

In summary, the management of metals through demetallation either chemically or physically has been demonstrated to improve product slates and improve catalyst costs and to allow increased residue in feeds to cracking units. The demetallation processes may have an additional impact on the metals industries, particularly nickel recovery from those catalysts used to process the high nickel crudes from the far east.

5. TESTING AND SELECTION

Evaluation of resid cracking catalysts adds additional parameters to testing techniques which are often more art than science. The ability to handle high levels of contaminants (both carbonaceous and metallic) while maximizing liquid transportation fuel yields becomes the primary selection criteria. Catalytic cracking catalyst generally represents the major single catalyst cost for most refiners; with resid addition this cost increases dramatically, further emphasizing the critical nature of proper resid catalyst selection.

During development of the RCC@ process, catalyst testing protocols were reviewed, modified, and developed as needed [36,371. Comparison of these and literature results with commercial and demonstration-scale experience readily demonstrated that classical techniques gave poor and/or often erroneous results. These techniuues included:

Microactivity Testin? (MAT): These techniques are, and have been, industry standards; they are very well described in Chapter 7 of this volume. They have been shown, however, to have problems predicting bottoms cracking, coke selectivity, and performance with metals. In addition, conflicting results are often obtained when comparing MAT results from differing catalyst types (e.g., USY vs. REY, and/or catalysts from different vendors). While some MATS have been modified for use with residual feeds, most have not. Material balance problems are common when charging residues. Pilot Plant Evaluations: Pilot evaluations are generally valid, but are expensive and often time consuming. Equilibrium catalysts are generally required for testing, limiting the utility of this tool to catalysts already in the marketplace and precluding any control over their history. Finally, most pilot systems are incapable of charging residue feeds and therefore cannot simulate expected commercial environments. Commercial Trials: Commercial results are the desired product of any selection program. However, the cost of this mode of testing can be very high in the case of a catalyst failure; a 1% shift in yield from gasoline to slurry in today's market costs a 50,000 BPD FCC unit $4,00O/day. In addition to this high risk factor, consistent commercial operations can be

318

very dificult to attain due to changes in crude slates and other refinery requirements. It is very rare to obtain clean, comparable commercial data from two differing periods on the same catalyst, much less different ones.

Consideration of these factors plus the requirements of processing residue feeds led to definition of the following requirements for our testing program:

Utilize virgin catalyst samples.

Charge actual residue-containing feeds.

Minimize total testing cost.

Be correlatable to commercial operations.

Operate at realistic metals levels.

Minimize the probability of a commercial failure.

Based on these requirements, a testing protocol was developed which met all of these criteria. When a new catalyst is received, physical and chemical properties are measured and compared to those values we have defined as desirable. If these values are satisfactory, the catalyst then passes to a sequential, 3-phase testing protocol.

Phase I: The Phase I protocol is based on conventional MAT evaluations. Virgin cracking catalyst samples are steamed at varying severity and metals content followed by standard MAT testing. The major modification to the standard ASTM MAT is use of a heavier feedstock simulating a full- range vacuum gas oil refinery feed. Results from these tests are compared to reference standards from historical catalyst evaluations. The primary performance categories used for catalyst selection are activity, metals and hydrothermal stability, transportation fuel yields, and bottoms cracking ability. Development of this protocol demonstrated several weaknesses in the MAT procedures - poor commercial coke selectivity correlation, inconsistent bottoms cracking performance, and relatively narrow performance windows leading to limited resolution between similar catalyst types. Phase I evaluations do have the advantages, however, of rapid turnaround and limited total cost. Even considering these cost advantages, we no longer select candidates by Phase I evaluation due to the errors we have encountered during testing. We do, however, use this protocol to monitor day-to-day commercial operations. As discussed in Section 3.5, the amount of feed that can be processed through the cracking unit is often dictated by heat release in residue operations. The amount of heat produced in the regenerator is affected by the carbon on the catalyst and the resulting CO2/CO ratio of the flue gas. High CO2/CO ratios can result, especially in units without feed preheat control, in reduced feed throughputs because of temperature limitation in the regenerator. A test has been developed [22] which was able to predict which catalysts will produce high COz/CO ratios. We now routinely screen RCC catalyst candidates for CO2/CO ratio prior to further testing.

Phase 11: Phase I1 evaluations were developed to address these shortfalls in the Phase I data. Two tests were developed to correspond with

319

commercial operations - an FCC test, utilizing a catalyst with 3000 ppm nickel plus vanadium on an FCC feedstock containing 1.5% .Ramsbottom carbon content, and an RCC test which uses 8000 ppm nickel plus vanadium catalyst on an Arabian Light reduced crude with a 6.5% Ramsbottom carbon content. The simplified procedure is shown in Table 6. The objectives of these tests are to rapidly screen catalysts such that poor performers can be quickly and economically rejected while minimizing the probability that a good candidate is rejected.

Table 6. Simplified Phase I1 Catalyst Evaluation Procedure

Calcine a t 1100°F for 4 hours. Vacuum impregnate with vanadyl naphthenate and nickel octoate in cyclohexane:

Calcine in air at 1100°F for 4 hours. Steam deactivate a t conditions known to produce a target MAT activity on the reference catalyst

TestMode: FCC RCC Target MAT: 75 65 Temperature, OF: 1400 1375 Time, Hours: 5 4.75

* Atmosphere: -97% steam, 3% air- Fixed fluidized bed test at 960°F and 4.5 catalyst to oil ratio, standard conditions, after discarding the first cycle to allow catalyst activity to stabilize. Vary catalyst-to-oil ratio as necessary to provide a conversion range of 70-82 volume percent.

FCC: 1125 ppm nickel, 1875 ppm vanadium RCC: 3000 ppm nickel, 5000 ppm vanadium

Virgin catalyst samples are calcined, impregnated to the selected metals level, calcined, and steamed to produce a pseudoequilibrium catalyst. The impregnation technique is a modified Mitchell method [381 which results in a metals distribution and activity which are different than encountered commercially. As a screening tool for further testing, these weaknesses can be reconciled. However, if final catalyst selection is to be based on results at this testing level, cyclic metallation procedures are recommended [39]. We have implemented cyclic metallation techniques but find their costhime requirements to be difficult to justify.

The impregnated catalyst is steamed in a fluid-bed steamer at constant time and conditions. These parameters were defined in the initial test development as those which produce a typical, commercial activity level on our reference catalyst. This procedure results in variable activity levels, with activity a function of the steam and metals tolerance of the catalyst being tested. Of particular importance, air should be present during the steaming process to fully allow normal metals/steam/catalyst interactions. This is graphically demonstrated in Figure 13; steaming of a zero metals, virgin catalyst at 1400°F for seven hours resulted in a 25% surface area reduction; the same conditions with 0.5% impregnated vanadium reduced the final

320

surface area by 40%. The most striking response was steaming with 0.5% impregnated vanadium and 3% air - a 75% surface area reduction was observed. Other studies have shown the amount of air to be relatively unimportant. The significant factor is that, for full vanadium interaction to occur as it does commerciallv, oxygen must be present.

Figure 13.

Effect of Steaming Atmosphere

On Metals Impregnated Catalyst

Surface Area, sq. d g 250 I I

I 0 1 2 3 4 5 6 I a

Steaming Time, Hours

0 '

+ Steam t N2 ~ x - 0.5% V, Steam t N2 a 0.5% V, Steam + Air

Reference [361

Performance testing of the pseudoequilibrium catalyst is accomplished on a fixed-fluidized bed (FFB) system, Figure 14. This unit consists of a single reactor which is cycled through purge, feed, purge, and regeneration sequences. Liquid products are collected, the gas measured and sampled, and carbon content of the catalyst analyzed. Initial performance testing demonstrated relatively high metals activity, with modifications required to catalyst-to-oil ratios and provision of a "burn-in" cycle for the final protocol. Each catalyst is evaluated in a multipoint (2-3 conversions based on catalyst- to-oil variation) test resulting in curves of all major component yields versus conversion, which are compared to reference standards.

This test was validated by obtaining feedkatalyst samples from several of Ashland's cracking units, both FCC and RCC. These materials were evaluated, at a single severity, and each yield component compared to the commercial yield. Examples of these data are shown in Figures 15 and 16, without correction for conversion differences; good correspondence was noted for all yields.

Finally, the test was baselined by evaluating the best and poorest catalysts of their respective types by this protocol. As shown for example in Figure 17, a relatively broad performance window was defined to distinguish "bad" from "good" catalysts.

Overall, the Phase I1 protocol has proven to be markedly superior to Phase I due to (1) an improved performance window, (2) use of moderate (FCC)

321

CATALYTIC CRACKING PROCEDURE Laboratory Fixed-Fluidized Bed System

Process Feed for 3 Cycles: Purge

0 Reaction: Temperature 960 OF Residence Time 23 seconds Feed Weight 188 grams

Purge Regeneration -- 1100-1300"F

Mass Balance Each Test: Gas -- Gas Chromatography Liquid -- Simulated Distillation

0 Carbon -- Leco Octane -- Gas Chromatography

S T E A M V A P O R 1 2 E R

M E T E R

P R O D U C T C O O L A N T C O N D E N S E R

P R O D U C T R E C E I V E R

F I X E D F L U 1 D I Z E D B E D C A T A L Y T I C C R A C K 1 N G U N I T

FIGURE 14

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Figure 15.

Correlation of FFB and Commercial Gasoline Yields

C m m e r

0

F B

B %

: m m e r

i I %

40 45 50 55 Phase 11, Wt %

On Corresponding Feeds and Catalysts At Constant Conditions

60

Figure 16. Correlation of FFB and Commercial

Coke Yields

12 0

0

4 4 6 8 10 12

Phase 11, Wt 96 On Corresponding Feeds and Catalysts

At Constant Conditions

and high (RCC) bottoms containing feeds, (3) better definition of bottoms cracking performance, (4) reasonably reliable coke selectivity trends, and (5) good prediction of unit (versus MAT) catalyst activity. This protocol has proven to be an excellent screening tool, reliably weeding out poor performers; however, definition of true catalyst economics and in particular coke and gas selectivities remained relatively weak areas. This test also has passed several catalysts which were shown in Phase I11 to be poor performers. We now feel that this test primarily stresses the zeolitic component of the catalyst, and that major matrix contributions are not fully evaluated until Phase 111.

Phase 111: The Phase I11 protocol, circulating pilot plant evaluation, has been developed as our final qualifier. Catalyst candidates which pass the

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Figure 17. Phase II RCC Gasoline Yield

42

40

38

36

34

32

iasoline Yield, wt %

RCC Catalyst

VGO Catalyst

30 55 60 65 70 75 80

Conversion, wt %

Phase I1 screening steps are obtained in drum quantity, a pseudoequilibrium catalyst prepared, and full yield (selectivity) curves developed. FCC catalysts are tested at 3000 ppm Ni+V, and RCC catalysts a t 8000 ppm.

Preparation of a valid pseudoequilibrium catalyst is the key to proper testing of these materials. The most sophisticated performance tests in the world are worthless if the materials being tested do not represent commercial operations. Our experience has shown that, for good correlation to occur, the pseudoequilibrium catalyst must have an age distribution with a small proportion of fresh catalyst plus metals levels representative of the commercial operation being simulated, also with a representative age distribution. We have also found that the metallation and testing must be correlated with the specific commercial unit for which catalysts are being qualified, since operating practices can have significant impact on the success or failure of a particular catalyst formulation. However, since timeliness and cost are a significant factor in any evaluation, techniques must be used which limit the total time requirements for the testing cycle.

These factors have resulted in a multipart Phase I11 test (Table 7). The procedure begins by steaming several batches of virgin catalyst using 100% steam. Operations are continued (monitoring surface area hourly) until an 80 MAT activity is attained. This step was found to be necessary to produce representative MAT activity levels at the end of the metallation sequence. We have compared catalysts prepared with and without presteaming and find their selectivities to be identical. The lower activity catalyst is much easier to

324

Table 7 Phase I11 Catalyst Evaluation Simplified Procedure

Steam deactivate a t 1450°F to 80 4- 2 MAT activity. Metallate on circulating pilot unit using enriched reduced crude feed to 3000 ppm nickel and 5000 ppm vanadium.

1325-1350°F regenerator temperature.

FCC performance test on a fully circulating pilot unit:

RCC performance test on a fully circulating pilot unit:

Add 5 % steamed catalyst every 12 hours. Metallate to 3000 ppm total nickel + vanadium, retain half of the batch for FCC testing. Mix equal amounts of 3000 ppm metallated catalyst and fresh (steamed) catalyst. Continue metallation (with fresh catalyst addition) to a total of 8000 ppm nickel and vanadium.

Sweet domestic hydrotreated VGO with 10% vacuum bottoms. Vary temperature and catalyst-to-oil ratio to produce a minimum conversion range of 70-82%.

Arabian Light atmospheric reduced crude. Vary temperature and catalyst-to-oil ratio to produce a minimum conversion range of 70-82%.

test, and utilizes realistic conditions during the testing procedures. In one comparative case, an unsteamed sample produced a pseudoequilibrium catalyst activity of 79 MAT. We have also recently experienced catalysts which have zeolites with extreme steam stability. Because of our concerns about control, relative kinetics, and matrix effects a t very long steaming times, we have somewhat arbitrarily placed an upper time limit of 9 hours on the steaming step. Since as noted above we are confident that selectivity effects are relatively unaffected, we test these higher stability materials at the higher activity. Most of our conventional samples require 3%-6 hours to reach the 80 MAT activity level.

A large (nominal one barrel per day) circulating pilot plant is used for catalyst metallation and equilibration. This unit is equipped with full on-line reaction and regeneration systems. About 80 pounds of dry, steamed catalyst are charged to the unit. A sweet reduced crude (metals generally in the range of 2-5 ppm nickel and 3-10 ppm vanadium with a Ramsbottom carbon content of about 3-4) is used for ease of operation, with the metals increased to about 240 ppm nickel and 400 ppm vanadium by addition of nickel octoate and vanadium naphthenate. This combination was selected based on the desire to use a real feed, with real contaminants, but balanced by the need to build metals at a rapid rate. Normal operation is established at catalyst-to-oil ratios of 5-10 and regeneration temperatures targeted for 1300-1325°F. Five percent (5%) fresh (steamed) catalyst addition is practiced every 12 hours, and operations are controlled to produce an FCC sample (nominal 3000 ppm nickel + vanadium) in 36 hours.

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An RCC sample is prepared by blending equal amounts of the FCC catalyst prepared above with steamed virgin. This blend is charged to the large pilot unit, and metallation performed in a manner analogous to that used for the FCC catalyst. Metals levels in the feed are adjusted to provide 8000 ppm metals in 48 hours of operation, again while adding 5% fresh (steamed) catalyst every 12 hours. When the desired metals levels are reached, a material balance test is again performed prior to cooling and collection of the catalyst sample.

We have extensively evaluated these materials to determine their suitability for use. In general, compared to true commercial equilibrium samples, we find them to be stable, with representative levels of metals. Unit cell sizes are stable and consistent with commercial experience. We see no indication of preferential loss of any particular chemical constituent, in particular rare earths. We do see significant increases in iron content on the pseudoequilibrium samples. Since age distributions are a designed portion of this test, we have separated these materials according to the method of Palmer & Cornelius [20]. They have shown RCC catalysts to exhibit increasing nickel and vanadium contents with catalyst age (Figure 18). Our RCC pseudoequilibrium samples, Figure 19, show a similar trend but have a flat region in the mid-age range. This is probably due to the use of 50% 3000 ppm plus 50% steamed catalyst as the source for the RCC sample. While not completely simulating commercial operation, we feel that the age distribution demonstrated is adequate for testing purposes.

These pseudoequilibrium catalysts are then tested on a smaller, %- barrel per day fully circulating pilot FCC. The unit is computer controlled, with regeneration air set by offgas oxygen content and catalyst condition. Catalyst-to-oil ratios and temperatures are varied to produce at five or more material balances spanning at least the range of 70-82 volume percent conversion. Liquid products are accumulated through sequentially colder vessels, ending in an ethanoydry ice bath. Sequencing of temperatures, maximum contact surface area, and aerosol collection are critical for accurate measurements. The gas stream is reheated to ambient temperature before metering and analysis on an on-line gas chromatograph. After each material balance (generally six hours duration), liquid products are collected and batch fractionated. GC analysis allows quantification of light components collected in the gas, as well as the distillation properties of the product liquid.

We emphasize that the use of real feeds, particularly those containing vacuum bottoms, are critical to accurate testing. This was clearly demonstrated several years ago when we showed good performance for an alternate catalyst (Figures 20-21) - excellent gasoline yield and bottoms performance, very competitive with our (more expensive) preferred catalyst. However, we had performed these tests on a hydrotreated gas oil from a sweet domestic crude. We were somewhat uneasy with recommending this catalyst, since we knew this material to have little matrix function and since our operations routinely charge some residuum.

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Figure 18. Commercial Equilibrium Catalyst

Metals Distribution M e 1.6

1

6 I I

0 20 40 60 80 100 Cumulative Percent by Relative Density

0-.. %

Nickel -.s- Vanadium

Reference [201

Figure 19. Evaluation of Phase I11

Pseudoequilibrium Metals Distribution

M 11

a

0.2- I I I I

0 20 40 60 80 100

Cumulative Percent by Relative Density -- Nickel - Vanadium

t 0' %

327

Figure 20. Pilot Plant Performance Comparison

On Hydrotreated Gas Oil

Figure 21.

Retesting with a mild reduced crude showed a remarkable difference (Figures 22-23) - very degraded gasoline and bottoms cracking when compared to our standard. As a result we routinely use real feeds (Table 8) from our operations, even though these materials are more difficult to obtain, handle, and store. We also recommend periodic testing of these materials, as they are not storage stable particularly if heated; about two years' shelf life has been

328

45

40

our experience. Monitoring of carbon content and/or C7 insolubles has been of most value for our purposes.

-

-

18

14

10

"-60 65 70 75 80 85 >430°F Conversion, Wt%

-

~

-

I 1 I I I

Figure 23. Pilot Plant Gasoline Yield On Mild Reduced Crude

LCO, Wt.%

90

0

329

Table 8 Feedstock Properties

Identification FeedType CrudeType Source Date Obtained IDNumber RDBNumber

Feed Characterization Gravity,API Viscosity @ 210"F, CST Ramsbottom Carbon, Wt% BS&W,Wt% Pour Point, "F Heptane Insolubles

Distillation - D1160, "F

-

Volume % 5 10120 30140

50 60170 80190

94 Elemental Analysis

Sulfur,Wt% Total Nitrogen, WPPM Basic Nitrogen, WPPM

Metals, WPPM Nickel Vanadium Sodium Iron

HPLC, Wt% Saturates Mono-Aromatics Di-Aromatics >Di-Aromatics Polars Asphaltenes

Canton FCC Charge Super Sweet Crude FCC Charge Pump

ABL Reduced Crude #5 Crude Unit

2/89 2/2/89 PPFAO 115 PPFAO 114

RDB0381-RDB0410 PP-16

27.0 6.82 1.74

Trace +70 0.37

580 6261676 722/756

792 8331886

95311030 Cracked @lo50

0.16 698 245

67.9 17.2 3.0 8.8 3.2 NA

17.6 26.46 19.66

7.1 Trace

+20

604 664736 792/845

902 96ll997 96ll997

3.38 580 489

8 28

2 8

34.5 18.8 8.2

24.9 8.1 5.5

Another critical factor in successful pilot testing is correlation to the commercial operation for which tests are being performed. The operating philosophy for different FCC units vary widely, and different catalystsltesting protocols can produce significantly different results. Our test units and procedures are correlated commercially in three ways: (1) on identical catalysvfeed pairs, (2) by baselining with known good and poor performers, and (3) by post-audit of catalysts selected for commercial trial.

330

The original validation of our pilot cracking unit was by collection of actual refinery FCC feeds and equilibrium catalysts. These feedcatalyst pairs were then evaluated at a range of severities, and the yields of plotted versus commercial at the same conversion. These results are shown by example in Figures 24-25. Even though the commercial data are for multiple FCC plus RCC operations, the correlations are remarkably good - particularly when one factors in the less sophisticated balance techniques used during that time. An important additional point, however, is that the actual yields are not the same - all cracking units have their own unique yield structures, whether they be pilot, demonstration, or commercial units. As a result, comparisons should be made only on differential yields - the increment between a known standard and a new material.

Figure 24. Correlation of Phase I l l and Commercial

Gasoline Yields

Commercial, Vol % 65

. . . . . . . . . . . . . . . . . . 60

45 50 55 60 65 Pilot Plant, Vol %

On Corresponding Feeds and Catalysts At Comparable Conversions

65

60

55

50

45

Commercial, Vol %

-- . . . . . . . . . . . . . . . . . . . . .

45 50 55 60 65 Pilot Plant, Vol %

On Corresponding Feeds and Catalysts At Comparable Conversions

Figure 25. Correlation of Phase I l l and Commercial

Coke Yields

Commercial, Wt % 12

10 . . . . . . . . . . . . . . . . . .

0 0

. . . . . . . . . . . . . . . . . . . .

. . . . . .

A 0 I

I 4 6 8 10 12 14 16

Pilot Plant, Wt % On Corresponding Feeds and Catalysts

At Comparable Conversions

-t

33 1

Secondly, the overall protocol was validated (including pseudoequilibrium preparation) by testing known "good" and "poor1' performers on black oil from our commercial experience. Figure 26 summarizes these results. We demonstrated not only that we can distinguish between "good" and "bad", but we have also established the magnitude of the differences which provide the window of resolution between catalysts.

and new catalysts began to be evaluated in this program. Final proof of concept was provided by post-audit of commercial results from catalysts recommended by this program. FCC results showed good directional agreement when comparing differentials between the before/after results as shown in Table 9.

These demonstrations were used to demonstrate validity of our tests,

Table 9. Post Audit of Commercial vs. Pilot FCC Results Expressed as Differential Yields

Phase I11 Commercial Dry Gas, Wt% 0.0 +0.3

Gasoline, Vol% +1.8 +1.7 Wet Gas, Vol% -1.9 -0.5

LCO, Vol% -0.4 -1.6 Slurry, Vol% +0.4 -0.2 Coke, Wt% +0.3 +0.2

These commercial results are (as always) clouded by differences in feed, conversion, and catalyst conditions, but are as close as can be obtained. The value of obtaining a 1.7-1.8% gasoline yield improvement can certainly pay for a lot of catalyst testing.

Table 10 compares a similar post audit for RCC operations.

Table 10. Post Audit of Commercial vs. Pilot RCC Results Differential Yields

Phase I11 Commercial

Wet Gas, Vol% -0.6 +0.5 Dry Gas, Wt% -1.2

Gasoline, Vol% +0.8 +1.0 Cycle Oil, Vol% +0.8 +0.9 Coke, Wt% -0.3 -0.2

332

FIGURE 26. Phase I11 RCC Baseline Performance Windows Wet Gas Yield, wt %

25 I

VGO

10

5 ’

4E

40

35

30

iasoline Yield, wt %

50 60 70 80 90 50 60 70 80 90

Conversion, vol % Conversion, vol %

Cycle Oil Yield, wt % 20 I

15

10

Coke Yield, wt % 25 I

50 60 70 80 90 50 60 70 80 90

Conversion, vol % Conversion, vol % -

333

Again, remarkable agreement was obtained. Gas yields for the baseline commercial operation were invalid due to compressor problems during that material balance period. Overall, a total of five post audits have demonstrated excellent agreement for gasoline and coke in particular, with good agreement for bottoms cracking (LCO at constant conversion). Gas yields, particularly wet gas, have proven to be the most difficult to predict, although they have generally been reliable directionally.

Finally, the ultimate success of any testing protocol is its baseline - when in doubt, do it again! We have historically run a complete baseline (including new pseudoequilibrium preparation) for each new feed batch, and generally a t least once a year. This has allowed us to define our yield uncertainties to improve our discrimination between catalysts. It has also pointed out unit problems which were unknown at the time of evaluation.

Catalysts are selected based on these yield curves. For our purposes, yield componentsttrends are ranked in order of importance:

Gasoline t Coke I Bottoms Cracking I Increasing Wet Gas I Importance Octane I Dry Gas I

We have found that a single parameter is very advantageous when ranking one catalyst "better" than another. While any subject as complex as catalytic cracking can never be completely characterized by a single variable, we have found the concept of a "value index" to be quite useful. In this manner, different feed batches can be reconciled to a single set of values, and "good" performers are instantly recognizable. Figure 27 presents this concept in graphical form, including approved catalysts, failed catalysts, and catalysts which are being developed.

334

Figure 27. Value Index Comparison of RCC Catalyst Candidates

Value Index 120

. . . . . . . . .

. . . . . .

. . . . . .

. . . . . . bl . . . . . .

. . . . . . . . . . . . . . . . ~ ~ ~ .

. . . . . . . . . . . . . . . . . . . .

o Approved Catalysts m Other Catalysts

These procedures have now been used to successfully qualify four RCC catalysts and seven FCC catalysts (one of which was predicted to fail, and did). These results have ensured that Ashland utilizes the best catalysts available for our operations. Further, we have avoided frequent catalyst changes and the potential for degraded performance when the wrong catalyst is used. For example, consider the following evaluation of a catalyst recommended for use in our Reduced Crude Conversion (RCC) Process. All data are for catalysts metallated to a nominal 8000 ppm nickel plus vanadium.

In all of our tests, candidate catalysts are compared to a commercial RCC catalyst known for its excellent gasoline yield, bottoms conversion, and coke selectivity. As shown in Table 10, initial MAT evaluation of the metallated sample was promising.

Table 11. Phase I Evaluation, Jtem. Wt% Reference Candidate Dry Gas 2.6 2.2 Wet Gas 10.3 11.6 Gasoline 44.0 47.5 Cycle Oil 19.0 17.3 Slum 17.6 15.1 Coke 6.7 6.4 Conversion 63.5 67.7

335

These data were compared at constant condition, and suggested the potential for high gasoline yield, competitive bottoms conversion, and potentially improved coke and dry gas selectivities. As a result, the candidate was accepted into the RCC catalyst program for further testing. Unfortunately, the Phase I1 results were poorer than predicted by the MAT test.

Table 12. Phase I1 Evaluation.

Item. Wt% Dry Gas Wet Gas Gasoline Cycle Oil Slurry Coke Conversion

Reference 3.5

10.9 41.0 16.2 13.1 15.4 70.8

Candidate 3.6

10.0 40.3 15.0 12.4 18.6 72.6

These results (compared at constant conditions in Table 11) suggested performance, at best, equivalent to the reference rather than the hoped-for improved catalyst as predicted from the MAT results. In fact, this test showed the potential for poorer gasoline performance and, most significantly, much poorer coke selectivity.

Table 13. Phase I11 Evaluation.

Item. Wt% Reference Candidate Dry Gas 4.2 5.2 Wet Gas Gasoline Cycle Oil Slurry Coke Conversion

11.5 41.7 15.8 14.2 12.7

70

15.3 33.4 13.7 16.3 15.0

70

The candidate failed the Phase I11 evaluation. Very poor gas, gasoline, and coke selectivities were observed in pilot testing, with significantly impaired bottoms cracking performance. Different catalysts, and catalyst families, respond very differently to testing procedures; the use of MAT tests alone can lead to a significant number of plant test failures.

In conclusion, we now rely on Phase I11 evaluations for all final catalyst recommendations. We feel that any program of this type must include:

Use of real feeds. Correlation to commercial operations.

Preparation of a valid pseudoequilibrium catalyst.

Definition of performance measurement windows.

336

Periodic (frequent) baselines. We are convinced that these techniques are accurate, valid, and cost effective. These data suggest that our program has gained about 1% gasoline yield for Ashland during the period of this program.

Blind post-audit of actual results.

6. SUMMARY

In this chapter the characteristics and complexities of resid cracking catalysts have been discussed. Optimum resid cracking catalysts contain a balance of zeolite and matrix activity. Large pore volumes and bimodal pore size distributions are required to accommodate the large resid molecules into the acidic sites for selective bottoms cracking as opposed to non-selective thermal cracking. The interesting restriction of COz/CO ratio was discussed, particularly for those FCC units which may not have catalyst coolers. A model has been developed to relate virgin properties to performance in a predictive way to facilitate catalyst development and evaluation. In the model alumina content was the most important single variable, and the model called for high levels of rare earth and low soda concentrations. These undoubtedly reflect the need for acidity in the matrix to achieve the selected bottoms resid cracking. The use of bottoms cracking additives to improve marginal catalysts and provide flexibility both to the refiner and the catalyst manufacturer was shown to be a feasible approach.

Throughout all the discussions on the catalyst, however, was the concept of balance between the zeolite activity and the matrix activity. It is felt that this balance is important with the fresh catalyst and it is important to maintain as closely as possible the optimum balance as the catalyst ages. Relative to the zeolite, too little matrix activity would force higher severity to be practiced to achieve conversion and this would lower selective cracking leading to gas and coke. Too much matrix activity would also lead to gas, coke, or reduced throughput of the unit in order to achieve the desired product mix.

The effect of metals and metals poisoning were discussed with metals traps and passivators being deferred to the following chapter. Metals removal from catalysts through either chemical removal or magnetic separation was shown to be feasible in certain circumstances and can lead to lower catalyst cost and improved product slates.

The testing and selection of resid cracking catalysts was shown to be more difficult than VGO or HVGO testing. The protocols for preparation of the fresh catalyst candidates in any laboratory or pilot plant evaluation program are quite important. Any shortcuts or halfway measures are bound to mislead both qualitatively and quantitatively in the attempt to develop or select proper resid cracking catalysts.

337

There is still much work to be done. Invariably the processing of resid implies sulfur and metals disposal. Both additive and engineering approaches to solve the environmental implications will put additional demands on the catalyst formulation and manufacturing process. In the past several years truly outstanding cooperation has developed between refiners especially interested in resid cracking and the catalyst manufacturers. Solving future challenges will require a continuation of these fine efforts.

REFERENCES

1. Thompson, G. J., R. H. Hensen, C. N. Cabrera, and E. J. Houde, 12th World Petroleum Congress, Houston, John Wiley & Sons Ltd. (1987).

2. Thiel, P. G., Davison Catalagram, 1-12 (1983). 3. NPRA Survey of US Gasoline Quality and US Refining Industry Capacity

to Produce Reformulated Gasolines, Part A (1991). 4. Cotterman, D. W., Davison Catalagram #83 (1992). 5. Avidan, A. A., Oil & Gas Journal 90 (201, 59-67 (1992). 6. Wrench, R. E., and C. F. LeRoy, NPRA Annual Meeting, Paper AM-85-31,

(March 24-26, 1985). 7. Murphy, J . R., Third Annual Katalistiks FCC Symposium (May 26-27,

1982). 8. Shaffer, A. G., Jr., and C. L. Hemler, Oil & Gas Journal 88 (22), 62-69

(1990). 9. Santner, C. R., Stone & Webster Canada, Ltd., Technical Symposium,

Toronto (1988). lO.Santner, C. R., Hydrocarbon Processing 69 (12) 75-78 (1990). ll.Nieskens, M.J.P.C., F.H.H. Khouw, M.J.H. Borley, and K.H.W.

12.Beck, H. W., J. D. Carruthers, E. P. Cornelius, R. A. Kmecak, S. M. Kovach,

13. Hettinger, Jr., W. P., "Fluid Catalytic Cracking: Role in Modern Refining,"

14.Johnson, III., C. A., J. M. Kersey, H. F. Moore, and Mitchell, Jr., M. M.,

15.Humphries, A., and J . R. Wilcox, Oil & Gas Journal, 87 (61, 45-51,

16.Maselli, J. M., and A. W. Peters, Cataly. Rev.-Sci. Eng. 26 (3&4), 525-554

17.Beck, H. W., J. D. Carruthers, E. B. Cornelius, R. A. Kmecak, S. M.

18.Hayward, C.M.T., and W. S. Winkler, Hydrocarbon Processing &j (2) 55-56

19. Humphries, A., J. R. Wilcox, "Zeolite/Matrix Synergism in FCC Catalysts,"

Roebschlaeger, Oil & Gas Journal 88 (24), 37-44 (1990).

and W. P. Hettinger, Jr., US Patent 4,588,702 (1986).

M. L. Occelli, ed. ACS Symposium Series, vol. 375, p. 308ff (1989).

Div. Petr. Chem. Prepr. ACS 37 (3) 689-697 (1992).

February 6,1989,.

(1984).

Kovach, and W. P. Hettinger, Jr., US Patent 4,480,047 (1984).

(1990).

NPRA Annual Meeting Paper No. AM-88-71 (1988).

338

20.Palmer, J. L., and E. B. Cornelius, Applied Catalysis 35,217-235 (1987). 21.Riley, K. L., Div. Petr. Chem. Prepr. ACS 23 (3) 1104 (1978). 22.Doolin, P. K., J. F. Hoffman, and M. M. Mitchell, Jr., Applied Catalysis 11

23. SAS Institute, Inc., Cary, NC (1985). 24.Mitchel1, M. M., Jr., H. F. Moore, and T. L. Goolsby, AIChE Spring

National Meeting, Paper 60C (1990). 25.Ellison, T. W., Jr., E. J . Demmel, C. A. Steves, and C. R. Johnson, NPRA

Annual Meeting AM-93-52 (1993). 26.Lowery, R. E., C. M. Fu, and M. K. Maholland, US Patent 4,929,336 (1990). 27.Elvin, F. J. and S . K. Pavel, NPFUAnnual Meeting, AM-91-40 (1991). 28.Elvin, F. J., AIChE Annual Meeting, November 1991. 29. Elvin, F. J., AIChE Annual Meeting, November 1992. 30.Ushi0, M., US Patent 4,359,379 (1982). 31.Ushi0, M., US Patent 4,482,450 (1984). 32.Hettinger, W. P., Jr . and R. Benslay, US Patent 4,406,773 (1983). 33.Kowalczyk, D., R. J . Campagna, W. P. Hettinger, Jr., S . Takase, and M.

34.Hettinger, W. P., Jr., Catalysis Today, 13, 157-189 (1992). 35.Goolsby, Terry L., H. F. Moore, M. M. Mitchell, Jr., D. Kowalczyk, Warren

S . Letzsch, and R. J. Campagna, AIChE Spring National Meeting, Paper 64e (1993).

233-246 (1992).

Ushio, NPRA Annual Meeting, AM-91-51 (1991).

36.Hettinger, W. P., Jr. et.al., Oil & Gas Journal (1984). 37.Mitchel1, M. M., Jr., and H. F. Moore, ACS Div. of Petroleum Chemistry,

38.Mitchel1, B. R., I&EC-PRD, 19 (21, 209-213 (1980). 39.Haas, A., W. Suarez, and G. W. Young, AIChE Annual Meeting (1991).

Preprints 33 (4) 547 (1988).


339

CHAPTER 10

METALS PASSIVATION

RICHARD H. NIELSEN and PATRICIA K. DOOLIN

Research and Development Department Ashland Petroleum Company

P.O. Box 391 Ashland, Kentucky 41 114

1. INTRODUCTION

An important factor behind the growth of resid cracking discussed in C h a p t e r 9 is t h e development of heavy meta ls passivat ion technology. Compared to gas oils, resids usually contain high concentrations of heavy metals (nickel, vanadium, and iron) primarily in the form of porphyrin complexes and salts of organic acids. Under cracking conditions, metals, notably nickel and vanadium in gas oils as well as resids, deposit on the cracking catalyst and catalyze undesirable dehydrogenation reactions. High sodium levels poison acid sites of the cracking catalyst. Vanadium (and possibly sodium) under the conditions of the FCCU regenerator destroy the zeolitic component of the catalyst. Active metals reduce the yield of gasoline and increase the yields of hydrogen and coke. Since most cracking units can handle only limited amounts of hydrogen and coke, the level of active metals on the catalyst must be controlled in order to achieve maximum throughput and profit. Metals passivation is the process of mitigating the deleterious effects of contaminant metals thereby improving the catalyst activity and/or selectivity to more desired products. In the case of passivation of vanadium, the average life of the cracking catalyst particle is prolonged. With nickel, passivation decreases dehydrogenation activity. The refiner usually has several ways to utilize the benefits of reduced yields of hydrogen and coke achieved through metals passivation. One economically attractive option for many refiners has been t o crack resid blended with the traditional gas oil feed.

Since 1976, the successful commercial use of antimony passivation agents has provided refiners with a cost effective process requiring little capital and operating costs to manage mainly nickel. Bismuth and cerium compounds a re now also commercially successful. Since 1982 tin and, more recently, rare earth

340

compounds have been commercially successful in passivating vanadium. A number of other elements are now known to provide some degree of metals passivation.

Other ways to handle heavy metals in the catalytic cracker a re catalyst withdrawal and replacement (flushing) with lower metals equilibrium or fresh catalyst, optimizing process conditions to temper the effects of the metals, and using a more metals tolerant cracking catalyst. Catalyst flushing has been com- monplace in the United States and Europe creating a market for low metals equilibrium catalyst often used for replacement. However, using high catalyst replacement rates or process conditions is generally limited by economics and effectiveness to relatively low metals levels.

Even when hydrotreating the cracker feedstock, metals passivation is of ten beneficial , i.e. b y reducing t h e economically opt imum d e g r e e of hydrotreater severity. Operating the feed hydrotreater to meet sulfur specifications instead of low metals results in slower rate of contaminant metal build up on the HDS catalyst, extending the bed life.

While i t is well known that substantial benefits can be gained by passivating high metals containing catalyst, recent experience shows significant benefits may be realized a t low nickel loadings, about 500 ppm nickel, o r in some instances when hydrogen is produced a t only 40-55 S C F B F F [1,2]. However, each FCC unit requires detailed evaluation to determine the benefits of passivation.

Metals passivation is accomplished today in two ways: 1) metering the metals passivation agent, usually via the fresh feed, into the cracker where it deposits on the catalyst or 2) incorporating additives during manufacture into the catalyst formulation or into a separate carrier particle which is blended with the catalyst. Passivation additives are normally referred to as metal traps.

This chapter reviews the present knowledge of metals passivation of FCC catalysts. Useful information for the refinery engineer, the manager and the researcher is presented. Metals deactivation chemistry and laboratory and commercial passivation experience a re discussed, primarily for nickel and vanadium.

2. CHEMISTRY OF METALS DEACTIVATION

Both nickel and vanadium function as dehydrogenation catalysts a t FCC reactor conditions [3]. The dehydrogenation activity of vanadium is generally thought to be about one-fourth to one-fifth that of nickel [4]. Traditionally the relative dehydrogenation activity of contaminant metals is expressed in the passivation literature as a single parameter for metal concentration: four times the nickel level plus the vanadium concentration (4Ni+V).

34 1

Vanadium reacts destructively with the zeolitic componeiit of the cracking catalyst causing loss of crystallinity which is a more critical problem than its dehydrogenation activity. Nickel does not cause s t ructural damage to the zeolitic cracking catalyst, but significantly alters product selectivity to increase coke and gas yields. High hydrogen production reduces gasoline volumetric yield and limits compressor throughput. When determined separately, the relative activity of contaminant metals to degrade catalyst surface area are ranked as Ni < Fe < Na << V (Figure 1 [5] ) . Iron is usually associated with the catalyst primarily as a tramp metal. A small quantity is deposited by iron porphyrins which are present in the crude oil. With sodium removed by a properly working desalter, nickel and vanadium are the primary contaminant metals of a typical FCC feedstock.

CATALYST R

200 r

1450 F/SHours IRON

NICKEL

0 VANADIUM

I I I

0 5000 10,000 15,000 METAL, ppm

Figure 1. Metals degrade surface area. Reprinted with permission from [5]. Copyright 1983 American Chemical Society.

2.1 Nickel Under FCC reactor conditions, nickel is a dehydrogenation catalyst pro-

ducing high yields of hydrogen and coke. The ex ten t of dehydrogenation depends upon the nickel content, the age of the nickel and the cracking catalyst type. Although nickel is not considered to be a major factor causing loss of cata-

342

lyst activity, studies have shown that nickel containing catalysts produce more heavy cycle oil which is indicative of a reduced ability to crack the heavier feed components. As heavy oil conversion is normally associated with the catalyst matrix, this would imply that nickel poisons mild acid sites present on the matrix as well as stronger sites on the zeolite exterior surface. This explanation would be in agreement with the observed location of nickel (external to zeolite) on commercial catalyst using secondary ion mass spectroscopy (SIMS) [6]. Nickel has also been shown to be catalytically active in the regenerator producing high concentrations of carbon dioxide [71. As the oxidation of carbon to carbon dioxide produces approximately 3.5 times the heat of the oxidation of carbon to carbon monoxide, processing high nickel feed can result in reduced unit throughput.

As with other metals, commercial cracking catalysts vary in their sensitivity to nickel poisoning. Electron spectroscopy for chemical analysis (ESCA) studies show that nickel on equilibrium catalyst exists as Ni+2 and Ni+3 [6]. X- ray photoelectron spectroscopy (XPS) studies indicate that Ni interacts with clay and gel components of FCC cracking catalysts to form NiA1204 surface species. In steam-aged catalysts, silica is found to migrate to the surface where, in the presence of Ni, it forms inert NiSiO3-like species. The ability of a catalyst matrix (non-zeolitic component) to minimize nickel dispersion or of a clay to form inert nickel species will determine the Ni tolerance of the catalyst. The most active nickel species a re produced when nickel combines with alumina o r extraframework material present in modified zeolite [8].

Variation of nickel dehydrogenation activity was observed on different supports. Various laboratory techniques have been utilized to effectively study nickel interactions. Temperature programmed reduction (TPR) studies have been useful to investigate these different nickel species. Nickel on alumina was found to be difficult to reduce. The extent of reduction of nickel on alumina supported catalysts increased with increasing nickel loading 191 and increasing reduction temperature [ 101. Bartholomew and Pannell [91 reported 29% reduction for a 0.5% Ni on alumina and 75% reduction for a 9% Ni on alumina catalyst reduced a t 450". For Ni-silica catalysts the extent of reduction was usually higher than that for alumina-supported systems [9,11], due to a lower interaction with the support. The active nickel species producing hydrogen and coke was found to vary on the cracking catalysts based upon these factors. Nickel was generally more active on catalysts which contained alumina species which could interact with the nickel.

Imaging SIMS has shown tha t nickel tended to be immobilized once deposited on the catalyst particle [3]. Palmer and Cornelius 1121 correlated catalyst age with nickel content using data from equilibrium catalyst which they fractionated by gradient density separation. The amount of nickel deposited was found to be the product of the feed rate and the feed nickel content divided by the inventory over time in the unit. This relationship can be used to determine

343

the average age of catalyst in the unit by dividing the observed nickel content by its deposition rate. As nickel ages, it losses dehydrogenation activity and the required amount of passivation agent decreases [13].

2.2 Vanadium Some of the largest natural resources of vanadium are certain crudes from

Mexico and Venezuela. However, vanadium is present to some extent in virtually all crudes. Vanadium compounds in crude oil are primarily porphyrin complexes or napthenates. Napthenates decompose fully below 525-30°C; however, por- phorins decompose only after one-half hour at these conditions [14]. Cracking may o r may not be complete in the short contact time of the modern FCC riser. Whether or not complete decomposition occurs through riser cracking or by combustion in the regenerator, it is generally agreed that vanadium is deposited on the exterior of the catalyst particle due to the polar nature and size of the porphorin molecule. Since the coked catalyst is carried into the regenerator of the fluid catalytic cracker, a portion of vanadium present on the catalyst is oxidized to V+5. ESCA studies of equilibrium and metal-impregnated fresh catalysts show vanadium only in the +5 valence state [15]. Another study indicates that approximately 5% of the vanadium is present as VO+2 species on a steamed Y zeolite [15]. However, it is generally agreed that the primary species of vanadium is +5 after steaming. The vanadium oxidation state is independent of the vanadium source contained in the crude. As the cracking catalyst is repeatedly transported from the regenerator t o the reactor and back again, the vanadium continually undergoes valence changes between +5 and +4. Once formed in the regenerator, V+5 does not readily reduce to a +3 valence under normal fluid catalytic cracking reactor conditions [15].

Vanadium deposition on the cracking catalyst results in substantial loss of catalyst surface area and activity. As the zeolite component is the highest surface area component of the modern cracking catalyst, a decline in surface area is primarily associated with loss of zeolite crystallinity. Catalytic activity is effected similarly by contamination by sodium or vanadium (Figure 2 [5]) although caused by zeolite acid site poisoning versus zeolite destruction, respectively. Vanadium has been reported t o be less destructive to the zeolite in the presence of nickel [4,111.

Vanadium deposited on the catalyst exterior gradually migrates from the matrix surface to the zeolite crystal where it reacts destructively with the zeolite. The mechanism of this attack is a subject of considerable controversy. Several papers published in the early 1980's propose interaction of vanadium pentoxide (V2O5) with the zeolite to form a low-melting eutectic [14]. Vanadium pentoxide is known to have a low melting point, 690"C, which is lower than the average FCC regenerator temperature of 720°C. Hettinger and coworkers [5] clearly demonstrated that an oxidative atmosphere is necessary for zeolite destruction. Therefore, a V+5 species is usually assumed to be responsible for

344

R 20a

3 W cr; 1oc

C

\ NICKEL

SODIUM

VANADIUM \\

I I \\, 5000 10,000 METAL, ppm

Figure 2. Relative activity loss due to metals. Reprinted with permission from [5]. Copyright 1983 American Chemical Society.

the destruction of zeolite. Most of the damage to the zeolite is expected t o occur in the regenerator. Further evidence is that the destructive properties of vanadium can be mitigated by using hydrogen as a reducing agent at high temperatures [5]. These conditions are thought to reduce the V+5 species and, therefore, prevent its effect on the zeolite.

Vanadates of rare earths or aluminum were identified in studies conducted using physical mixtures of catalyst or catalyst components and vanadium pentoxide powder. The mixtures were calcined to high temperatures and characterized. Studies using laser Raman spectroscopy, XPS, and X-ray diffraction (XRD) on equilibrium catalyst showed a variety of phase changes occur with the destruction of zeolite by vanadium. Calcined rare-earth-exchanged Y (CREY) collapsed with the formation of cerium orthovanadate (CeV04) whereas HY formed mullite (AlgSi2013) and silica (tridymite) 1151. Mullite formation was also observed in steam-aged V-loaded gels but not when nickel was present. These studies led to theories of zeolite destruction by the formation of such compounds with abstracted atoms from the zeolite which lead to structural collapse. For example, Pompe [141 proposed the destruction of REY resulted when V2O5

345

attacked the rare earth component of the zeolite forming a low melting point RE-vanadate phase in which RE-ions were incorporated in varying proportions because of their chemical similarity. Although these compounds were formed, these studies did not establish whether their formation caused zeolite destruction o r whether they were simply formed subsequent t o the zeolite destruction.

These theories were later questioned by researchers who found that V2O5 did not cause zeolite structural damage in the absence of steam [16]. If sintering were the operative mechanism, destruction should have taken place in dry air. In the absence of steam, no zeolite destruction occurred. In 1986, Wormsbecher, Peters and Maselli proposed vanadic acid as the vanadium species responsible for zeolite destruction. The acid, H3V04, would be formed under FCC regenerator conditions by the reaction V2O5(s) + 3H20(v) --> 2H3VOq(v) [16]. This hypothesis incorporated both the oxygen and steam requirements. Since vanadic acid is a strong acid analogous to H3P04, acid attack of the zeolite via hydrolysis of the Si02/A1203 framework seemed plausible. The instability of zeolites to acid attack was well documented. However, this theory did not explain why catalysts which contain high sodium levels were even less vanadium tolerant than those with low levels. Sodium ions would be expected to have a neutralizing effect and to improve vanadium tolerance.

Vanadium was found t o be equally destructive whether added to the catalyst by napthenate impregnation o r by physical mixture of V2O5 powder. X-ray adsorption spectroscopy (XAS) studies found the vanadium adsorption edges were identical for steamed catalyst exposed to vanadium by impregnation or physical mixture, indicating the same oxidation state and coordination geometry [17]. Electron microprobe studies showed that after steam treatment vanadium was evenly distributed throughout the catalyst particle in each case. Wormsbecher contended that a volatile species must be responsible for a small amount of V2O5 powder to cause the same destruction as vanadium impregnation [16]. Liquid wetting or solid-state reaction could not account for deactivation by small amounts of V2O5 powder. To prove that zeolite destruction was caused by a volatile species, transport experiments were carried out in a flowing tube reactor. In these experiments, the zeolite containing catalyst was physically separated from a source of V2O5 powder. High temperature water was injected above the V2O5 in flowing air. Even though the vanadium source and the catalyst did not come into contact, after several hours the zeolite had completely lost crystallinity. Hence, the precursor for vanadium poisoning must involve H20 vapor and V2O5; the resulting species must be volatile. Compounds of vanadium with oxidation states lower than +5 were not considered as they did not exist at FCC regenerator conditions.

Recently Pine [18] studied vanadium destruction using a solid-state kinetics approach. He proposed that pentavalent vanadium simply served as a catalyst for the steam destruction of zeolite. The rate constants for crystallinity loss

346

were found to be directly proportional to the vanadium concentration. This would be true whether the role of vanadium was that of a reactant or of a catalyst. However, the fact that very small amounts of vanadium have a large effect on the reaction rate without being consumed was more consistent with a catalytic role. Pine extrapolated the rate constants obtained with vanadium to zero concentration and found agreement with rate constants taken without vanadium. This fact was consistent with the conclusion that the reaction was the well- known steam destruction of zeolite. To further understand the location of vanadium attack in the zeolite, rate constants were determined for silicalite, CREY and USY (ultrastable Y) in the presence and absence of vanadium. Silicalite was found to have a low vanadium tolerance. CREY and USY were found to have the same vanadium tolerance even though the CREY had almost 5 times as many framework aluminum atoms per unit cell. Based on these findings, the Si- OH bond was considered the more probable site of attack. This was consistent with the lower steam stability of a small particle Y zeolite which would have a high surface area to volume ratio. In addition, contrary to other studies [19], sodium and vanadium independently were found to have the same catalytic activity for steam destruction of zeolite, and together they acted synergistically. From the kinetic results Pine concluded both materials enhance the rate of reaction of steam with the zeolite. However, Pine did not explain the mechanism of this syngeristic effect.

At the time of this writing, a known mechanism of vanadium attack which explains all observed phenomena is not available. Although current knowledge is inadequate from an academic perspective, the factors which lead to vanadium deactivation of cracking catalysts such as oxidation state of the vanadium, and the presence of steam and high temperature are clearly defined.

3. FCC PASSIVATION ADDITIVES

Numerous strategies to deal with the deleterious effects of metals, primarily nickel and vanadium, have been developed. These include hydrotreat- ment to remove metals from the resid FCC feed [20], operational changes to alter the oxidation states of metals [5 ] , and passivation agents (for list of suppliers see [Zl]). The use of metal passivation has become an established practice. A passivation agent is a compound which can be utilized in an FCC unit under normal operating conditions. Passivation additives can also include metals traps or scavengers which are mixed with the catalyst or compounds which a re incorporated in the catalyst during manufacture.

4. NICKEL PASSIVATION AGENTS

Nickel passivation agents are normally injected into the FCC feedstock to react with the contaminated catalyst. Although a large number of elements a re

347

claimed in the patent literature as effective agents for nickel passivation (as discussed in section 6 of this chapter), only antimony, bismuth and cerium based compounds have been utilized commercially. Compounds are available with the active ingredients in an organic solvent or an aqueous solvent.

4.1 Antimony Research focused on heavy oil cracking and contaminant metals on crack-

ing catalyst led Phillips Petroleum Company in the late sixties to mid-seventies to the discovery of several metals passivation agents. Antimony containing compounds discovered by Marvin M. Johnson and Donald C. Tabler consistently were outstanding metals passivation agents [1,22]. The fist commercially used additive was an oil-soluble compound containing antimony, phosphorus, and sulfur in a hydrocarbon solvent developed by Phillips Petroleum Company called Phil-Ad@ CA. The active compound was antimony trisdipropyldithio-phosphate [23]. The antimony content was typically 10.5 t o 12.5 wt.%; sulfur, 17.5 wt%, and phosphorus, 7.5 wt.% minimum.

Tests in bench scale, semi-batch, micro confined fluidized bed units [24] demonstrated large decreases in hydrogen and coke yields accompanied by corresponding increases in gasoline yield. The relationship with antimony concentration at a constant metals loading was nonlinear (Figure 3 1251). A similar nonlinear relationship was observed in the hydrogen yield variation with antimo- nyhickel ratio in commercial tests [l]. Pilot plant transfer line reactor tests [25] confirmed the bench scale yield results and that multiple cracking-regeneration cycles could be run without a significant decrease in passivation benefits.

In a series of laboratory experiments with passivation agents impregnated on equilibrium catalyst, antimony trisdipropyldithio-phosphate was compared with triphenyl antimony, antimony trithallate, and colloidal antimony pentoxide dispersed in a hydrocarbon. Although passivation was observed with each compound, the antimony trisdipropyldithio-phosphate produced significantly more gasoline and less hydrogen and coke than the other compounds [23]. This suggested sulfur or phosphorous or both improved the passivation. While this enhanced passivation is small, pilot plant studies have shown it to be economically significant.

An organo-antimony compound, antimony tricarboxylate, and a colloidal dispersion of antimony pentoxide in a hydrocarbon-based solvent [26] were compared in a circulating pilot plant and three commercial units. In the pilot plant antimony was cracked onto an equilibrium catalyst containing 1000 ppm nickel and 3100 ppm vanadium with gas oil. No significant differences in conversion or yields of gasoline, hydrogen or coke were found between the antimony tricarboxylate and the colloidal antimony pentoxide. However, the laydown efficiency was found to favor the antimony tricarboxylate which started with an efficiency near 100% and declined to about 50% within 50 hours while the colloidal antimo-

348

-40 0 0.1 02 0.3 0.4 05

Antimony on catalyst, wt. %

Figure 3. Antimony passivates metals on FCC catalyst. Data for 75 vol.% conversion of West Texas topped crude on equilibrium catalyst from HOC at 95OOF. Reprinted by permission from the September 1977 issue of HYDROCARBON PROCESSING, page 97, copyright Gulf Publishing Co., 1977, all rights reserved.

ny started with about 80% efficiency and declined to about 40% during the same time period.

In commercial practice water-based antimony agents were found to be as effective as hydrocarbon-based agents [271.

Despite the fact that, from an industrial point of view, antimony passivation of nickel is a well-known process, the influence of accompanying elements (co-catalysts) is less well understood. A recent laboratory study of antimony passivation with and without sulfur andor phosphorus was conducted [28]. XPS analysis found a decrease in the surface nickel atoms in those samples passivated with antimony compounds containing sulfur and phosphorus with respect to the unpassivated catalyst. This was attributed to the fact that antimony complexes containing sulfur and phosphorus were more active for forming Ni-Sb alloys than antimonyhulfur or antimony alone. Nevertheless, the passivation of nickel was slightly more effective with the antimony organo complex containing only sulfur. At low levels of antimony, no differences in passivation were observed between the complexes studied. Therefore, the role of sulfur and phosphorus as co-catalysts for antimony passivation of nickel remains unknown.

The Research Institute of Petroleum Processing, SINOPEC (RIPP) has developed an antimony-containing additive 1291, MP-25 [30], comparable in per-

349

formance to an imported additive in a test at the Jiujiang Refinery. Before testing MP-25, the Luoyang Refinery tested an oil soluble passivator with low antimony content, MP-85, and found the hydrogen yield decreased 35%) the coke yield decreased, and the yield of light cycle oil increased slightly [31]. Based on the successful use of MP-85, higher antimony containing MP-25, which was soluble in alcohols and esters but not hydrocarbon solvents, was tested [31]. Care was taken to avoid exposing the MP-25 t o air. The MP-25 was charged at a high rate initially for four days and then set to the maintenance rate t o build the antimony concentration from 500 ppm to 1500 ppm. The antimony laydown efficiency was 79%. The hydrogen content of the dry gas decreased after only one day of passivator injection. Comparing the before and after periods of MP-25 use, the hydrogen content of the dry gas decreased 38% from 51.3 to 32%. The' yield of coke decreased slighty and the yield of slurry oil ("oil paste") decreased approximately 1%.

4.1.1 Antimony-Nickel Interactions Since the advent of commercial use of antimony additives, the interaction

of antimony with nickel has been the subject of detailed studies. Dreiling and Schaffer [32] examined catalysts having nickel loadings in the range of 1.9-4.4% in weight and Sb:Ni ratios varying from 0.0 to 0.41. From XRD results the authors suggested the formation of Ni-Sb solid solutions with a high level of Sb present on the nickel surface. Geometric and electronic effects were invoked to explain the results.

Parks et al. [33,341, working with different types of Ni and Ni-Sb on catalyst at high levels, suggested the formation of an alloy. Hydrogen chemisorption on nickel was effectively poisoned by the presence of antimony. XPS showed that both antimony and nickel were present on surface sites. Three types of nickel and two types of antimony were found. On cracking catalyst with high metal levels, the antimony forms were: 1) a non-reducible antimony oxide, probably existing as a mixed metal oxide catalyst, 2) a reducible species, well dispersed on the catalyst and 3) reducible antimony which forms an alloy with nickel upon reduction. In view of their findings, the following was proposed: a) geometric blocking of nickel sites by the antimony present on the catalyst, b) alter- ation of the electronic properties of Ni surface atoms by the presence of Sb in such a way that their catalytic activity was significantly reduced and c) that the amount of antimony available to passivate the nickel is determined by the equilibrium between antimony interacting with the support and with nickel.

Goldwasser [9] studied the effect of antimony addition on the structure and chemisorption properties of nickel. The addition of antimony substantially reduced the chemisorption properties of nickel. Complete reduction to metallic nickel and metallic antimony was found for the Ni- and Sb-rich samples. Nickel increased coke and hydrogen yields in isooctane cracking. The presence of antimony reduced the amounts of coke and hydrogen produced by nickel. Site block-

350

age of nickel and weakening of the Ni-C bond by the addition of antimony was suggested to explain the results. Electronic effects were proposed to explain the strong Ni-Sb interaction. Due to these effects the back bonding capacity of the nickel was reduced by the presence of antimony, producing a weakened Ni-C bond strength, thereby decreasing the amount of chemisorbed carbon monoxide.

4.1.2 Antimony Commercial Experience The first commercially used additive was Phil-Ad@ CA whose active

ingredient was oil-soluble. The additive was injected into a compatible diluent carrier stream such as a light cycle oil stream and pumped to the fresh feed line t o the FCC riser. The first antimony passivation plant trials were conducted at Phillips Petroleum Company's Borger Refinery Heavy Oil Cracker (HOC) in June, 1976. At that time, HOC capacity was 24,000 B/D. A second catalytic cracker unit in the refinery was a 30,000 B/D FCC unit cracking primarily gas oil. The two units were interconnected through slurry oil recycle of oil and catalyst fines. The metals on the HOC catalyst were about 1.5 times those on the FCC catalyst [25].

The overall test plan was to operate the unit normally but to make three base tests for a short-term comparison of conversion and product yields. Base test conditions were developed jointly by Research and Development and Refinery Process Engineering personnel. A base test was conducted at controlled feed rates, riser, and regenerator temperatures. In addition to the base tests, computer and operator panel board readings of the normal unit operation were logged for long term monitoring.

The Phil-Ad@ CA antimony additive was metered from a small tank into the inlet of the HOC fresh feed pump. The additive was initially injected a t a high rate (Figure 4 1251) to build up the additive concentration on the catalyst. After two days, the rate was reduced to a maintenance rate and held. Six days later a second series of base tests were made. The antimony on the catalyst was then increased to the second level and a third set of base tests were conducted,

The HOC unit's response t o the Phil-Ad@ CA injection resulted in a decline in the hydrogen yield as noted in the gas compressor speed readings. During the first 24 hours, hydrogen yield decreased from 300 to 120 SCF/B. Within a few hours, a measurable increase in raw gasoline yield was reported. Within two days, the regenerator temperature started decreasing indicating less coke make.

35 1

June July

Figure 4. Phil-Ad@ CA injection. Reprinted by permission from the September 1977 issue of HYDROCARBON PROCESSING, page 99, copyright Gulf Publishing Co., 1977, all rights reserved.

The short-term, base test results were summarized in Figure 5 1251 as a function of conversion of fresh feed. The incremental increase in gasoline yield was less at higher reactor temperatures; 6.1% improvement at the lower temperature and 2.7% a t the higher temperature. On the same basis, the hydrogen yield decreased 46.6% and 38.6%, dry gas decreased 27.5% and 21.7%, and coke decreased 15.3% and 16.7%. The yield of isobutane increased while the C3 and C4 olefin yields were not significantly changed. Operation of the fractionator influenced the split of light and heavy cycle so the effect of passivation on these streams was inconclusive. These data indicated the optimum operating conditions changed when the unit was passivated due to new heat and material balance steady states and unit restraints.

An extended evaluation (Figure 6 1351) made under typical refinery operating conditions confirmed the base test results. The percent change of the mean values for the ten month period just before passivation was compared with the ten month period following. The HOC unit experienced a 20.4% increase in gasoline while producing 9.1% less coke and 61.2% less hydrogen. During the passivated period, the fresh feed rate increased 11.9% while the conversion increased 5.4%. Before passivation, the HOC unit was limited by air blower and gas compressor capacities (Figure 7 [25]). By reducing the HOC coke yield by 15%, passivation allowed charging 15% more long residuum feed without reaching a gas compressor limit. Later, operation a t a higher catalyst metals level (23,500 ppm Ni+V vs. 17,000 ppm) was demonstrated at an increased fresh feed rate (Figure 8 [35]).

The FCC equilibrium catalyst was also passivated a t the time of the HOC unit passivation [25,36]. The FCC unit equilibrium catalyst contained about two- thirds the metals level of the HOC catalyst. The passivation of the FCC unit

352

I I I I I I

3501 I I I I I I

5.0 Phil-ad CA wmsivsu

4.0 - Barn

3.0. I I I I I

-

=..a 221 I

71 72 73 74 75 76 77 Conversion, lv. % fresh feed

Figure 5. Base test results show passivation benefits. Reprinted by permission from the September 1977 issue of HYDROCARBON PROCESSING, page 100, copyright Gulf Publishing Co., 1977, all rights reserved.

353

HYDROGEN YIELD S CFB

CONVERTED

METALS ON EQUIL. CAT.

4 N I + V , ppm x 10-3

PHIL-AD CA RELATIVE INJECTION

RATE

HYDROGEN YIELD 1 300

100

- 1 I

:::I1, I I I , I I --rj PHIL-AD CA RATE

S N J M M J S N J M M J S N J O D F A J A O D F A J A O D

0

1975 1976 1977

Figure 6. Extended evaluation confirmed base tests. Reprinted by permission from [35].

was accomplished at times using antimony on the catalyst in the HOC slurry oil [361, as well as by direct injection of Phil-Ad@ CA. During the passivation period, the unit charged 5.9% more fresh feed which contained 22.5% more long residuum than during the prior ten month period. The unit produced 9.3% more gasoline, 10.2% less coke and 27.8% less hydrogen. Comparing the two units, greater metals passivation benefits were obtained in the unit with the higher metals on the catalyst.

Overall the tests demonstrated that metals passivation improved selectivity with higher conversion, increased feed throughput, and allowed the utilization of heavier feedstocks.

Since the first commercial application, antimony metals passivation has been applied successfully t o a variety of catalytic cracker unit designs including Thermofor Catalytic Cracker Units, to gas oils and resids from a large range of crudes including hydrotreated feeds, and to a variety of types of catalysts a t a wide range of metals levels [l] as discussed below. Additional benefits were also demonstrated: a large increase in C4 production through more severe cracking conditions, and the maximization of distillate production at low conversion oper-

354

FRESH FEED, 27,000 BPD

26,000

25,000

24,000

3 1,000

- AIR BLOWER -

- 4 NI + V = 17,000 ppm

- I

29,oooc

27,000

26,000

25,000

\

Air blower before

Phil-ad CA injection

I 1

64 66 68 70 72 74 76

Conversion, lv. %

Figure 7. Passivation allows thruput and conversion increases [25]. Reprinted by permission from the September 1977 issue of HYDROCARBON PROCESS- ING, page 100, copyright Gulf Publishing Co., all rights reserved.

Figure 8. Increased Phil-Ad@ CA injection allowed higher metals on catalyst. Reprinted by permission from [35].

355

ation [l]. Optimization of metals passivation benefits is discussed by Bohmer, et al. [l].

Antimony material balances showed that antimony was retained on the catalyst, normally concentrated on the fines [l]. Extensive sampling during unit operation and during turnarounds has detected negligible antimony dissolved in the liquid hydrocarbon products or sour water [1,37]. Antimony was not detected in the ambient air samples collected in the regenerator during turnarounds [l]. Stibine was not detected in welding plume samples. Deterioration or embrittlement o r cracking of steels were not detected during metallurgical examinations. As a result of information gathered a t over seventy refineries, the use of antimony for metals passivation in an FCC unit was found to be a safe and acceptable practice.

In addition, antimony usage efficiencies can be increased by recycling catalyst fines back to the unit. Catalyst particles from an FCCU treated with antimony serve as a solid passivation additive [36]. Both fines in the slurry oil and the regenerated fines are effective.

4.1.3 Antimony Effects At High/Low Nickel Loadings Antimony passivation is effective at low nickel levels (<6000 ppm 4Ni+V)

as shown in seven commercial experiences [l]. The relative metal levels of the equilibrium catalyst in these tests ranges from 2100 to 5000 ppm 4Ni+V. Hydrogen yield decreases by 26 t o 40 vol.%, even in two units producing only 40- 55 SCFBFF without passivation. ! b o plant tests are briefly discussed below.

The first FCCU's catalyst contained 900 ppm nickel and 600 ppm vanadium for 4200 ppm 4Ni+V [1,21. A significant production of hydrogen caused the wet gas compressor to be operating a t its limit. Passivation with antimony reduced the yield of hydrogen by 31%. The yield of coke decreased 5% and the yield of gasoline increased 1.5%. With the wet gas compressor unloaded, 2.5% of the gas oil fresh feed was replaced with lower value resid. The cost benefit ratio of the improvements was greater than 1:30.

The second example [1,2] of passivation a t low nickel level was a unit whose catalyst contained 3,160 ppm 4Ni+V with only 490 ppm nickel and 1200 ppm vanadium. The unit operated against both its air blower and process gas compressor limits. The 92 SCFiB of hydrogen caused difficulty in maintaining the governor on control as well as affecting the heat control of fuel gas users. When passivated with antimony (Table 1 [2]), the hydrogen production dropped 37% to 58 SCFB, unloading the compressor, allowing the governor to function and stabilizing the fuel gas composition. Coke production was also reduced as indicated by a 17'F regenerator bed temperature reduction. The decreased compressor loading was utilized by increasing the fresh feed from 31,000 B/D to 33,000 B/D with a poorer quality feed. The metals on the catalyst increased with

356

Table 1 Antimony Passivation at Low Nickel Level [2].

After Before Passivation

Charge, B/D 30,774 32,943 Feed API Gravity 28.2 27.5 Feed Carbon Residue, wt.% 0.38 0.63 Regenerator Bed, "F 1,302 1,285 Metals on Catalyst, ppm

Nickel 489 611 Vanadium 1,203 1,585

Hydrogen, SCFBFF C2 & Lighter, LBBFF C3 & C4 Olefins, LV%FF

Gasoline, LV%FF LCO, LV%FF Coke, L B B F F

1c4, LV%FF

92 58 14.3 14.7 12.8 13.2

3.1 3.1 56.2 54.3* 24.2 26.5 19.7 18.3

* Separations conditions were changed to increase yield of LCO during the late fall and winter months.

Reprinted by permission from [2]. Copyright 1991 American Chemical Society.

the higher metals concentration in the feed. The yield of gasoline was lower due t o a seasonal end point adjustment.

Nickel passivation is particularly important when processing Chinese and other Pacific Rim crudes. For example, some major Chinese crudes are very high in nickel content and low in vanadium [29]. A Shengli atmospheric resid contains 36.5 ppm nickel but only 0.1 ppm vanadium. The percentage decrease in hydrogen production in commercial units is independent of the catalyst Ni/V ratio. Phillips' laboratory studies showed antimony interacts with vanadium to reduce its dehydrogenation activity [l].

High nickel-containing catalysts were effectively passivated with antimony added to a pilot plant feed. The test was conducted by the Research Institute of Petroleum Processing, SINOPEC. An antimony concentration of 1000 ppm on 3800 ppm nickel on a 6000 ppm nickel catalysts resulted in a one half to one third reduction in the yield of hydrogen and a significant reduction in the yield of coke (Figure 9 [29]>. The rate of increase of hydrogen and coke yields slowed down to a low rate above 6000 ppm nickel. However, the hydrogen-to-methane molar ratio continued to rise at about a constant rate. An 11,400 ppm catalyst required

357

i t , , l l l 0 40

0 1000 3000 5000 Sb on Catalyst, ppm

Figure 9. Effect of antimony on high nickel catalysts [29]. Reprinted by permission.

higher antimony levels (up to 4000 ppm Sb) and the extent of reduction in hydrogen and coke yields was significant, but not as great as with the lower nickel catalysts. Probably due to partial surface covering with antimony and antimony- nickel alloy, the MAT activity of the 11,400 ppm Ni catalyst decreased by 7 units with antimony addition. The pilot plant's yield of gasoline was still rising with the antimony level, however. A loss of catalyst surface area and activity was also noted with 12,000 ppm nickel plus vanadium [38], but indications were that the vanadium was mainly the cause and that the nickel was not effective in reducing the surface area.

4.1.4 Antimony Effect On CO Combustion Promoters Carbon monoxide combustion promoters, usually particulate substrates

containing a low concentration of platinum, are added in small quantities to the FCC catalyst inventory in order to attain increased or complete combustion of CO to C02. The effect of antimony or bismuth [39] on combustion promoters appears t o be promoter and unit dependent [39,401. More severe deactivation results for platinum-palladium containing promoters since palladium is more sensitive to antimony than is platinum [41]. Metals-free catalyst containing 0.5 wt.% promoter of unreported platinum concentration was not affected by 0.28 wt.% antimony impregnated onto the catalyst-promoter mixture [42].

358

While no significant change was observed in some units, in others, the steady state CO promoter efficiency decreased as much as 60% [39]. Decreased efficiency was especially likely during the initial injection period at high antimony addition rates [43]. Doubling the rate of combustion promoter at those times was recommended [41]. At the lower maintenance rates of passivator addition, the combustion efficiency was often restored to the prepassivation level by optimizing the antimony rate to avoid an excess of passivator. However, some cases required up to double the rate of promoter [39,40]. The effect seemed to be more pronounced with lower concentrations of metals on the equilibrium catalyst.

Perhaps the most widely held theory is that passivation of the platinum promoter by excess antimony (or bismuth) occurs and results in lower catalytic activity for oxidation of CO. While nickel and vanadium both promote the oxidation of CO to C02, they are orders of magnitude less active than platinum [44] but the lower activity is partially compensated because they are present in concentrations considerably higher than platinum. Moreover, in a study of regeneration properties of metal contaminated cracking catalysts, nickel is shown to be a very active catalyst for the direct conversion of carbon to C02 [7]. Antimony is found to effectively passivate the catalytic effect of nickel in the regenerator as well as in the reactor [7]. Therefore, if a cracker is utilizing antimony passivation, lower concentrations of C02 are expected during the catalyst regeneration, especially for high nickel concentrations on the catalyst. Addition rates of CO promoters above the pre-passivated optimum promoter level may then be required. Excess antimony above some concentration may promote CO combustion has also been suggested [421. The use of antimony has not affected the efficiency of SOX reduction catalysts in units with complete CO combustion [l].

4.1.5 Evaluation Of Passivation Benefits Since the reduction in the yield of hydrogen is the most pronounced fea-

ture of metals passivation with antimony seen in commercial units, the selectivity improvement in coke and gasoline are related to the hydrogen level expected for a given metals level [45,46,471. Having determined the selectivity benefits, the utilization of the unloaded air blower and process gas compressor through increased feed rate, cracking lower quality feed, or increasing conversion can be estimated through heat and material balance calculations [45,461. The latter two benefits require commercial or pilot plant relationships for the yields of feed blend components such as the gas oil and resid or for the conversion and product yields as a function of cracker severity (i.e., catalyst-to-oil ratio or temperature). Published examples are available [45,46,47].

A comparison of pilot plant, predicted and commercial results of passivation for three cases [l] showed good agreement for hydrogen reduction. The estimations of the changes in coke yield were acceptable. Gasoline increases were hard to measure in commercial units due to changes in operating conditions. Overall, the pilot plant and predicted values were good estimates of the

359

magnitude of commercial benefits obtained.

Modeling to predict passivation effects was discussed by Bohmer, McKay and Knopp [2]. Modeling was recommended for evaluating commercial test data that yielded unexpected results due to feedstock or operational changes which may have masked the passivation effects.

4.2 Bismuth Bismuth (and manganese) compounds are reported as effective nickel pas-

sivating agents developed by Gulf(Chevron) [48]. Little has been published about the chemistry of the interaction of these materials with nickel. Varying reports are given of the commercial benefits of the additive compared to antimony and the Betz DimetalIicB 9P2 [13,27,40]. The main benefits of bismuth are it is less toxic than antimony and it currently is not listed by the United States Environmental Protection Agency as a hazardous chemical. The volatility and leachability of bismuth from the cracking catalyst is reported to be less than antimony [49]. At the time of this writing, one-half dozen or more FCC units in the United States are using CMP-112@ bismuth additive marketed by Intercat [Sol.

4.2.1 Bismuth Commercial Experience A commercial t e s t of a bismuth additive by Chevron Research and

Technology Company was reported [49] for a FCCU which had not previously been passivated. The testing procedure was to charge a "small amount" of resid a few days before the start of bismuth injection, to inject the bismuth additive for 30 days, to stop the passivation additive and to allow the bismuth concentration on the catalyst to decay.

Due to the resid feedstock, the nickel content of the catalyst rose steadily during the test from just less than 1000 ppm to about 1450 ppm (Figure 10 [49]) over sixty days. Without passivation, a steady increase in the yield of hydrogen was expected. With bismuth injection, the yield of hydrogen trended lower (Figure lo), to values (52-62 SCF/B) lower than experienced with the initial nickel level (59-74 SCF/B). After about two weeks, the bismuth to nickel concentration reached a point of maximum passivation benefit. With the further increase in catalyst nickel, the ensuing rise in the yield of hydrogen was still a t a lower rate than expected without passivation. When the bismuth addition ceased, the yield of hydrogen rapidly increased to values (72-81 SCFD3) higher than before the passivation period due to the higher catalyst nickel content. Following an evaluation of the test results and a decay period, the refiner resumed bismuth addition.

Long term data obtained on another FCC unit showed the decrease in the yield of hydrogen as a function of nickel at two levels of bismuth (Figure 11 [49]). In agreement with bench scale and pilot plant results, incremental benefit from

360

-20-10 0 10 20 30 40 50

E 16001

1500-

2 1400- 8 1300-

c)

; 1200- 2 1000- t:

Bismuth

Bimnuth

Bismuth

1

u -20 -10 0 10 20 30 40 50 60 Time, Days

Figure 10. Commercial Bismuth Test. Reprinted by permission [49].

Catalyst Nickel: ppm

Low Bimmuth

+ 0" C*tniy.t

Emtimated Unpuivptsd

...+...

0

Figure 11. Effect of Bismuth Level. Reprinted by permission 1491.

361

low to high bismuth levels was smaller than the incremental benefit of the low bismuth level compared t o no passivation.

At Mapco Petroleum Company's Memphis Refinery, Chevron and Mapco conducted a commercial test of bismuth immediately following antimony passivation [51]. A water-based antimony passivation agent was used prior t o switching to bismuth in the form of Intercat's CMP-112@ additive. The injection system was cleaned or replaced before loading the liquid additive. Like the antimony additive, the bismuth additive was metered into a feedline of the FCCU. Due to the high catalyst turnover rate and the high feed nickel content, CMP-112@ was injected at a higher build-up rate for several days and then reduced to a maintenance rate.

In this test immediate results comparing antimony and bismuth were difficult due to the variation of feedstock quality over the time required to purge the unit of the antimony containing catalyst. The crude runs from a previous period with antimony were similar to the crude processed during the bismuth period. The fraction of equilibrium catalyst in the make-up trended from about 30% at the start of the antimony period to 100% by the end of the bismuth period. Despite the increased usage of equilibrium catalyst during the bismuth period, the nickel was significantly newer (Table 2 [51]). Fresher nickel has more dehydrogenation activity, but expected hydrogen yields could not be estimated. No difference could be found in the catalyst MAT activity between

Table 2 Effect of Catalyst Makeup Rate Change on Catalyst and Nickel Age Profiles 1511.

Percent of nickel in inventory with aee less t han t

Percent of catalyst inventory If total If total with age less than catalyst catalyst

t Ni = 2000 ppm Ni = 3000 ppm Bi Sb Bi Sb Bi Sb

fdavs Period Period Period Period Period Period

7 60.9 51.7 54.5 48.0 56.5 49.2 14 84.7 76.7 75.8 71.2 78.7 73.0 21 94.0 88.7 84.1 82.4 86.6 83.3

Reprinted by permission.

362

I-( 1401 Dailyhydrogen ~

4

k . 3 20: W Antimony period

A Bismuth period O 0 ' 140 . Test-run hydrogen

-Regression, all data g 0 120;

g 20: u o

1,500 2,500 3,500 4,500 5,500 Catalyst nickel, ppm

Figure 12. Effect of Bi and Sb on hydrogen yield. Reprinted by permission [51].

the antimony and bismuth periods. Periodically test runs were made in addition t o the normal, daily data unit logging of results. The deposition efficiency of the water based antimony additive was 36-44%. The laydown efficiency of the Intercat CMP-112@ was about 50% during the initial high dosing period but increased to about 65% at the steady state maintenance rate.

Comparison of the results of the test runs as a function of the nickel content for both elements showed equivalent results for the yield of hydrogen (Figure 12 [51]). The yield was reported to be fairly independent of the changes in feed and operating conditions. The yield of coke (Figure 13) and the ratio of coke yield to the feed Conradson carbon residue was said to compare well for the two passivation agents. Considering data scatter and unit changes, differences in the conversion and yield of gasoline were undetectable.

363

5.0 5.8 5.6

5.2 % 5.0

; 5.4

u

5.2 ,Test-run coke I L . -

A - - . . 1 . . - - A

. H Antimonypriod

4.8

Coker per feed ~ - R s g r a * i o n ’ a l l & L . I 3.0rConcarbon . A

A Lmuthpriod

1.81 ,. I

1,500 2,500 3,500 4,500 5, 3 1 I I 1 I

Catalyst nickel, ppm 00

Figure 13. Effect of Bi and Sb on coke. Reprinted by permission [51].

4.3 Betz Nickel Passivator Betz Process Chemicals, Inc. has developed and marketed a proprietary

non-antimony containing additive Dimetallica 9P2, formally called DM-1152, for the passivation of nickel on cracking catalysts. The active ingredients were considered to have low toxicity [62]. This additive was commercially tested successfully in 1988 at the FCCU at the Coastal Eagle Point Oil Company refinery [52] following antimony pretreatment. This nominal 50,000 B/D unit had a catalyst inventory of about 450 tons. No purchased equilibrium catalyst was used during the trial and fresh catalyst additions were varied little. A change in catalyst suppliers occurred about 75 days from the start of Dimetallica. The feed rate and other unit conditions were relatively constant during the Betz treatment period.

Dimetallica was first injected for 10 days at twice the anticipated maintenance rate and then reduced to 1.5 times the rate for five days before setting the anticipated maintenance rate. After a line-out period, the rate was periodically adjusted based on unit performance and averaged about 12% less than the target with deviations usually to lower rates. The catalyst analyses indicated that the antimony concentration declined prior to the Dimetallic maintenance level, and was essentially below detection levels three weeks later.

364

During the Dimetallic additive test, the riser outlet temperature averaged 956°F overall with three roughly 50-75 day subperiods averaging about 962"F, 950"F, and 956°F. Excursions were generally less than 10 degrees for any period of time. The start of the 950°F riser subperiod corresponded to the switch in catalyst. During the whole DM-1152 evaluation period, the regenerator temperature ranged from 1300°F to 1400"F, averaging 1345°F which was typical for this unit. The catalyst-to-oil ratio varied as typical between 4 and 6, averaging about 5.0 during the antimony and DM-1152 periods. Generally ranging from 60 to 66 vol.%, the equilibrium catalyst MAT activity was unchanged during the antimony or Dimetallic periods and was unaffected by the switch in catalyst.

The yield of hydrogen (60-100 SCF/B) remained unchanged varying with the active metals levels on the catalyst. It averaged slightly below 80 S C F B with antimony, Dimetallic, and the change in catalyst supplier. The unpassivated yield of hydrogen was estimated to be about 150 SCF/B for the level of active metals on the catalyst.

4.4 Cerium Certain cerium compounds were found unexpectedly to be effective passi-

vators for nickel as well as for vanadium [53]. The compounds included both cer- ous and ceric oxidation states from an array of organic and inorganic anions. Cerium was claimed to be less toxic than antimony. Water or organic solvents solubilized or suspended the compounds. The agent was injected into the cracker's fresh feed stream. Patent claims of the cerium levels on the catalyst ranged from 0.005-240 ppm or an atomic Ce:Ni ratio of 0.05:l to 1:l. The mechanism of the cerium-nickel interaction was not discussed.

Betz obtained commercial success with both antimony and cerium based passivation agents [40]. Strangely, both antimony and cerium agents performed equally well in some units, but in other units antimony was more advantageous while in still others cerium was the more effective.

5. VANADIUM PASSIVATION AGENTS

Vanadium is the most damaging of the contaminant metals to the cracking catalyst resulting in high catalyst replacement costs to the refiner. For this reason, research efforts have been extended to develop an effective passivation agent for vanadium. Although a number of materials have been cited in the literature for vanadium passivation, only tin additives are commercially available. Data from laboratory studies or short commercial trials are available for oil-soluble titanium and rare earth compounds. These vanadium passivation agents dissolved in a solvent are injected into the cracker feed stream to react with the incoming vanadium before i t can destroy the zeolite.

365

5.1 Tin Since catalyst activity appears t o monotonically decrease with increasing

vanadium levels due to zeolite destruction, passivation of vanadium to prevent catalyst damage is of prime commercial interest. While the passivation of vanadium has not been as successful as nickel, tin does reduce the deleterious effects of vanadium contaminants [49,54]. Vanadium destroys the zeolitic component of the catalyst. Chevron claims the rate of zeolite destruction can be decreased by tin. While the interaction of tin with vanadium is thought to be rapid, the major benefits are not observable until the vanadium damaged zeolite becomes depleted from the unit inventory. Although data on the effectiveness of tin are mixed, when correctly used, tin can effectively reduce the harmful effects of vanadium by 20-30 percent [ll.

The mechanism of tin passivation was not well understood, except for the general assumption that inert compounds were formed on the FCC surface [55]. Recently the effects of Sn on V-contaminated model catalysts were studied using Mossbauer spectroscopy [56] and electron paramagnetic resonance (EPR) measurements [57]. Tin-119 Mossbauer spectroscopy indicated that Sn-V interactions take place only during steam-aging. Mossbauer results indicated that tin was present as a Sn+4 species. Occelli proposed that Sn+4 formed ligands t o vanadium through oxygen bridges. The S n N complex formed in (V+5-0-Sn+4) units [58]. V-Sn alloys were not observed. In laboratory studies the order of deposition of tin and vanadium had little effect on the nature of the resulting species.

Since tin forms oxides which are stable up to 900°C, it was reasonable to assume that tin oxides would form at FCC regenerator conditions. Based upon this assumption, recent studies have applied Lewis acid-base oxide reaction concepts to explain Sn passivation of vanadium [59]. However, molten salt tests showed SnO2, presumably because of its acidic nature, essentially nonreactive with V2O5 or Na2O-V205. Since no chemical reaction between bulk Sn02 and V2O5 was observed a t FCC temperatures, the author concluded any tendency to form inert compounds between only Sn, V and 0 would be unlikely. A "three- way" complex between SnOg, V2O5 and the zeolite surface was proposed. Although an interaction of oxides did not occur in the molten salts experiments, the effect of steam atmosphere was not investigated. Therefore, the study did not discount the possibility that interactions would occur in the presence of steam.

In another study, XPS experiments were conducted to scan only the surface layers of the catalyst [60]. A strong tin signal was detected indicating tin remained predominantly on the catalyst surface. Measurements of surface vanadium concentrations were also made on samples containing vanadium only and vanadium plus tin. Results indicated about 2.5 times as much vanadium on the surface of the vanadium-only samples than with the vanadium-tin sample. This

366

was despite the fact that equal amounts of vanadium were added to both samples. These data indicated tin formed a thin layer on the catalyst, coating a portion of the vanadium.

5.1.1 Tin Commercial Experience Laboratory and pilot plant tests demonstrated the benefits of tin passiva-

tion of vanadium. In addition to passivating vanadium, laboratory MAT tests showed tin also passivated moderate levels of sodium [60]. This was consistent with the hypothesis that both vanadium and sodium catalyze the steam deactivation reaction [18]. At 1.0 wt.% sodium, tin increased conversion 4.0 vol.% and increased gasoline by 0.4 vol.%. Hydrogen increased slightly and coke decreased by 0.2 wt.%FF.

Since 1982, tin has been tested and used commercially [60]. In some early tests, mixed results were obtained partially since tin passivation was not analogous to nickel passivation. Tin prevented zeolite damage from vanadium whereas antimony diminished nickel's catalytic dehydrogenation activity. Antimony effectiveness commenced with addition to the unit and the effects were measurable in a short time. Tin interacted with vanadium and resulted in the preservation of the zeolite activity. It was ineffective on the pre-existing vanadium on the catalyst. The benefits of tin were only observed when the damaged catalyst was replaced which took several weeks and was dependent on the make-up rate. Some early tests were not run long enough to experience the benefits of tin passivation.

The Valero Refinery tested tin in early 1985 as a vanadium passivator in their HOC with hydrotreated resid feedstock [60]. The catalyst was a metals tolerant, very hydrothermally stable grade. Antimony passivation was also used during the test to passivate nickel. A previous baseline of the MAT conversion of the equilibrium catalyst (without tin) showed a loss of three MAT numbers with an increase in vanadium levels from 2500 ppm to 3000 ppm. During the tin test base case in October, 1984, the vanadium on the catalyst was increased from 3000 ppm to 3755 ppm causing the MAT activity to decrease from 73 to 68 vol.% converted. A 4.0% drop in unit conversion to 76.9 % was observed at the plant. Tin passivation was initiated again in February, 1985, and the vanadium levels were again allowed t o rise from 3000 ppm.

By early May, the vanadium level reached 3600 ppm. The MAT activity remained stable a t 73% whereas without tin passivation, the MAT declined to 68% (Figure 14 [60]) and the plant conversion had only declined by 0.5-1.0% to 81.2%. The surface area of the tin passivated catalyst was maintained. Of the five number improvement in MAT conversion, about 2.4 numbers was attributed to a higher catalyst makeup rate (from 0.65 to 0.91 lb/B [SO]) used to hold the vanadium on the catalyst to the 3500 ppm target. A net difference of about 2.6 points was thus attributed to tin passivation. This occurred even though the

367

2001

74 4 0

0

Figure 14. Tin maintains catalyst surface area and MAT activity [601. Reprinted by permission.

catalyst sodium level had increased substantially. The total light gases decreased by 32 SCF/BFF t o 241 SCF/BFF. The hydrogen yield decrease accounted for 10 SCF/BFF while methane increased by 9 SCF/BFF which resulted in a decrease in the hydrogen-to-methane ratio from 1.09 to 0.77. Further conditions and product data were reported, but the results were considered to be less useful than the MAT data due t o the feedstock and operating con-

368

Table 3 Tin Passivation f601. Unit Yields Without Tin With Tin

Gasoline, vol. % 57.2 58.3 Coke, wt. % 6.1 7.0

H2/C1, mole ratio 0.68 0.61

Conversion, vol. % 71.7 75.9

H2, wt. % 0.09 0.09

. . Qperatinz Conditio n s Regen. Bed Temperature, OF 1243 1261 Riser Outlet Temperature, O F 950 952

Recycle Rate, %FF 6.5 6.1 Feed Preheat, OF 517 393

Feedstock API Gravity 22.3 22.6 Sulfur, wt. %FF 1.13 1.16 Nitrogen, wt.%FF 0.16 0.14 Carbon Residue, wt.%FF 0.42 0.47 Aniline Pt., O F 176.7 175.6

Catalvst Ni, ppmw v, PPmw Sb, PPmw MAT Avg. Makeup, lbhbl.

1200 1200 1600 1600

170 160 67.8 69.2

0.09 0.09

Source: Chevron Research Company Reprinted by permission.

ditions changes. Due to the substantial benefits observed, the use of tin was continued for at least a year.

A relatively low level vanadium (1600 ppm) t e s t was conducted by Chevron Research Company. The equilibrium catalyst also contained 160-170 ppm antimony to passivate 1200 ppm nickel. An economically significant gain in MAT activity of 1.4 points (Table 3 [601) resulted which was in excellent agreement with predicted results based on pilot plant studies. The commercial unit conversion increased from 71.7 vol.% to 75.9 vol.%. The yield of coke increased slightly due to a higher activity catalyst and a decline in feed quality.

369

70-

65 - B E B B

B B S B B P P

E B B R, P P

P

*, *c P

*c B P

B PI.. **

*c B B B B B BB

Figure 15. Tin maintains activity at higher vanadium levels [60]. Reprinted by permission.

A tin passivation test was performed at high catalyst vanadium levels in a modern resid cracker feeding substantial quantities of atmospheric bottoms and vacuum gas oil. At a constant MAT activity of 57-67 vol.% conversion, the catalyst vanadium level was increased from 2500-5500 ppm without tin to 5000-7500 ppm with tin (Figure 15 [60]). Antimony passivation was also utilized during the test to mitigate the effects of the high nickel loadings.

5.1.2 Antimony And Tin Combinations of passivation agents in the FCC unit have also been used

successfully. In a test program [l], Phillips Petroleum Company evaluated an antimony-tin system. The raw results from the plant data indicated a 2.9% conversion increase, 2.2 percent increase in the yield of gasoline with only a 0.2% increase in the yield of coke. The plant data was further evaluated using a model to account for variations in fresh feed properties, slurry recycle and other process conditions. Adjusted for process conditions, the results improved to a 4.4% increase in conversion, a 2.6% higher gasoline yield, and only a 0.3% increase in the yield of coke due to higher conversion of lower quality oil charged. The unit fresh feed rate did not decrease.

By 1989, some FCC units having 1000-3000 ppm or higher concentrations

370

of vanadium on the catalyst routinely used antimony-tin passivation. Tin concentrations ranged from 110 ppm to 460 ppm [l]. Antimony:tin atomic ratios ranged from 3-11. Regenerator temperatures were 1200-1330°F. Incremental improvements over antimony alone were: conversion up to +3.0%, gasoline yield up to +2.4%, and yield of coke up to a 0.5% decrease.

The Department of Chemistry of Oil and Organic Catalysis of Moscow University, the All-Union Scientific Research Institute for Oil Refining and the Institute of Organic Chemistry, Academy of Sciences of the former USSR developed a one-step synthesis of nickel passivators from readily available reactants on the semi-industrial scale [61]. Results were comparable to the best foreign additive [61,62]. The water soluble passivator contained antimony and other elements. The agent was claimed as unique in that the ratio of ingredients could be varied to achieve optimum benefits for specific conditions [62]. For example Formulation A maximized the yield of gasoline while Formulation B, with the same catalyst and vacuum gas oil, maximized the reduction in coke and hydrogen yields. Oil-soluble formulations based on antimony, bismuth and tin were also developed [63] with the goal of low-toxcity and high efficiency.

5.2 Titanium And Zirconium Metal additives such as titania and zirconia are reported to tie up vanadi-

um through the formation of high melting point binary oxides with V2O5 [64]. With titanium and vanadium, no true compound could be identified. It is postulated that substitution of Ti+4 into the crystalline structure of V+4 occurs leading to the disappearance of the titania and the vanadium pentoxide X-ray patterns.

Limited commercial experience using TYZOR@ titanium additive from DuPont showed an increase in equilibrium catalyst MAT activity with the titanium [65]. The organo-metallic compound was introduced via the cracker feedstock. Studies conducted in circulating pilot units with TYZORB feed addition did not show an improvement in catalyst activity [66]. The riser time appears to be fast compared to the time required for titanium to interact and tie up vanadium on the catalyst.

5.3 Rare Earths A number of studies [67-701 have shown the benefits of ra re earths in

metal passivation. Feron [71] reports that hydrocarbon soluble rare-earth compounds, dysprosium and samarium napthenate and lanthanum octoate, are efficient vanadium passivating agents. Treatment a t a lanthanium: vanadium ratio of one preserved greater than 90% of the zeolite structure after high temperature-hydrothermal treatment (7OO0C, 20% steam). Using this practice, the catalyst composition would not require modification since the passivating agent is added in proportion to the concentration of vanadium in the feed. This permits a more flexible operation since cracker feedstocks of different vanadium content

371

could be processed without a catalyst changeout. Since rare earths react readily with vanadium to form chemically stable compounds, lanthanum or other rare- earths are effective passivators. Rare earth vanadates are stable even at the high temperatures reached in the FCC regenerator. The rare-earth vanadium compounds are thought to be distributed homogeneously on the different components of the catalyst, avoiding clogged pore mouths on the external surface.

6. ADDITIONAL AGENTS

Numerous passivsting agents have been claimed in various patents to passivate nickel, vanadium or sodium. The more well-known commercially tested passivation agents were discussed above. A list of other elements includes germanium [72,73], gallium [74], tellurium 1751, indium [75,761, aluminum [73,771, barium “781, zinc [75], boron [79,80], phosphorous “791, tungsten [81], tantalum [64], lithium [82] and cadmium 1751.

A ranking of various passivation elements relative to antimony for hydrogen reduction is shown in Table 4 [2]. The agents were primarily organo-metal compounds o r oxides of the elements. Arsenic did not effectively passivate the metals [l].

Table 4 Ranking of Elements for Reduction of Hydrogen

Element Hydrogen Reduction Relative t o Antimony Passivation

Sb T1 Bi P Sn In Ca Te Ba Ge A1 Si

1.0 0.8 0.7 0.6 0.5 0.4 0.4 0.3 0.3 0.2 0.2 0.2

~~~

Reprinted with permission from [2]. Copyright 1991 American Chemical Society.

372

7. PROCESS VARIABLES

Studies have shown that metals of lower oxidation state have less activity for dehydrogenation reactions to produce undesirable products during cracking [5]. Therefore, the activity of contaminant metals on cracking catalysts can be moderated during processing. An example is passing the regenerated catalyst through a reducing zone prior to returning the catalyst to the cracking zone [83,84]. Reducing atmospheres of hydrogen, hydrocarbons, or carbon monoxide in a broad range of concentrations are suggested as appropriate.

The use of sulfur-containing compounds capable of associating with the metal has been proposed [85]. These compounds can be used by themselves or in conjunction with the traditional antimony passivation. It is suggested that the metal may be converted to a sulfide or oxysulfide [85]. Hydrogen sulfide is typically used. However, o ther sulfur compounds such a s lower alkyl thiols, thioethers, and disulfides may also be used. A contact time of a t least three seconds is required, according to pilot plant studies. A separate vessel prevents hydrogen sulfide (or other sulfur compounds) from entering the riser.

Another method utilized to reduce the deleterious effects of nickel, vanadium and iron is to contact the catalyst with a lift gas containing not more than 10% C3 or heavier hydrocarbons [86]. These hydrocarbons selectively carbonize active contaminating metal sites on the catalyst to reduce hydrogen and coke production. The process proves particularly useful for heavy residual feedstocks. This method was in commercial use in over six units in 1991 [21]. When the vanadium is maintained in the +4 valence state, improved vanadium tolerance for the equilibrium catalyst is observed. V+5 oxidation state can also be prevented by operating the regenerator in a partial burn mode such that the CO concentration is maintained a t 1 mole%.

Unit design and process conditions can make a significant commercial impact on the activity of contaminant metals. Two stage regeneration reduces the rate of hydrothermal deactivation promoted by vanadium. The process conditions of the Reduced Crude Conversion (RCC) unit also significantly deactivate equilibrium nickel [27] so that only a 0.2 antimony to nickel ratio on the equilibrium catalyst is required. However, process conditions, such as steam in the riser [87,88] alone, are insufficient for economically processing of any but low metals feedstocks. Metals passivation agents or metals tolerant catalysts o r both can further reduce the effects of nickel and vanadium.

8. METAL TRAPS

For a number of years manufacturers have been working to increase the metals resistance of FCC catalysts and have been quite successful in developing a number of improved catalysts. One way these advances have been accom-

373

plished is the addition of diluents capable of selectively sorbing metal contaminants. These metals traps may be added to the FCCU catalyst as separate particles or incorporated directly into the catalyst particle during preparation. The physical blends are typically prepared at the plant prior to shipping. A variety of materials are identified in the literature as metals traps.

These trap components, which are typically inorganic oxides, are not so environmentally objectionable as are certain other compounds. A few of the inorganic oxides exist as naturally occurring minerals. The addition of a metals scavenger to the FCC catalyst generates a dual function cracking catalyst (DFCC) that initially (because of dilution effects) is less active. The vanadium scavenger selectively sorbs and immobilizes the vanadium contaminant. Nickel traps cause nickel to agglomerate o r become incorporated in the trap itself as shown by microscopic analysis of contaminated catalyst. For cracking catalysts which contain metals traps (DFCC) deactivation occurs a t a much slower rate than that of the typical FCC catalyst, as the metal concentration in the feed increases. At some point, depending upon the effectiveness of the metal trap and the feed metal concentration, the DFCC becomes more active as shown in Figure 16 [89].

Improvement in the metal tolerance of cracking catalysts can be achieved by changes in catalyst formulation as well as by the addition of separate particle metal traps previously discussed. Catalyst formulation changes include the incorporation of active elements or compounds into the catalyst matrix. These elements react with the incoming metal atoms to produce inert compounds on

METALS (Ni and V)

Figure 16. DFCC is more active at high metals than ordinary catalyst [89]. Reprinted with permission from ACS Div. Pet. Chem. Prepr. 32(3-4), 661 (1987). Copyright 1987 American Chemical Society.

374

the matrix surface, preventing the metal from being catalytically active for hydrogen and coke production. Vanadium which becomes tied up on the matrix outside the zeolite cage structure is effectively immobilized.

For incorporation catalysts, these additives are generally added to the catalyst by slurrying with the other catalyst components, faujasite, clay, binder, active matrix, and the like, before spray drying into particles. For cracking catalysts formulated in-situ with pre-formed particles, the passivation compounds may be impregnated or precipitated on the catalyst particle. In particular, the additive tends to concentrate on the outer surface of the particle.

8.1 Nickel Alumina is often used to improve the tolerance of commercial cracking

catalysts. Increased nickel tolerance is observed when large crystal, low surface area alumina is incorporated into the catalyst matrix which allows nickelto agglomerate into low surface area nickel crystals. The large nickel crystals leave substantially fewer active surface sites t o cause dehydrogenation reactions. The low surface area alumina can be directly prepared or can be produced by back-filling with silica causing pore blockage of the alumina and effectly reducing its surface area. Recently reported by Lam et al., nickel may also be tied up by encapsulation into a non-active Ni-alumina tetrahedral spinel structure [go].

Katalistiks International reported the results of a commercial trial of their nickel trap-containing catalyst in a high resid, moderate Conradson carbon unit [91]. During the period with the new catalyst, the feed nickel roughly doubled from 6 to 10 ppm which doubled the equilibrium catalyst nickel content from about 3000 to 6000 ppm. Antimony continued to be charged during the first part of &he trial as the inventory of trap-containing catalyst built-up. The antimony Concentration on the catalyst was allowed to decrease from a high of 2000 ppm shortly after the introduction of the new catalyst to a low of 700 ppm before the antimony addition was discontinued altogether for three months.

With the trap-contzining catalyst, conversion remained within the base range and approximately constant on a time averaged basis. Hydrogen yield (Figures 17 and 18 [91]) indicated the catalyst was more resistant to nickel poisoning. Decreasing the antimony level increased the hydrogen yield by only 0.03 wt.%. Hydrogen production was independent of nickel concentration with the new catalyst (Figure 18). The relative yield of gasoline, considering the feedstock variations, was insensitive t o antimony. Akzo's "Advanced Catalyst System'' traps and encapsulates the nickel [131 in low activity large pores reducing its dehydrogenation activity. It is reportedly most effective at higher levels. It is commerically used (up to 7,000 ppm) extensively in Europe [27].

375

.25

.20

Ep .15

8 .10

*05 J A S 0 N D J F M A M J J Month

Figure 17. Nickel trap catalyst (D) only marginally benefited from antimony; nickel increased with time [91]. Reprinted by permission from the June 1991 issue of HYDROCARBON PROCESSING, page 91, copyright Gulf Publishing Co., 1991, all rights reserved.

I I L~Avidusc'atallystl I I o Catalvst D base vanadium

- 1 1 1 1 1 1 1 1 1 1 -

1,500 2,500 3,500 4,500 5,500 6,500 Ni, PPm

Figure 18. H2 yield with Ni t r a p catalyst D was mainly due to V [91]. Reprinted by permission from the July 1991 issue of HYDROCARBON PRO- CESSING, page 91, copyright Gulf Publishing Co., 1991, all rights reserved.

8.2 Vanadium A naturally occurring material which has been reported as an effective

vanadium trap is sepiolite (Mg8Si1203o(OH)4). Sepiolite is a fibrous clay mineral composed of tetrhedral chains of silica lying in planes, joined together through shared oxygens by magnesium and aluminum atoms in octahedral coordination [92]. Raman spectra of 5% Vlsepiolite after exposure to steam aging conditions show no V2O5 or tetrahedral V04 species. Spectral bands are consistent with V2O7-. The compound that vanadium forms with sepiolite is dictated by the purity of the sepiolite. The V-compound on pure sepiolite is identified as MgV206, whereas in an impure sepiolite (containing CaC03 and MgCO3) the

376

vanadium also forms Mg3V208. Sepiolite is also an effective trap for basic nitrogen compounds [93]. Other selected magnesium compounds such as attapulgite, hectorite and chyrsolite are also effective vanadium scavengers. However, in the presence of sulfur dioxide, magnesium traps form very stable sulfates. For this reason, their use in units which process high sulfur feeds may not be feasible.

The effectiveness of pure magnesium oxide as a vanadium trap was investigated [94]. Test samples were prepared by mixing steamed cracking catalyst with MgO, impregnating the mixture with vanadium and nickel napthenates followed by drying, calcining and re-steaming. MAT tests measured higher activity on the preparations with MgO. Atomic force microscopy (AFM) analysis of the catalyst with added MgO revealed only the catalyst and a forsterite phase. Apparently a reaction between silica and MgO occurred during the steam treatment to form Mg2Si04. Scanning transmission electron microscropy energy dispersive X-ray (STEM-EDX) analysis revealed that all detectable vanadium was associated with the forsterite indicating its effectiveness as a vanadium scavenger.

The benefits of using calcium and magnesium dual components as vanadium scavengers in cracking catalysts were reported [95]. The preferred calcium/magnesium component material was dolomite while the preferred magnesium-containing material was sepiolite. Separate particles containing dehydrat- ed magnesium-aluminum hydrotalcite [ 961 or an alkaline earth metal oxide in combination with alkaline earth metal spinel, preferrably magnesium aluminate spinel, were effective as vanadium traps as well as SOX transfer catalysts.

Based upon the hypothesis that vanadium pentoxide reacted with steam to produce a volatile acidic species, alkali and alkaline earth metals were examined as traps for vanadium. Davison's "DVT vanadium tolerance" additive was based upon this chemistry [971. The additive was typically added to the cracking catalyst as a physical mixture in 80-95% catalyst to 20-5% DVT. Studies showed the DVT additive improved the vanadium tolerance of Super D catalyst by a factor of three, based on the vanadium level required to give the same activity as the catalyst without the additive. Limited commercial testing demonstrated the ability of DVT to scavenge vanadium in a regenerator. Microprobe analysis of typical catalyst and DVT particles from a commercial trial showed that DVT particles had higher concentrations of vanadium than the catalyst particles [97].

The interaction of rare earth oxides with vanadium has been well established [14]. Soluble rare earth compounds have been used as passivation agents [71]. Physically discrete particles of lanthanum oxide or other rare earth oxides have been identified as effective vanadium scavengers when mixed with cracking catalysts [98]. The oxide particles can be physically blended with the catalyst or added to the catalyst matrix or the feed stock [98].

377

The ability of vanadium to attack and destroy the zeolite in cracking catalysts is reduced significantly by nonionic deposition, such as precipitation, of rare earths rich in lanthanum on the catalyst prior to or during contact with vanadium containing feed 1703. This study finds that ion exchanging lanthanum or other rare earths into the zeolite does not provide as effective preservation of zeolite (as determined by surface area or X-ray intensity measurements), or provide the reduction in hydrogen production that precipitation gives at the same vanadium level.

In contrast, another study [69] found in laboratory tests that ionic deposition of the rare earths cerium, praseodymium, neodymium, or gadolinium and their compounds in an amount on a mass basis greater than the lanthanum present in the catalyst was effective. Cerium remained more effective a t higher metals levels (10,000 ppm), while praseodymium and gadolinium declined in effectiveness. The rare earth metals were added to the matrix of the catalyst by impregnating a fresh catalyst or by cracking the agent along with the cracker feedstock 1681.

Shot or sponge coke (preferrably) have been shown to adsorb metals when physically mixed with cracking catalyst [38], being more efficient for vanadium scavenging. Upon regeneration, the outer surface of the metallated coke particle is removed along with the metals. The oxidized metals are entrained as small particles which exit the regenerator as a fine dust in the flue gas stream.

Akzo reported in the mid 1980's [99] improved conversion over a base catalyst using a proprietary incorporated vanadium trap, 62.1 wt% compared t o 45.2 wt%. Relative yields of hydrogen and coke were better with the trap: 11.8 versus 14.9 and 1.8 versus 2.0, respectively. When a nickel trap was also incorporated, conversion further improved, 64.1 wt%, while the relative yields of hydrogen and coke dropped to 4.7 and 1.2, respectively. Since these results were reported, Akzo has developed a third generation proprietary trap, the first generation being MgO- or CaO-based with second generation BaTiO2.

Intercat currently markets a vanadium trap, V-Trap@, which was developed by Chevron [loo]. This new generation material is reported to be more effective to selectively demobilize vanadium than earlier versions. The manufacturers of V-Trap@ claim it is effective against vanadium even in units which process high sulfur feeds.

8.3 Scavenging Traps Scavenging traps containing proven active passivating elements such as

antimony and tin have been proposed. Separate particle traps were prepared with both antimony and tin compounds on a diluent material such as alumina, magnesium compounds, or titanium compounds [ lol l . The trap should exhibit a MAT activity of about 1.0. The antimony and/or tin compound was physically

378

mixed with or dissolved in a solvent an impregnated onto the carrier.

Laboratory studies using thermogravimetric analysis (TGA) show that at the temperatures and conditions of the FCC unit, antimony and vanadium are volatile while tin and nickel are not [58]. Nickel and vanadium deposit on both the catalyst particles and the trap particles during cracking. The nickel deposited on the separate particle will react with the antimony and form inert compounds. Unfortunately the nickel deposited on the catalyst particle will tend to remain in place. However, antimony on the diluent which has not reacted with nickel will vaporize.and react with nickel on the catalyst o r in the oil. Vanadium which has been deposited on the catalyst but has not reacted with the zeolite will become volatile [16] in the regenerator. The volatile vanadium species can interact with the tin on the trap particle and become immobilized.

8.4 Layered Catalysts The most recently devised class of metals tolerant catalysts is layered or

encapsulated catalyst particles comprising a conventional matrix-zeolite inner core covered with a porous outer shell or layer. The layer may be of various materials that act as metals traps [68]. A low activity, attrition resistant shell is composed of clay with a silicate binder [102]. A benefit of layered catalysts over catalysts with the metal trap embedded in the matrix is the removal of metal from the FCCU through attrition of the shell. Fresh surfaces for trapping metal are formed, extending the life of the core catalyst.

Several other layered-type catalysts have been patented. One uses a porous, rare earth oxide/aluminum oxide/aluminum phosphate as a suitable coating [67,103]. Another, a more complex shell catalyst, has up to three core components: a dealuminated Y-type zeolite, a shape selective zeolite like HZSM-5, and a shape selective aliphatic aromatization material such as gallium ZSM-5 within an alumina-rich matrix shell [102]. The alumina shell adsorbs metals. The shell may also contain vanadium traps such as magnesium oxide or lanthanium oxide. The shell coating gradually attrites off, thereby removing the metals from the unit with the fines.

9. SUMMARY

Passivation of nickel and vanadium on cracking catalysts is an established industry practice. Implementation of a metals passivation program requires little capital and only a small increase in operating expenses. The antimony passivation of nickel has been demonstrated in a variety of unit designs cracking a range of feedstocks containing from low to high metals. When properly conducted, the process is considered safe and environmentally acceptable. Newer passivation agents offering less toxicity and less environmental risk than antimony are based on bismuth and cerium. Commercial tests have shown their efficiency to be equivalent t o antimony in some units while other units respond better to

379

antimony and others to one of the other additives.

The passivation of vanadium with certain tin compounds complements the passivation of nickel. Since tin reduces the rate of destruction of zeolite, the time necessary for measuring a significant benefit is dependent on the replacement rate of the aged, poisoned catalyst with fresh, treated catalyst. By maintaining the zeolite's integrity, the conversion and selectivity are maintained a t higher levels than without passivation.

The mechanism of nickel passivation by antimony is not well understood but has been proposed to involve antimony-nickel alloys with surface enrichment of antimony. Geometric blocking of the nickel sites and accompanying electronic effects reduce nickel's chemisorption properties and catalytic activity. The mechanism for vanadium passivation by tin is also not well understood. Inert compounds of vanadium and tin are generally assumed to forrnm the catalyst surface. XPS measurements suggest tin forms a thin coating over a portion of the vanadium.

An alternative to passivation agents is the more recently developed metals traps which immobilize nickel or vanadium on specific sites on the carrier or on the catalyst particle. Layered or shell coated cracking catalysts are a newer route to metals control. Several passivation techniques may be combined to obtain the optimum product distribution for a particular situation. For example, nickel passivation agents injected into the oil feed or separate trap particles can be added quickly in response t o an increase in feedstock metals, whereas additives incorporated into the catalyst would be expected to have a longer response time due to the need to build unit inventory.

Passivation will change a unit's heat and material balances resulting in new steady states and possibly new operating restraints. New steady states may appear, for example, as an increased catalyst circulation rate is required to maintain the regenerator bed temperature due to the reduction in the yield of coke. New unit restraints may be an air blower limiting capacity instead of the process gas compressor. When passivating catalyst, especially at moderate to high metals levels, the refiner should plan for the expected operating changes.

The refiner may utilize the benefits of metals passivation directly through increased selectivity at constant conditions: the reduction in hydrogen and coke are accompanied by an increase in yield of gasoline. Alternatively, conditions may be changed to maximize the economic benefits resulting from the improved selectivity. Where the catalyst make-up rate has been dictated by the metals loading, it may be possible to reduce the catalyst make-up rate. Unit capacity for the same feedstock is increased. Lower quality feedstock may be processed. The operator may use higher reactor severity to increase conversion in units previously below their optimum.

380

10. FUTURE DIRECTIONS

Researchers may find opportunities for fundamental and applied studies in metals passivation. As is the case in refinery process history, the commercial application is b e t t e r understood than the basic chemical mechanisms. Knowledge will increase as newer, specialized analytical tools become available which allow surface analysis at the metals levels found on equilibrium catalysts. A mechanism that accounts for all observations has yet to be published. A better understanding of the differences in laydown efficiencies between units and agents may suggest ways to improve overall efficiency.

In a more applied research area, improved process mathematical models would be helpful in predicting the passivation benefits and how the protocol needs to change in response to variations in feedstock or the desired product slate. The challenges of reformulated gasoline requirements include producing less olefins in the gasoline pool with more butenes and isobutane for alkylation. As a result, riser temperatures are likely to increase, increasing diolefin production. For example, a pilot plant experiment a t 1050°F reactor temperature revealed that metals passivation reduced the butadiene yield by 90%, increased the propylene yield by 20% and the total butenes by 10% when a catalyst with 10,000 ppm of metals was passivated [2].

The behavior of metals and passivators in ultra-short contact reactors which are under development will require study. Metals laydown and passivation efficiencies would be expected to differ with shorter reaction time.

Future work will inevitably evolve due t o increasing environmental demands. For example, waste miminization programs may require processing of used lubricating oils in cracking units. The development of new passivating agents for metals such as cadmium, zinc, copper and lead will be needed. In view of forseeable regulations on TCLP for the disposal of spent cracking catalysts, leaching of metals from the catalyst surface is of concern. More effective ways of tying up metals on the catalyst surface area may be required, especially for high metals resid operations.

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1990 NPRA Question & Answer Session on Refining and Petrochemical Technology, Ed. M.L. Batchelder, Gerald L. Farrar & Associates, Inc. (Tulsa, OK), 59-61 (1990). A. Corma, M.S. Grande, M. Iglesias, C. del Pino and R.M. Rojas, Appl. Catal., a, 61-71 (1992). L. Zaiting, Katalistiks' 8th Ann. Fluid Cat Cracking Symp. (Budapest), (1987). Jiujiang Refinery, Research Institute of Petroleum Processing, Shiyou Lianzhi (Petroleum Processing), &,28-32. Translated from Chinese by

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R.H. Nielsen, D.L. McKay and G.H. Dale, U.S. Patent 4,148,712 (1979). R.E. Wrench, J.W. Wilson and C.F. LeRoy, NPRA Ann. Meet., Paper AM-

J.B. Rush, NPRA Ann. Meet., Paper AM-81-43 (1981). 1988 NPRA Question & Answer Session on Refining and Petrochemical Technology, Ed. M.L. Batchelder, Gerald L. Farrar & Associates, Inc. (Tulsa, OK), 46-48,51-3, 68-70 (1988). 1991 NPRA Question & Answer Session on Refining and Petrochemical Technology, Ed. G.L. Farrar, Gerald L. Farrar & Associates, Inc. (Tulsa,

1987 NPRA Question & Answer Session on Refining and Petrochemical Technology, Ed. M.L. Batchelder, Gerald L. Farrar & Associates, Inc. (Tulsa, OK), 61-2 (1987). L.M. Il'ina, M.N. Pervushina and V.N. Erkin, Chem. Technol. Fuels Oils,

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1983 NPRA Question & Answer Session on Refining and Petrochemical Technology, Ed. M.L. Batchelder, Gerald L. Farrar & Associates, Inc. (Tulsa, OK), 65-6 (1983). W.C. McCarthy, T. Hutson, Jr. and J.W. Mann, How t o estimate the benefits from the Phillips passivation process, Phillips Petroleum Co., R&D (Bartlesville, OK), (undated). W.C. McCarthy, T. Hutson, Jr. and J.W. Mann, Katalistiks 3rd Ann. Fluid Cat Cracking Symp. (Amsterdam), (1982). C.K. Teran, NPRA Ann. Meet., Paper AM-88-70 (1988). T.C. Readal, R.A. McKinney and R.A. Titmus, U.S. Patent 4,036,740 (1977). A.R. English and F.A. Pettersen, Petrol. Ref. and Petrochem. Processing Intl. Conf. (Bejing), Technical Reprint (1991). Unpublished information supplied by Intercat, (1992). R.S. Heite, A.R. English and G.A. Smith, Oil Gas J., 88(231, 81-3,86-7 (1990).

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65 66

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68 69 70 71 72 73 74 75 76 77 78 79

80 81 82

R.C. Barlow and JJ. Lipinski, NPRA Ann. Meet., Paper AM- 89-17. (1989). D.R. Forester, U.S. Patent 4,913,801 (1990). A.R. English and D.C. Kowalczyk, Oil Gas J., 8 2 0 , 1 2 7 - 8 (1984). J.M. Stencel, Raman spectroscopy for catalysis, Van Nostrand Reinhold (NY, NY), p. 123 (~1990). M.W. Anderson, M.L. Occelli and S.L. Suib, J. Mol. Catal., a, 295 (1990). M.W. Anderson, M.L. Occelli and S.L. Suib, J. Catal., 1120 ,375 (1990). M.L. Occelli, S.M. Naraghi, V. Krishnan and S.L. Suib, J. Catal., I%, 325- 31 (1992). R.L. Jones, J. Catal., 129,269-74 (1991). F.W. Denison 111, J.F. Hohnholt, A.R. English and A.S. Krishna, NPRA Ann. Meet., Paper AM-86-51 (1986). E.A. Karakhanov, S.V. Lysenko and K.M. Minachev, Petrochem. Symp. of Socialist Countries, Materials of [sic] Symp. (Kozubink, Poland), I, 8-14 (1988). E.A. Karakhanov, A.A. Bratkov, S.V. Lysenko, L.M. Il'ina, A.E. moiref and T.K. Mellik-Akhnazarov, VI Petrochem. Symp. of Socialist Countries, Materials of [sic] Symp. ,I, 523-6 (1988). E.A. Karakhanov and S.V. Lysenko, ZhVKhO, fQ, 622-25, Translated from Russian by Ralph McElroy Co., Custom Division, P.O. Box 4828, Austin, TX 78765 (1989). H.W. Beck, J.D. Carruthers, E.B. Cornelius, W.P. Hettinger Jr., S.M. Kovach, J.L. Palmer and O.J. Zandona, U.S. Patent 4,432,890 (1984). W.P. Hettinger, S.M. Kovach and H.W. Beck, U.S. Patent 4,496,665 (1985). E. Chao, Unpublished memorandum, PC 87-11, Ashland Petroleum Company, R&D (Ashland, KY), (1987). P. Chu, A. HUSS, Jr., H. Owen, J.A. Herbst, G.W. Kirker and P.H. Schipper, U.S. Patent 5,077,253 (1991). P. Chu, A. HUSS, Jr. and G.W. Kirker, US. Patent 5,001,096 (1991). Z.C. Mester, U.S. Patent 4,900,428 (1990). W.H. Beck, C.F. Lochow and C.W. Nibert, U.S. Patent 4,515,683 (1985). B. Feron, P. Gallezot and M. Bourgogne, J. Catal., U, 469-78 (1992). B.J. Bertus and D.L. McKay, US . Patent 4,490,299 (1984). D.R. Forester, U.S. Patent 5,019,241 (1991). J.S. Roberts, D.L. McKay and B J . Bertus, US . Patent 4,377,504 (1983). C.F. Bertack, U.S. Patent 4,522,704 (1985). J.S. Roberts, B J . Bertus and D.L. McKay, U.S. Patent 4,256,654 (1981). J.S. Yoo, U.S. Patent 4,337,144 (1982). B.J. Bertus and D.L. McKay, U.S. Patent 4,377,494 (1983). V.A. Durante, D.J. Olazanski, W.J. Reagan and S.M. Brewa, U.S. Patent 4,430,199 (1984). H. !h, U.S. Patent 4,295,955 (1981). D.L. McKay, B J . Bertus and H.W. Mark, U.S. Patent 4,290,919 (1981). H. !h, U.S. Patent 4,364,847 (1982).

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88

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95 96

97

98 99

100 101 102 103

D.F. Tatterson and W.D. Ford, U.S. Patent 4,298,459 (1981). G.D. Myers, W.P. Hettinger, Jr., S.M. Kovach and O J . Zandona, US. Patent 4,432,863 (1984). A.A. Avidan and A.A. Chin, U S . Patent 4,986,896 (1991). H.U. Hammershaimb and D.A. Lomas, U S . Patent 4,479,870 (1984). 1981 NPRA Question & Answer Session on Refining and Petrochemical Technology, Ed. M.L. Batchelder, Gerald L. Farrar & Associates, Inc. ('hlsa, OK), 56,58-60,77 (1981). 1982 NPRA Question & Answer Session on Refining and Petrochemical Technology, Ed. M.L. Batchelder, Gerald L. Farrar & Associates, Inc. ('hlsa, OK), 44,50-1,58-9,70,75-6 (1982). M.L. Occelli, ACS Div. Pet. Chem. Prepr., 32(3-4), 658-62 (1987). C. Lam, P. O'Conner and C.P. Smit, in Akzo Chemical Symposium '88, May 29 - June 1,1988, Ed. H.J. Louvink, (The Netherlands), (1988). W.S. Letzsch, L.L. Upson and A.G. Ashton, Hydrocarbon Process., W,

R.E. Grim, Clay Mineralogy, McGraw-Hill (N. Y., NY), (~1968). A. Corma and F.A. Mocholi, Appl. Catal., a, 31-46 (1992). J.B. Hall, and E.H. Hirshberg, in Proceed. of the 45th Ann. Meet. of the Electron Microscopy Soc. of America, Ed. G.W. Barley, San Franscisco Press (San Franscisco), 200-1 (~1987). J.V. Kennedy and L.W. Jossens, US. Patent 5,002,653 (1991). A.A. Chin, I.D. Johnson, C.T. Kresge and M.S. Sarli, US. Patent 4,889,615 (1989). R.E. Ritter, L. Rheaume, R.F. Wormsbecher, A.W. Peters, D.N. Wallace, E.T. Habib and P.G. Thiel, NPRA Ann. Meet., Paper AM-85-47 (1985). A.A. Chin and M.S. Sarli, U S . Patent 4,921,824 (1990). 1985 NPRA Question & Answer Session on Refining and Petrochemical Technology, Ed. M.L. Batchelder, Gerald L. Farrar & Associates, Inc. (lblsa, OK), 40-2 (1985). A.K. Rhodes, Oil Gas J., 90 (41),41-48 (1992). M.L. Occelli and H.E. Swift, U S . Patent 4,466,884 (1984). D.M. Stockwell and G.S. Koermer, U S . Patent 5,082,814 (1992). A.W. Chester, P. Chu, A. HUSS, Jr. and G.W. Kirker, U S . Patent 4,919,787 (1990).

89-92 (1991).


CHAPTER 11

UNIT DESIGN AND OPERATIONAL CONTROL: IMPACT ON PRODUCT YIELDS AND PRODUCT QUALITY

LAWRENCE L. UPSON, CHARLES L. HEMLER, DAVID A. LOMAS

Process and Systems Development Department

UOP 25 E. Algonquin Road, Des Plaines, Illinois, 60017

I. INTRODUCTION

In 1992, the fluid catalytic cracking process celebrates its half-century anniversary. This colorful process has a history filled with contributions from a vast array of individuals and corporations. The early pioneers created a unit design that has since evolved into a configuration that incorporates a blend of science, art, engineering skill, equipment improvement, and operating know-how. Many of the early characteristics and features of the unit design are still recognizable, but others have undergone some dramatic changes.

Throughout its history, one basic design concept has remained as the dominant feature of the fluid catalytic cracking unit (FCCU): the reaction zone and the catalyst regeneration zone have remained intimately connected to each other. Catalyst is first contacted with hydrocarbon feed in the reaction zone. Coke, which is one of the reaction products, quickly deposits on the catalyst surface, producing a large drop in catalyst activity. The catalyst is then separated from the volatile reaction products and stripped with steam. Next the catalyst is transferred to the regeneration zone, where the coke is burned off, and the activity is restored. The catalyst is then transferred back to the reaction zone, and the cycle is repeated.

Several important features of this reactor-regenerator cycle make the FCCU a unique processing tool within a refinery complex. First, the time of a complete cycle is short, compared to other cyclic refinery processes. In modern FCCU design the average catalyst contact time in the reaction zone can be as short as two to three seconds, and the entire cycle can last less than 10 minutes (1). To achieve this rapid cycling between reactor and regenerator, the catalyst circulates at high rates, which can be as high as 1 ton/second in large units.

Second, the reactor and regenerator operations are linked by the heat balance. The exothermic heat released from burning the coke from the catalyst in the regeneration zone provides the heat needed in the reaction zone. This heat is

385

386

carried into the reaction zone by the circulating catalyst. Thus, the reactor and regenerator performance are interdependent on each other, with the relationship between the two being complex. Changes in the design of one part of the system have significant effects on the performance of the whole system.

The purpose of this chapter is to describe what changes have occurred in FCCU design over the years. The features of newly designed units are emphasized along with a description of how these design changes and the resulting changes in process variables effect the results that are achieved in an FCCU.

11. REACTOR DESIGN

A. Historical DeveloDment

On May 24, 1942, the first commercial FCCU was put on-stream at the Baton Rouge refinery of Standard Oil of New Jersey. The purpose was to produce aviation gasoline to fuel Allied planes during World War 11. The FCCU concept was an immediate success. By the end of the war, 33 other FCCU’s of this first design and of a modified design (Model 11) had been put into operation).

1. Early Design

The heart of any catalytic process is the reaction section where, feedstock and catalyst are brought into contact. Although the early designers envisioned a system with a few-seconds contact time, the activity levels of the early catalysts did not allow such a choice. Thus the early operations required a large amount of unconverted product recycle to compensate for the low catalyst activity and the low conversion per pass that resulted.

Even with these process limitations, a number of units of the Model I1 design were built during the 1940’s through the combined effort of various organizations (the Recommendation 41 Group). The widespread commercial use of the FCCU process began with this Model I1 design. This design (Figure 1) featured a reaction section with a dense, fluidized bed of catalyst. Following World War 11, different organizations began to develop designs that made use of a dilute transport zone, where feed and catalyst were contacted before entering the dense-bed reactor. This transport zone, where the hydrocarbon and catalyst mixture was lifted in dilute flow through the pipe leading to the dense, fluidized bed of catalyst, came to be known as the riser reactor section.

2. Importance of Riser Cracking

The stacked unit design introduced by UOP in 1947 (Figure 2) was a successful early design in which a relatively long riser served as this transport zone. A dense- phase reactor of catalyst still followed the riser. Measurements taken along the riser (2) showed that a substantial portion of the conversion was taking place and the selectivity to desired products was more favorable within the dilute-phase section (Figure 3). This design configuration became popular, and elements of this

387

Figure 1

Downffow Model 11 Catalytic Cracking Unit

Air

I ' Column Bottoms Rcprlnled fmm Stmlkr, el al., Ref. 2. UOP zcc+i

configuration have remained until the present time, although the contribution of the dense-phase reactor section has been minimized over the years.

The Shell organization took a big step toward more complete cracking within the riser when they introduced a new unit design in 1955 (3). In this new design, the riser discharged directly into cyclones attached to the end of the riser. Unconverted heavier liquid fractions continued to be recycled, although now this recycle was directed into a catalyst bed within the reactor rather than to the riser as in past operations. However, again the catalysts available in the middle 1950's did not allow the full potential of such a design to be achieved.

3. Changes in Catalyst and Operating Conditions

The introduction of higher activity zeolitic catalysts during the 1960's has had a profound effect on FCCU design and performance over the years. The higher activity and much improved coke and gas selectivity, which resulted from the switch from amorphous silica-alumina catalysts to zeolite-based catalysts, allowed refiners and designers to consider much higher, once-through conversion levels and higher reactor temperatures than had previously been used.

Attention began to shift back to a system with a shorter reaction time. Recycle rates were lowered, and reactor temperatures increased. Once again, design changes began to emphasize the contributions of the riser and special attention was given to riser configuration as a way to reduce undesirable backmixing. As the 1970's began, risers were extended well into the reactor, and the dense bed of catalyst was

388

60

4 40

30

0

Figure 2

UOP Stacked FCC Unit

,-onversion--, - ---mmmm --I--- A Ovnall ~ 57.3 1111111111111111-

Gasoline 46.0 LV-W

Overall

Riser Time Appror Dense Bed Time w

4*

I I I I I I I l l

50:/;

Stripper

389

5.0 18.7 45.4 21.5 9.4

eliminated from the reaction path. Risers were equipped with cyclones and other mechanical devices to more quickly separate the reaction products from the catalyst.

Operating conditions were also undergoing changes. Starting in the early 1970’s, complete combustion conditions within the regenerator led to regenerated catalyst with a low carbon content. At that same time, unleaded gasolines and the emphasis on higher octanes pushed reactor operating temperatures even higher. More active catalysts with improved selectivity led to the elimination of cycle oil recycle and even higher once-through feed rates.

3.8 17.3 49.8 21.7 7.4

This combination of design changes, catalyst changes, and different operating conditions produced a significant shift in the product distribution for the typical gas oil operation (Table 1) (4). Higher conversion levels, more selective operations, and significantly higher feed rates were all manifest in the changing operations.

30.0 54.4 20.0

104.4

Table 1 Changes in FCCU Yields as

Designs, Catalysts, and Operating Conditions Change

27.8 59.3 20.0

107.1

Time Period

Design Features Cracking Mode

Combustion Mode Feed-Cat Mixing

Catalyst Type

Yields, wt-% c2-

C,s and C,s Gasoline Cycle Oils Coke

Yields, Liq. vol-% C,s and C,s Gasoline Cycle Oils

Total C, t

Gasoline RON

1960 I 1970

partial partial

92.0 I 91.0

1980 I 1990

riser

complete poor

USY Zeolite

riser with rapid disengaging

complete good

USY Zeolite

with Active Matrix

4.0 17.9 50.9 21.8 5.4

3.3 17.9 52.5 21.5 4.8

28.8 60.4 20.0

109.2

28.8 62.6 20.0

111.4

92.5 I 93.0

390

For many refiners the quality of the FCCU feedstock also changed during the turbulent 1970’s, as soaring crude prices forced many refiners to process heavier and more contaminated fractions (5). This trend continued into the 1980’s when refinery economics demanded even more selective operations. The 1990’s brought a need for gasoline reformulation to satisfy increasing environmental constraints on automotive fuel. This need has led to a corresponding interest in even more severe FCCU processing conditions.

B. Modem Reactor Desim

In modern reactor design, the intent is to maximize primary reactions, which produce high yields of the desired liquid products: gasoline, light cycle oil (LCO), and light olefins. At the same time, these designs attempt to minimize secondary hydrogen-transfer (H-transfer) reactions, which result in the saturation of olefins and the overcracking of gasoline and LCO and hence a reduction in the yield of these desirable products.

1. Maximizing Plug-Flow Riser Cracking

Riser cracking, which occurs in an essentially plug-flow, short-residence time regime, maximizes the desirable reactions. Post riser cracking on the other hand, occurring in the reactor vessel, takes place in a backmixed regime, which provides a much longer exposure of catalyst and hydrocarbons and consequently facilitates the occurrence of these undesirable secondary reactions. Thus, modern reactor design has evolved toward systems that provide rapid, postriser separation of hydrocarbon from catalyst to minimize these nonselective, postriser, backmixed reactions.

A variety of different mechanical arrangements have been developed to more closely approach all-riser cracking. Exxon uses the terminology transfer-line cracking to describe this situation (6) Mobil has developed a mechanical system they call closed cyclones (1). Shell showed a related configuration (Figure 4) in a recent publication that describes various Shell reactor designs over the years (7), Ashland’s patented vented riser (8) provides a rapid but simple catalyst-vapor separation using an open-ended riser. Other commercial applications use descriptive nomenclature such as direct-connected cyclones (9) for this service.

With cracking now taking place primarily in the riser, the contact time between catalyst and hydrocarbons is quite short, on the order of just a few seconds (1). Achieving high conversion in these systems would thus be expected to be more difficult than in the older designs, which had substantial postriser contact between catalyst and hydrocarbons.

2. Yield Effects

The yield advantages for these separational systems depend on the extent to which the different arrangements provide rapid separation of hydrocarbons from catalyst and then provide subsequent containment and rapid removal of the hydrocarbons from the reactor. Invariably a side-by-side comparison of these

391

Figure 4

Shell FCCU Designs

1970's 1980's 1990's 1940 - 1960's

Reprinted from Naper, eta!., Ref. 7. W W +

UOP 2038-4

392

Yields

HZS c2-

c3s c4s Gasoline Cycle Oils Coke

Total

Conversion

RON MON

systems is difficult, but recent literature describing units that have been revamped to the short contact time systems, have indicated the magnitude of the yield shifts when compared to some of the older conventional reactor configurations (1,9). Table 2 shows the results of an FCCU reactor revamp at Cenex Oil, where a riser, T-disengager system was replaced by a system in which the riser discharged directly into catalyst separating cyclones (9). As a result of the change, conversion decreased, but both gasoline yield and research octane increased. The more selective cracking, the higher product olefinicity, and the higher octane are due to minimizing postriser cracking and suppressing secondary reactions.

Wt-% Liq. Vol-% Wt-% Liq. Vol-%

0.5 - 0.5 - 3.8 - 3.0 - 5.0 9.0 4.7 8.3 9.1 14.3 10.0 15.6

47.9 57.8 48.7 58.8 26.5 24.6 28.4 26.0

100.0 '105.7 100.0 108.7

73.5 75.4 71.6 74.0

- - - - 4.7 7.2 - -

92.9 94.1 80.6 80.5

Table 2 Results From FCCU Revamp at Cenex Oil

I Before I After

C. Catalvst Strimer Design

Although much of the attention for these new separational devices has been focused on the reactor, many of these approaches recognize the need for an initial displacement of hydrocarbons from the catalyst and then a well-designed stripping section to remove whatever hydrocarbons can be separated from the catalyst before passing to the regenerator. Thus, displacement and prestripping have become significant, and renewed attention has been directed toward an effective multistage stripping section.

393

Actually the catalyst stripper has been an often overlooked section of the unit design. The effective and simple movement of the particulate solids requires the presence of a fluidizing medium. Hydrocarbon vapors leaving the reactor provide this fluidization. One of the functions of the catalyst stripper is to displace these hydrocarbons and replace them with steam. The stripper also provides a zone in which adsorbed hydrocarbons can desorb from the catalyst surface, thereby helping to lower the amount of hydrocarbons actually carried to the regenerator. Thus, over the years, efforts in the stripping section have been aimed at improving the contacting and displacement efficiency and in providing an adequate desorption residence time. For example, in UOP designed units, stripper performance has been improved by using a series of alternating, horizontal baffles with elongated skirts (63). This design eliminates channeling and improves stage efficiency.

D. Feed Distribution

Poor mixing of feed and catalyst causes some loss of cracking selectivity in the lower section of the riser. However, in the long-residence-time reactors of earlier design, good mixing of the feed and catalyst was not a critical item. The loss of selectivity as a result of poor mixing in the riser was overshadowed by the more extensive loss of selectivity created by backmixed cracking in the bed or in the postriser vapor space.

With all-riser cracking, intimate and uniform contact between feedstock and catalyst throughout the riser and particularly at the point of initial contact is essential to achieve the maximum selective benefits that this design offers. Thus, the modern feed distributor makes use of a multinozzle injection with adequate steam and pressure drop to atomize and disperse the feed. Early dispersion systems were not much more than appropriately sued pipes through which feed was added. Good feed-catalyst mixing was not achieved. Multiple nozzle systems were shown to offer advantages in cross sectional area coverage particularly for large risers (10). Multinozzle distributors, combined with appropriate distribution steam and pressure drop also offered yield advantages because the feedstock was better dispersed (11). One approach to feed distributor design today uses very large quantities of dispersion steam and very high distributor pressure drop in an effort to produce small oil droplets (12). Most other commercial feed distributors did not go to this extreme to provide feed dispersion but made use of appropriate combinations of steam and pressure drop to achieve good feed dispersion and contact with the catalyst.

A feed dispersion system is more than just a feed distributor, particularly when heavy, high-boiling feed components are involved. For example, the Ashland RCC unit (Figure 5) , uses a light hydrocarbon gas to first lift and condition the catalyst before the feedstock is brought into contact with the regenerated catalyst (13). This lift gas both assists with feed dispersion and helps to passivate the highly metal contaminated catalysts that are common in a residue operation. Specific details of the effect of lift gas are described in Section IV.B.5 of this chapter. An alternate system uses selected recycle streams partway up the riser to provide a higher

394

Figure 5

RCC Process

Reaction - Products

Reactor 4 Riser

- Residueand Diluents

c-- LiftCas

W P 1385A.4 UOP2WB-5

reaction mix temperature at the base of the riser (14). Institut Frangais du PBtrole (IFP) claims advantages for this system as an essential portion of the EUR process.

Interest in alternative feed systems and variations in feed-catalyst contacting is not just a recent phenomenon. An early design by Texaco used a dual riser system, which featured a selective, low-conversion first riser and then subsequent cracking of the remaining unconverted product in the second riser (15). Gulf advocated cracking individual feed components at distinct positions along the riser by segregating the feedstocks (16). UOP patented another alternative (17), which

395

involves the cracking of hydrotreated LCO recycle (the J-Cracking process) as a way to achieve selective gasoline yields and enhanced octanes.

111. REGENERATOR DESIGN

A. Evolution of Desim

The main purpose of the regeneration zone is to remove the coke from the spent catalyst coming from the reactor. The catalyst activity of the fluidized particles is thereby reestablished prior to being returned to the reactor riser.

Over the years, the regenerator section has undergone considerable fundamental design and mechanical improvements. Early designs were of the bubbling-bed type with superficial gas velocities in the regenerator bed of less than 1 ft/sec. The temperature in the regenerator bed was typically in the 1200-1250°F range. Commercial regenerations are now typically run at higher superficial gas velocities in the turbulent-bed or fast-fluidized-bed regimes along with operating temperatures up to 1400°F. Figure 6 contrasts a modern high-efficiency combustion design with a typical less desirable configuration of the past. The primary vessel elevations are now set to achieve maximum regenerator-reactor pressure differentials, because the process favors high regenerator pressure (to enhance combustion kinetics and power recovery) and low reactor pressure (to provide the best product yields and selectivities). The main regenerator improvements are aimed toward:

Enhanced and controlled coke burning kinetics in the regenerator Reduced catalyst inventory in the regenerator Ease of start-up and routine operability Uniform radial carbon and air distribution Limited afterburning and temperature differentials between cyclones Narrow catalyst residence-time distributions Additional heat balance flexibility - Two-stage regeneration - Dense-phase catalyst cooling Particulate, power, and waste-heat recovery

These benefits are derived from the combination of a more homogeneous gas and solids fluidization regime, reduced yet controlled particle residence-time distributions, and improved handling of the regenerator heat balance.

B. Basic Concepts in Regenerator Desim

1. Heat Balance Considerations

One of the key factors in regenerator design is a thorough understanding of the regenerator heat balance. Since the FCCU regenerator is intimately coupled with

396

Figure 6

FCC Unit Configurations

Past Present

Regenerator Reactor Regenerator Reactor

5 psig

\ 20 psig 1

J Air U

Steam & Diluents

UOP 19ffiE-2 UOP 20084

the FCCU reactor, the regenerator heat balance is strongly influenced by the reactor heat balance. Thus, the overall FCCU heat balance must be considered in defining the regenerator heat balance (Figure 7).

The heat released from the combustion of coke is the largest single component in the regenerator heat balance and in the overall FCCU heat balance as well. Thus, the factors that influence the amount of combustion heat that is generated must be well understood. These are the factors affecting the kinetics of burning coke to CO, CO,, and H,O.

397

Figure 7

FCCU Heat Balance

UOP 2008-7 Steam

a. Coke-Burning Kinetics

The coke deposited on the catalyst consists primarily of carbon and hydrogen along with relatively small amounts of sulfur, nitrogen, and metals (17) This discussion focuses only on the carbon and hydrogen species in the coke. With the use of high conversion, zeolite catalyst and modern stripper designs, the amount of hydrogen in coke has typically decreased over the years to levels of 6-7 wt-%.

Catalyst regeneration, as a carbon removal process, is widely accepted as being described by a first order kinetics expression with respect to carbon concentration and oxygen partial pressure as given in Equation 1 (18):

398

- - dc = KCP,,, dt

Where: K = Rate constant, hr-l atm" C = Carbon on catalyst, wt-% Pm = Oxygen partial pressure, atm

The reaction rate constant will depend on temperature and can be expressed in an Arrhenious-type equation:

A 6 - _ - dC = Koe RT c P , , ~

dt

Where: AE = Activation ene ra R = Gas constant T = Temperature, OR

This relationship for the rate of carbon burning from an FCCU catalyst was found experimentally to be valid over a wide range of temperatures. Diffusion to and from the catalyst surface and within the catalyst pores did not limit the carbon- burning rate.

A similar expression (Equation 3) can be used to describe the reduction in hydrogen content of the coke during combustion. The oxidation of hydrogen, however, proceeds more rapidly than that of carbon.

A 6 - _ - = KLe RT H poz

dt (3)

Where: K', > K,

As described in Equation 2, the rate of carbon burning depends on the carbon level on the catalyst and the temperature of the catalyst. The catalyst temperature, in turn, depends on the amount of carbon and hydrogen burned, because the heat of combustion of these components raises the catalyst temperature. Thus, the simultaneous solution of differential Equations 2 and 3, coupled with a regenerator heat balance, can provide insight into the impact and optimization of process variables, such as temperature, carbon and hydrogen content of the coke, and oxygen particle pressure, on the overall coke combustion rates. These equations can be solved with the aid of a few simplifying assumptions, such as a plug-flow regime, when a combustor-style regenerator is used (see Section C.l) or backmixed flow, when a bubbling-bed regenerator is used.

399

When discussing coke combustion rates and kinetics it is important to distinguish between the coke yield (as a fraction of the feed) and the coke concentration (as a fraction of the catalyst weight). The relationship between net coke yield and coke on catalyst is somewhat analogous to the relationship between heat transfer rate and temperature. The net coke yield is set by the unit enthalpy balance with little direct impact on the individual particle’s rate of coke combustion. The net coke yield as a fraction of feed can be mathematically represented by an overall FCCU heat balance.

b. Modeling the Heat Balance

The various heat balance regimes associated with the FCCU are depicted in Figure 7. As can be seen, individual heat balances can be taken around the reactor (heat balance no. l), around the regenerator (heat balance no. 2), and an overall heat balance can be taken around the entire system (heat balance no. 3).

Taking the heat balance around the entire FCCU and solving for the coke yield produces Equation 4:

Where: AH Rx = FCCd

- AHRx -

Enthalpy change of feed at reactor inlet conditions to reaction products at reactor exit conditions, Btu/lb of feed

Heat of cracking, Btu/lb of feed

Enthalpy change of any diluents (i.e. steam or lift gas) from reactor inlet to reactor exit, Btu/lb of feed

Enthalpy change of any recycle stream from reactor inlet to exit, Btu/lb of feed

Heat of combustion for burning coke to CO, CO, and H,O, Btu/lb of coke

Enthall ly change for air between inlet and outlet of the regencrator, Btu/lb of coke

Heat losses from both reactor and regenerator, Btu/lb of coke

Equation 4 shows that the coke yield is essentially independent of both feedstock and catalyst quality since neither of these items enter into the equation. This equation does show that changes in the coke heat of combustion (AH play

400

a major role in the coke yield. The heat of combustion depends on the amount of CO burned to CO,, frequently expressed as the flue gas CO,/CO ratio. The effect of the CO,/CO ratio on AH improved stripping efficiency, hence lower H, content in the coke, actually increases coke yield, because decreasing the H, content decreases the heat of combustion.

is shown in Figure 8. Equation 4 also shows that

Figure 8

Heat of Combustion us. C02 I CO Ratio in Flue Gas

- - - .. - - - - - - - - - Total CO Combustion - - - - 17,000

16,000 - r"

f! (16,500 Btu / Ib Coke) .

11,000 I I I I I 1.0 2.0 3.0 4.0 5.0 6.0 i

C% / CO Mol Ratio D

UOP lS25H-33 UOP 20086

The coke yield from Equation 4 can also be related to the change in coke content on the catalyst during regeneration. This change, known as delta coke (A C) is described by Equation 5:

Where: C,,, = Coke content of catalyst entering the regenerator - wt-% of the catalyst

catalyst C,, = Coke content of catalyst leaving the regenerator - wt-% of the

Delta coke can be related to the coke yield (as % of feed) if the ratio of catalyst circulation rate to hydrocarbon feed rate is known (Equation 6). This ratio is commonly known at the cat/oil ratio.

40 1

Coke Y i e l d = c a t / O i l

Delta coke is a dependent variable, which is difficult to accurately predict because of its dependency on feedstock quality, processing conditions, and catalyst formulation.

c. Effect of Process Variables on Regenerator Heat Balance

From a heat balance around the regenerator, Equation 7 can be developed. This expression relates the final regenerator temperature to the catalyst delta coke value.

Where: TRcgcn = Regenerator dense phase temperature TRX = Reactor outlet temperature A H’- = Heat loss from regenerator/lb of coke AC = = Csv, for many modern day operations

i. Regenerator Temperature and Coke Concentration

The impact of temperature alone on the rate of carbon combustion is exponential and quite dramatic, as shown in Table 3.

Table 3 Effect of Temperature on Carbon Burning Rate

Temperature, O F I Relative Combustion Rate

1100 I 1 .o ””” 1400

The burning rate decreases substantially as the carbon content on the catalyst decreases (Table 4) at constant temperature, oxygen partial pressure, and delta coke reduction.

402

Table 4 Effect of Carbon Concentration on Burning Rate

Burning

II II

11 0.15 I 0.05 I 0.10 I 0.1 11

ii. Oxygen Partial Pressure

This parameter is set by the combination of regenerator operating pressure, excess air, and oxygen concentration in the supply air. The relationship between the rate of combustion and oxygen partial pressure is linear. Figures 9 and 10 show the relative effect of total pressure and excess oxygen on the regeneration time at

Figure 9

Effect of Oxygen Concentration on Regeneration Time

100 90 80 70 60 50 40

30

20

t

10 I I I I I I I

1100 1125 1150 1175 1200 1225 1250 1275

Temperature, "F

0 Total Pressure Assumed Constant at 30 psig 0 Plug Flow Regeneration System 0 Carbon Reduced from 1.10 to 0.3 Wt-% on Catalyst 0 Excess 02 Change Made by Air Rate Adjustments

UOP 1W6E-11 UOP 2008-9

403

130

Figure 10

Effect of Pressure on Regeneration Time Requirements

Plug Flow ksumed Carbon Level Reduced from 1.10 to 0.3 wt-W -

110

T 90 8

70 s 5 5 0 - 3

30 p?

- - -

1150°F

1200°F 1250°F

- I I I I I I I I I I

___. .

10 ’ 0 5 10 15 20 25 30 35 40 45 50

Total PTessure, psia UOP 1906E-12 uoP2008-10

various temperatures. Figure 11 shows the effect of excess oxygen at various temperatures on the coke content of regenerated catalyst from a bubbling bed regenerator at constant residence time (19).

iii. Total Combustion and CO Promoter

In the absence of a CO combustion promoter, large variations in CO,/CO ratios in the regenerator flue gas are observed (20). At the catalyst surface the ratio of CO,/CO is believed to be an intrinsic function of the temperature at the burning site (Arthur’s ratio) (21). However, the CO exiting the burning site may be further oxidized to CO, at a rate that is dependent on temperature; CO, 0, and H,O partial pressures; active metals on the catalyst; carbon-oxygen distributions within the fluidized bed; and even effected by the presence of FCC catalyst (Figure 12).

This burning of CO to CO, in the dilute phase, known as afrer burning, can produce large increases in the flue gas temperature above those in the dense phase. This increase in temperature is due to the much larger heat of combustion from burning carbon to CO,, compared to burning to CO, (Table 5) , and to the absence of a large catalyst heat sink, such as is present in the dense phase.

404

0.5

3 0.4 Y-

2 a 5 0.3 F e 5 0.2 2

5 5

F

- Y

a

er

3 0.1

0.1% Excess 02 -

0.5%

1.0%

-2.0%

-

Conditions:

- 3secRxTime

1150 1200 1250 1300 1350 1400 1450 1500

Temperature, O F

UOP tWgE-14 UOP2MY)-l2

405

Table 5 Carbon Heats of Combustion

Heat of Combustion, Reaction Btu/lb of carbon

C+ 'hO, - .CO

c + 0 , d co, 14,150

In a bubbling-bed regenerator, the AT between dilute and dense phase is frequently 60" to 100°F when a CO combustion promoter is not used. When a CO combustion promoter is not used, partial CO combustion in the bed is not possible. However, complete CO combustion is possible without a promoter if a substantial excess of oxygen is used (3 to 4% 0, in the flue gas) and if regenerator dense-phase temperatures are kept in excess of 1,300"F (19).

In the early 1970's, researchers at Mobil patented the use of platinum at low levels (1 to 3 ppm) in an FCC catalyst inventory to catalyze the combustion of CO to CO, in the dense phase (22,23). With a CO combustion promoter present, CO combustion occurs readily at regenerator dense-phase temperatures well below 1,300"F. Guegin has reported dense-phase CO combustion occurring at 1,200"F in a commercial FCCU (24).

In a bubbling-bed regenerator, the use of a CO combustion promoter greatly reduces the degree of afterburning by transferring the burning of CO to CO, from the dilute phase to the dense phase. In one commercial operation, a reduction of 100°F in afterburn AT has been reported as a result of using a combustion promoter (25).

2. Regenerator Fluidization

a. Fluidization Regimes

In the various commercial FCC regenerator designs, several fluidization regimes exist (26). Figure 13 pictorially illustrates these regimes.

i. BubblingBed

The bubbling bed regime ranges from the minimum bubble velocity (Umb), which typically for an FCC type powder is from 0.02 to 0.1 ft/sec, up to about 1.0 ft/sec superficial velocity (27). Here three distinct yet interchanging gas phases exist: the bubble phase, the emulsion phase, and the gas phase within the catalyst pores. These three phases all flow at different relative velocities. Discrete bubbles of gas flowing through the bed produce abrupt pressure fluctuations at the surface of the

406

Figure 13

Fluidization Regimes

- Gas Velocity

UOP lgOBE-3 UOP W - 1 3

bed (28). The magnitude of these pressure fluctuations is determined by superficial gas velocity (U,) as shown in Figure 14. One characteristics of the bubbling bed is a relatively large pressure fluctuation, compared to higher velocity flow regimes.

Figure 15 represents a typical bubblingbed regenerator that has limited solids entrainment from the dense bed surface to the cyclone inlet. Most of the larger particles (>50 p) that are entrained return to the bed through the two-stage cyclone diplegs In such a bubbling bed, a distinct surface of bubbling catalyst separates the lower, dense phase from the dilute phase above the bed. In the dilute phase, the density of t h e entrained solids decreases as the distance above the bed surface increases. Eventually a height above the surface is reached where the amount of entrained catalyst particles reaches a minimum. This height is known as the transport disengaging height (TDH). To minimize the amount of catalyst fines carried into the cyclone separator, regenerators are designed to have the distance between the inlet to cyclone first stage and the surface of the bed to be greater than the TDH (29) (Figure 15).

ii. Turbulent Bed

With higher gas velocities, (1.0 to 3.5 ft/sec), the distinct bubble phase disappears and the bulk gas flows as described by Yerushalmi (1) "in voids which continually coalesce and split, tracing tortuous passages as they rise through the bed (26)." The upper bed surface is considerably more diffuse with reduced pressure fluctuations and substantially higher entrainment of solids into the region above the

407

4.0

3.0

2.0

1.0

0.0

-1.0

Figure 14

nansition from Bubbling to lhrbulent Bed

-

-

-

-

I I I I I

UOP 2oMI-14 Souce: Yemhalrni and Cankurt Ref. 28.

Figure 15

Bubbling-Bed Regenerator

Two-Stage . Cyclone

1 T.D.H.

Flue Gas

Catalyst

Z l z o r

Catalyst to

Reactor

Combuster Air

UOP 1D06E-5 w)P 2ooo.16

408

dense bed. Because of the requirement for higher coke-burning capacity and improved contacting efficiency, the vast majority of commercial regenerators are operating in the turbulent-bed regime. In this regime, the ultimate regeneration capacity is set by the increase in solids entrainment as gas velocity increases (Figure 16), and by the cyclone separation efficiency and dipleg hydraulics (30).

Figure 16

Maximum Dilute-Phase Entrainment in Vertical Gas-Solids Upffow

100.00

10.00 a" \

1 .oo d:

0.10

0.01 0.10 1 .oo 10.00 Ve2 / gDpPp2

eS = Maximum Entrainable Ibs. of solids/ft3 of Gas Ve = Effective Gas Velocity, ft/sec Dp = Particle Diameter, ft Pp = Particle Density, Ib/ft3 Pg = Gas Density, IbIft3 Ws = Ibs. of Soliddsec x ft2

B = Gravitational Constant, 32.2 Wsedsec

Reprinted from API Publication #931, May 1975.

iii. Fast-Fluidized Bed

UOP 1906E6 UOP2008-16

The fast-fluidized regime, with its superficial velocities in the region of 3 to 10 ft/sec, extends into a complex transport phase where the rate of solids carry over increases sharply as the transport velocity is approached. In the absence of any solid recycle, the bed would rapidly disappear. Beyond this velocity, catalyst fed to the base of the regenerator transverses it in fuUy entrained transport flow. The voidage

409

or density of the resulting suspension is dependent not only on velocity of the gas but also on the solids flow rate. This flow rate is frequently referred to as the solids flux rate, which is measured in units of mass per area per time. If the solids rate is low, dilute-phase flow will result. If on the other hand, solids are fed to the regenerator at a sufficiently high rate, for example by the recirculation of solids as in Figure 13, then the relatively large solids concentration referred to as the fast- fluidized bed may be maintained. Yerushalmi and Cankurt discuss these various regimes of fluidization in great detail (26).

The transport velocity may therefore be regarded as the boundary that divides vertical gas-solids flow regimes into two groups. Below the boundary, the bubbling and turbulent fluidized beds exist. Above the boundary lies the transport regime, which encompasses a wide range of states from dilute-phase flow to the fast fluidized bed depending on the solids flow. Because of the formation of catalyst particle agglomerates, the FCC catalyst transport velocity is approximately 20 times the terminal velocity of a single 50 p FCC particle.

b. Solids Mixing Characteristics in the FCC Regenerator

The mixing characteristics of the bubble, or gas-induced solids in the various fluidized-bed regimes have received considerable attention over the years (31). Although the bubble-induced vertical mixing rate of solids is extremely high, the radial mixing characteristics are relatively poor (because bubbles flow vertically).

This facet of the fluidized bed can cause problems because some commercial bubbling- and turbulent-bed regenerators exceed 50 ft in diameter and have relatively low length/diameter (L/D) bed ratios ( < 0.5). Severe gas-solid (carbon- oxygen) distribution problems that can be encountered result in:

0 Thermal gradients, both interparticle and within the bed 0 Variable residence-time distributions and short circuiting 0 Afterburn in the dilute phase and uneven cyclone temperatures

Using radioactive isotopes, Singer investigated catalyst mixing patterns in various commercial FCCU’s (reactor, regenerator, and stripper sections) (32). The measured distributions indicate that although regenerator dense beds approach good mixing, substantial deviations from perfect mixing are attributable to catalyst bypassing and regions of relatively immobile catalyst. Axial mixing problems in bubbling beds have also been discussed by Karl-Wirth and by Kunii (33,34).

Over the years, radial mixing characteristics and residence-time distributions (short circuiting) of the turbulent- and bubbling-bed systems, have been improved through the use of dual diameter regenerator vessels and higher L/D ratios in the regenerator bed section, by achieving more uniform air distribution through a better design of the air grid hole patterns, and by achieving a more uniform distribution of catalyst in the regenerator. The distribution of spent catalyst entering the regenerator has been improved by using deflector plates and baffles and by venting the spent catalyst inlet line to create tangential swirls. Catalyst bypassing has been

410

reduced by optimizing the location of the regenerated catalyst outlet port. Cyclone dipleg orientation has also been improved to provide more uniform distribution of the catalyst fines being discharged from the cyclones and returned to the bed.

C. Modem Reeenerator Desig!

1. High-Emciency Combustor Design

The UOP high-efficiency (combustor style) regenerator, a design introduced in the late 1970’s, has minimized most of the intrinsic mixing problems inherent in bubbling-bed regenerators. This design effectively uses the fast-fluidized regime mentioned earlier.

Figure 17 shows a typical high-efficiency regenerator. The lower combustor section can be operated either above or below the fluid cracking catalyst transport velocity. The resulting catalyst-air suspension in the combustor depends not only on

Figure 17

High-Efficiency Regenerator Flue Gas

Air

Catalyst to and from Reactor

W P WWE-7 WP2wB-17

41 1

the gas velocity but also on the solids flux rate. Slide valve control of the quantity of catalyst being externally recirculated to the inlet of the combustor can therefore be used to control the catalyst inventory and residence time in the combustor for various air rates.

a. Combustor Hydraulics

Figure 18 represents typical combustor hydraulics for various catalyst loadings and superficial gas velocities. The combustor density is measured across the entire combustor height and the catalyst loading is the summation of both the spent catalyst circulation and the combustor external recirculation. As seen in Figure 18, the gas superficial velocity (A) lies well below the catalyst transport velocity, but at a low solids flux (region x-y), dilute-phase flow exists. At condition (y), the solids flux is sufficient to choke the system. Any further increase in solids loading (region y-z) results in a substantial density increase.

At the higher gas velocities (B) and (C), in Figure 18, choking takes place at much higher solids flux. The result is a less abrupt change in combustor density with further increases in flux. This area is the fast fluidized region. Eventually, at%& higher gas velocities described by line (D), true transport flow is achieved. At this . point no reasonable solids flux can choke the system.

Figure 18

Combustor Operation

Catalyst Loading (W), lb/sedft2 UOP (-Ed UOP20oO-18

412

Figure 18 shows how flexible the system is in adjusting the combustor hydraulics (inventory and residence time) for various operating conditions, such as temperature, pressure, superficial gas velocity, catalyst circulation and carbon concentration. For example, if the superficial velocity is increased from (B) to (C) either by a change in pressure or combustor air flow rate, the density of the combustor inventory will decrease from (E) to condition (F) at constant solids loading. However, if necessary, the combustor catalyst density may be reestablished at condition (G) by opening the external recirculation slide valve and increasing solids flux. Any increase in combustor inventory results in a decrease in the upper regenerator level (surge inventory).

b. Combustor Solids Mixing Characteristics

Most of the intrinsic mixing problems mentioned previously are eliminated with the high-efficiency style regenerator. Figure 19 compares the theoretical particle residence time distributions of a modern fast-fluidized combustor with that of a conventional bubbling bed. The basic difference in response curves is due to the differences in the fluidization regime and to the differences in the solids entry-exit configuration. The combustor regenerator more closely approaches a plugflow system because of reduced backmixing and the impossibility of a spent catalyst particle exiting the regenerator without passing through the dilute-phase combustion zone.

Figure 19

Theoretical Residence Time Distributions

# 3 .-

Lr,

Approach to Plug Flow Y u c 3 Y

‘ic (100% Backmixed) W

Time, sec UOP 1-E-10 UOPZWU-19

413

c. Catalyst Recirculation to the Combustor and Precombustion Temperature

As discussed in a previous section, controlled catalyst recirculation to the combustor serves to:

Raise precombustion temperatures

Because the externally recycled solids are at final regeneration temperature,

Increase residence time via combustor hydraulics

they set the precombustion, or initiation, temperature of the spent catalyst. This precombustion temperature (TMX) is the mix temperature of the cold. spent catalyst, hot recycled catalyst, and the combustion air (Figure 20), as calculated by Equation 8.

Figure 20

Catalyst Precombustion Temperat ure

I I I I I 0.0 0.5 1.0 1.5 2.0 2.5

Catalyst Recycle Rate, (Ibflb Cat Circ.)

Basis: Reactor Temp. - 980°F BlowerTemp. = 320°F Ib AirIIb Coke = 14.2

Ha in Coke = 6.0% RegenTemp. = 1344°F ACoke = 0.8%

CP = 0.273

UOP 1WBE.S UOP po8-20

414

0.275F (CCR), Tkllm + 0.275 (CCR) TRT + 0.273 BT,

0.2738 + 0.275F (CCR) + 0.275 (CCR) rh =

Where: B = Airlb/hr CCR = Catalyst circulation, lb/hr F = Catalyst recirculation, wt-% CCR TB = Air blower outlet temperature, O F

0.275 = Specific heat of catalyst, Btu/lb/"F 0.273 = Specific heat of air, Btu/lb/"F

The mix temperature, which is strongly influenced by the amount of recirculated catalyst, in turn strongly affects the coke combustion rate. This interrelationship is shown in Table 6.

Table 6

Catalyst Regen Recycle Concentration, wt-% Temperature, O F Time, sec

1.5 CCR 0.8 to 0.05 1161

For identical conditions the overall combustion time is reduced by a factor of 3. In reality, the combustion time would be reduced further by additional internal recirculation (solids slip).

d. Combustor Coke-Burning Profiles

The interaction of the process variables can be mathematically modeled to calculate the performance of the combustor. Such a model combines a heat balance around differential longitudinal segments of the combustor with the kinetic equations 2 and 3 for carbon and hydrogen burning. As an example of such a calculation, Figure 21 shows the calculated time-dependent responses of the numerous products of combustion for a typical set of combustor operating conditions.

e. Commercial Experience with the High-Eflciency Combustor

The high-efficiency combustor design has been successful commercially: more than 35 units are operating worldwide. The advantages of this design over the earlier bubblingbed design have been described by Cabrera and Mott (35) as:

Lower catalyst inventory (approximately 50% of a bubbling-bed system) Quicker response to changes in catalyst inventory

Figure 21

Combustor Composition Pro files

q

& 0 a m * 1 Hydrogen, wt-%

Carbon,wt-% - Carbon, wt-% 0.8 -a= Oxygen, mol-%

= = Catalyst Temp.

- -

- 1350

0 10 20 30 40 50 60 Time secs -

415

L

Conditions: Regen Pressure = 30 psig Aidcoke, lbflb = 14.18 Ha in Coke = 6.Owt-% Coke on Cat = 0.8 wt-% RegenTemp. = 1344°F BIowerTemp. = 320°F Catalyst Recycle = 1.5 Reactor Temp.. - 980°F 02 in Air = 21.0mol-%

Assumes Total Combustion Plug Flow KIP 1W6E-13 K IP 2wB-21

0

0

Improved regeneration: lower carbon on regenerated catalyst Simpler operation and more stable unit control

2. Regenerator Design for High Coke-Making Feeds

Recently, considerable interest has been shown in increased feedstock flexibility, particularly to accommodate residue blends and other high-boiling components. Because residual feedstocks contain higher quantities of metals and coke precursors, they produce incrementally more coke per unit of conversion.

416

Therefore, to maintain optimum conversion and selectivities, additional flexibility to control the unit enthalpy balance is required. During the 1980’s two-stage regenerators and catalyst coolers were commercialized to provide this flexibility.

a. no-Stage Regenerators

In the Total reactor plus two-stage regenerator (R2R) design, the metallurgical temperature limits of 1400 “F employed in conventional FCC regenerator design can be exceeded with the use of refractory lined external cyclones attached to the second stage regenerator vessel. Final regeneration temperatures are unconstrained, with up to 1700°F being theoretically possible. Figure 22 from Mauleon, et al., illustrates this design (36). The first regenerator operates in a partial combustion bubbling bed regime (with coke burned partially to CO and partially to CO,). Operating in the partial combustion mode reduces the coke heat of combustion (Figure 8), compared

Figure 22

7luo-Stage Regenerator in lbtal R2R Process

2ndStage 7 FlueCas

2nd-Stage u Air 0

1st-Stage Flue Gas

417

to complete combustion of coke to C O , and keeps the maximum temperature in the first stage to approximately 1300°F. In the first zone, 60 to 80% of the coke and essentially all the hydrogen and sulfur are burned. The first-stage flue gas leaves the system through conventional two-stage cyclones to a CO boiler. Catalyst then flows via dilute-phase transport to the second-stage regenerator, where the remaining coke is burned at temperatures up to 1700°F. This elevated temperature is possible, from a metallurgical standpoint, because the second stage is designed without any temperature limiting internals.

The regenerator configuration (Figure 23), used in the Ashland-UOP RCC” reduced crude conversion design, also consists of two distinct turbulent-bed regeneration zones. In the RCC regenerator, both bed temperatures are maintained at 5 1350°F. This temperature control is achieved partly by limiting the amount of CO that is burned to CO, in the primary (upper) regenerator stage. In the RCC regenerator, the temperature is additionally controlled by using a catalyst cooler of

Figure 23

RCC 7tuo-Stage Regenerator

Flue Gas 9

Regenerated Catalyst

UOP 138SA-15 UOP 2DM-29

418

either a flow-through or backmix design to remove heat from the primary stage. The lower secondary zone operates in total CO combustion to achieve carbon levels on regenerated catalyst of less than 0.05 wt-%. The flue gas from the secondary zone flows into the upper zone, where the excess oxygen is consumed in the partial combustion, primary burning zone. This design, therefore, requires only one flue gas, two-stage, cyclone separation system.

b. Catalyst Coolers

Heat-removal systems have long been used in commercial FCCU’s (37). In the past, however, all these designs were plagued by either poor mechanical reliability or lack of range in heat removal (38). A relatively new, improved design using an external dense-phase flow-through or backmixed tube bundle has demonstrated excellent mechanical reliability (39). These modern coolers have achieved the same or better on-stream efficiency (3 years) as many of the other mechanical components of the FCC system and have demonstrated a wide range of easily controlled heat-removal duties. This unique all-vertical bundle design allows either a backmixed or flow-through configuration (Figure 24). The bayonet tube design uses forced-water circulation and allows for catalyst to flow on the shell side of the bundle. The hot catalyst in the catalyst cooler is kept fluidized by passing air up through the catalyst bed on the shell side.

Figure 24

Catalyst Cooler Configurations and Control

Flow- Through

Regenerator Regenerator

Backmu

/

Aeration Steam &Water

water

UOP 1979.19 UOP ZMYI-24

Water

419

Catalyst coolers of this design take advantage of two properties of fluidized dense-phase catalyst beds:

As a result, both the catalyst flux and the aeration velocity can be maintained

Fluidized bubbling beds mix well in the vertical direction Fluidized FCC catalyst acts as an excellent heat-transfer medium

at low levels, and the heat-transfer coefficient between catalyst on the shell side and water on the tube side can be closely controlled over the range of 10 to 100 Btu/(hr)(ftz)( OF). The independently controlled catalyst flow produces a countercurrent solid-gas (bubble phase) flow. mically a single cooler can be operated and controlled anywhere within a heat transfer duty of 20 to 150 x 106 Btu/hr. This precise control of the cooler duty is much more easily achieved than is control of the flue gas COJCO ratio (control of the heat of coke combustion) and, therefore, offers the most convenient method to control regenerator temperature.

Figure 25 shows a typical catalyst cooler and combustor configuration, where the cooler can be operated in either a flow-through or backmixed mode. In the flow- through mode, catalyst flows from the upper regenerator through the catalyst cooler and into the bottom section of the combustor zone. This recirculation of cooled catalyst provides direct control on the temperature of the regenerated catalyst leaving the combustor and entering the upper regenerator. In the backmixed mode, the cooled catalyst is returned to the upper regenerator at a point near the regenerated catalyst exit line. This form of catalyst recirculation strongly influences the temperature of the regenerated catalyst leaving the regenerator and going to the reactor but has little influence on the temperature of the catalyst recirculating back to the combustor and hence has little influence on combustor performance. The location of the cooler in relation to the riser disengager arms, recirculation standpipe, and regenerated standpipe is thus an extremely important factor in providing flexibility to operate in either of these two modes.

In a VGO-resid operation, the coke concentration on spent catalyst is high because of the large amount of coke precursors in the feedstock. Without the cooler, excessive catalyst temperatures would result in severe catalyst deactivation (particularly with metals present), possible damage to regenerator metallurgy, and low catalyst/oil ratios. The cooler is operated to recycle cooled catalyst back to the regenerator and thus removes heat during the combustion stage. Optimum kinetics, catalyst/oil ratios, and activity maintenance are achieved to maintain maximum product selectivities. This mode is illustrated in Figure 25 (Case B).

A catalyst cooler is not only useful for controlling regenerator performance for high coke-making resid type feeds but also has utility when processing more conventional, lower coke-making, vacuum gas oil (VGO) feeds. For a typical VGO operation, with the regenerator operating to achieve complete combustion of CO to CO,, the regenerator temperature would be expected to be in the range of 1300 to 1320°F. This temperature level provides good kinetics for rapid coke burning, has little potential for thermal deactivation for modern high-stability FCC catalysts, and is normally well below the design temperature limits for the regenerator metallurgy.

420

Figure 25

Catalyst Coo Ier Configuration and Operution

UOP 19061-158 UOP 2008-25

Thus there is no incentive to use the catalyst cooler to lower the regenerator temperature. However, there is a considerable incentive to lower the temperature of the regenerated catalyst before it contacts the hydrocarbon feed. As described later in Section IV.C.1 of this chapter, the Iower catalyst temperature would result in increased conversion (through higher catalyst/oil ratios), and improved product selectivities. The backmixed mode illustrated in Figure 25 (Case A) allows for this type of operation.

42 1

IV. FCCU Operational Control

A. General Process Control

The operator of an FCCU has control over several important process variables that strongly influence the product yields produced by the FCCU. Included in these variables are the temperature and pressure of the reaction zone, the feed-dispersion conditions at the riser inlet, and the carbon on the regenerated catalyst. In some units where catalyst coolers have recently been installed, the regenerator temperature can be independently controlled. In theory, fresh feed rate and feed quality are variables that can also be controlled, but in most refinery situations, these two factors are controlled by the overall refinery operating strategy.

To some extent, equilibrium catalyst quality can be controlled by the FCCU operator. The daily fresh catalyst addition rate and equilibrium catalyst withdrawal rate are under the operator’s direct control, and so the operator controls the equilibrium catalyst activity and the amount of contaminant metals on the equilibrium catalyst. By controlling the metals content, the operator can, to some degree, control catalyst selectivity. From a longer range point of view, the operator can change the type of fresh catalyst being used and thereby alter, even more significantly, the selectivity characteristics of the equilibrium catalyst. However, a change in fresh catalyst would normally require 3 to 6 months to fully implement. The influence of fresh catalyst quality on product yields from the FCCU is discussed in detail in Chapter 8.

The use of catalyst additives to control special functions, such as CO combustion control, octane enhancement, and flue gas SOX control, can also be controlled by the FCCU operator. Because these materials are normally added separately, by an addition system that is independent of the bulk catalyst addition system, their influence can be quickly altered by changes in addition policy. Details concerning the use of these specialized additives are covered in Chapters 10, 13, and 14.

The purpose of this section is to discuss the above-mentioned process variables, which are within the control of the FCCU operator, and to describe the impact of these variables on the product yields from the FCCU.

B. Reactor Process Variables

1. Reactor Temperature and Conversion Interaction

a. Temperature Control

Reactor temperature is the reactor parameter that is most readily controlled by the operator. In most units, reactor temperature is normally controlled by varying the catalyst circulation rate. By varying the amount of hot catalyst coming from the

422

regenerator that contacts the hydrocarbon feed entering the riser (and thereby varying the catalyst/oil ratio), the temperature of the catalyst and oil mixture at the inlet of the riser can be varied, and hence the temperature at the riser outlet is varied. Units that control catalyst flow rate by using slide valves on the catalyst lines leaving the reactor and regenerator normally control riser temperature in this manner.

Another method for controlling riser temperature is by controlling the temperature of the hydrocarbon feed to the riser. Units such as Exxon Model IV, which control catalyst flow rate through variations in the pressure balance between reactor and regenerator, normally control riser temperature through variations in hydrocarbon feed temperature rather than by variations in catalyst flow rate.

b. Temperature Effect on Conversion and Product Selectivity

The effect of increasing riser temperature is complex. The higher temperature alone will increase conversion; but if temperature is raised by increasing the circulation of hot catalyst, the increased catalyst presence will further increase conversion. Conversion increases of 1 to 2% absolute/lO"F increase in riser outlet have been reported for typical FCCU operations (40,41).

Further complexity arises because the reaction paths that produce the multitude of products derived from catalytic cracking have a wide range of activation energies, and thus the yield selectivities of these products varies as temperature is increased. For example, the reaction to produce isobutane has a positive activation energy, but the reactions to produce butylenes have a negative activation energy. Thus, at constant conversion, increasing riser temperature will decrease isobutane yields and increase butylene yields. This point is illustrated by riser pilot plant data in Table 7, adapted from Owen et al., for cracking a heavy VGO over a zeolitic catalyst (42).

Table I The Effect of Reactor Temperature on Product Yields

0.52

423

60% Conversion 70% Conversion

Gasoline 81.0 77.9

Total C,s 10.8 11.9

Total C,s 5.2 5.9

G- 1.2 1.7

Coke 1.8 2.7

c. Conversion Effect on Product Selectivity

Increasing riser temperature also increases the rate of conversion, which in itself shifts the product selectivities: gasoline selectivity decreases and light hydrocarbon selectivities increase. Selectivity data adapted from Upson and Sikkar illustrate this point, as shown in Table 8 for cracking an East Texas VGO over a zeolite catalyst in a laboratory Microactivity Test (MAT) at constant reactor temperature (43).

75% Conversion

75.5

12.8

6.3

2.3

3.5

Table 8 The Effect of Conversion on Product Selectivities

60 I 70 80

Product olefinicity also changes with conversion as data from reference 43 illustrate in Table 9.

Table 9 The Effect of Conversion on LPG Olefinicity

1 C,=/Total C, 1 0.86 1 0.78 1 &I: 11 C,=/Total C, 0.48 0.36

d. Temperature and Conversion Effects on Gasoline Octane

Riser temperature and conversion also strongly affect the quality of the gasoline produced in an FCCU. Reactor temperature is well known to have a strong

424

5 8 2 - - u 78

74

effect on gasoline octane, particularly research octane (RON). Values of 0.4 to 1.0 RON/lO°F reactor temperature increase have been reported for zeolitic catalysts (44,45). A commonly used rule of thumb is 1.0 RON/lO"C and 0.5 MON/lO"C (0.6 RON/lO"F and 0.3 MON/lO"F).

More recent data on FCC gasoline octane indicate that the gasoline octane response to temperature is strongly dependent on the cut point of the gasoline (46,47). Figure 26 shows the octane temperature response for a variety of gasoline fractions from several different refineries (47). In this figure, the research octane

C C5-250 0.4 D 175-355 0.5

mm----m-----m-lll I E c5-320 0.3 - - - - - -= F Cg-400 0.7

MON - G

H C C5-250 - 0.0 0.3

I C5- 420 - 0.4

--llllllm-l------- H 175-355 -

- J 300-410 1.1

-

I I I I I I I

Figure 26

Effect of Reactor Temperature on Octane

Source: Andreawn and Upson, Ref. 47. UOP 2008-26

response varies from 0.3 to 1.1 RON/lO°F reactor temperature increase. As seen in Figure 27, also taken from reference 47, a good correlation exists when the gasoline mid-boiling point is plotted against RON. These data indicate that the octane of the heavy gasoline fraction is more responsive to reactor temperature changes than is the octane of the light gasoline fraction. This difference in response can be attributed to the differences in olefinicity between the two fractions. The light gasoline fraction is highly olefinic and changes in reactor temperature have little effect on its composition. However, the heavy gasoline fraction is much less olefinic. An increase in reactor temperature can produce a significant increase in olefin content in the heavy fraction and hence an octane increase.

425

Figure 27

Effect of Gasoline Boiling Range on Octane Response

1 .o

0.0 I I I I I I I I I

150 175 200 225 250 275 300 325 350 Gasoline 50% Point, "F

Source: Andreason and Upson. Ref. 47. UOP 2008-27

Published data indicate that conversion also has a strong impact on gasoline composition and hence on octane (44-48). Values of 0.6 to 2.0 RON increases/lO% increase in conversion have been reported. An illustration of the changing hydrocarbon composition with conversion is shown in Figure 28. At high conversion, aromatics and paraffins (which to a large degree are isoparaffins) are increasing rapidly, and naphthenes and olefins are decreasing. This change in composition resulting from conversion effects causes a change in octane, particularly at high conversion, as illustrated in Figures 29 and 30 (62).

Figures 29 and 30 indicate that the increase in motor octane resulting from a conversion increase is approximately equal to the increase in research octane. This effect is in contrast to the effect of increasing temperature, where the RON increase is nearly twice as great as the MON increase. This equality of effect between RON and MON as a result of conversion was also noted by Witoshkin, et al., who quoted a typical value for octane increase as 0.8/10% increase in conversion for both RON and MON(49).

2. Catalyst Circulation Rate

Changes in the catalyst circulation rate can occur for a variety of reasons. In many FCCU's, increasing the catalyst circulation rate is the control procedure that is used to increase riser temperature. A decrease in feed preheat causes the catalyst circulation rate to increase so that a constant riser temperature is maintained.

426

Figure 28

Effect of Conversion on Gasoline PONA

30 40 50 60 70 80 90

Conversion, vol-%

Reprinted from Ref. 44. UOP 204)-28

Likewise, an improvement in unit coke selectivity, either because of a catalyst or a feed quality improvement, increases catalyst circulation rate. In this case, the improved coke selectivity produces a decrease in regenerator temperature, which then results in an increase in catalyst circulation rate to maintain a constant riser temperature.

At constant riser temperature, an increase in the catalyst circulation rate produces in the riser an increase in the amount of catalyst present relative to the oil it is contacting (increase in cat/oil ratio). The result is an increase in conversion. Murphy and Cheng indicate that a typical response in a commercial-riser FCCU operating at 980°F riser temperature, at the 75 wt-% conversion level, and in the 6- 7 cat/oil region, would be an increase of 3 to 4 wt-% conversion for an increase of 1.0 in the cat/oil ratio (41). For the special case of cracking a hydrotreated heavy VGO with a zeolitic catalyst in a short contact time (4.5 sec residence time) system, Venuto and Habib have reported that doubling the cat/oil ratio typically results in an 8 to 10 vol-% increase in conversion (50).

3. Reactor Pressure

Reactor pressure is theoretically an independent variable. In actual practice, however, the operator has little latitude in independently varying reactor pressure.

427

82

81

80

Figure 29

MAT RON Measurements

-

-

-

- I I I I

92

91

90

60 70 Conversion, wt-%

80

Reprlnted from Andreasson and Upson, Ret. Q.

UOP 2008-29

60 70 Conversion, wt-%

80

Reprinted tram Andreuson and Upson, Ret. Q.

UOP m-33

428

The reactor pressure, which is normally set by the limitations of the gas compressor, varies as the gas compressor suction conditions are changed. Thus, the feed rate to the FCCU and the reactor conversion level, items that influence the pressure drop through the main fractionator and the amount of gas handled by the compressor, largely dictate the FCCU pressure.

Of more interest than absolute pressure is the hydrocarbon partial pressure in the reactor. To improve the initial contacting of catalyst and oil, most modern FCCU's inject steam into the base of the riser to help disperse the feed as it enters the riser. 'Qpically, a dispersion steam equivalent of about 2 wt-% of the feed is used, which on an average volume basis through the length of the riser, amounts to about 25% of the total volume in the riser. Thus, the use of dispersion steam has a significant effect on the hydrocarbon partial pressure in the riser.

Directionally, increasing hydrocarbon partial pressure has much the same effect as increasing the H-transfer characteristics of the catalyst: conversion increases, coking tendency increases, olefinicity decreases, and octane decreases. Schlossman et al. studied the effect of partial pressure in a pilot plant riser-cracking unit by varying unit pressure (51). Figures 31 and 32 illustrate the decrease in C, and C, olefinicity that they found when partial pressure was increased. They also saw a decrease in gasoline olefinicity, from which a threefold increase in partial pressure produced a 2 number decrease in RON (Figure 33). Similarly, Schwanenbek et al. reported a 2 RON drop for a 5 psi increase in oil partial pressure (52). Because motor octane is less affected by olefinicty changes, it is less affected by pressure changes. A 0.5 decrease in MON for a 5 psi increase in oil partial pressure was reported by LeRoy (53). Murphy and Cheng reviewed a variety of pilot plant and commercial operations and determined a typical response to a 5 psi increase in hydrocarbon partial pressure (Table 10) (41).

4. Carbon on Regenerated Catalyst

Catalyst coke formation with zeolitic catalysts is believed to be due to H- transfer from unsaturated species adsorbed on the catalyst surface to gas-phase olefins. The adsorbed species are the source of hydrogen used to saturate the other olefins and of carbon that then becomes coke (54). Some of this carbonaceous material remains adsorbed at the active sites and prevents those sites from further catalytic participation. Other carbonaceous species cyclize to aromatic compounds, which can then diffuse to the exterior of the zeolite crystal and block access to the pores leading to the active sites (55). Both mechanisms result in a decrease in catalyst activity as coke builds up on the catalyst. This effect has been frequently reported in the literature (56-58). Wachtel et al. reported that increasing carbon on regenerated catalyst (CRC) from 0.1 to 0.4 wt-% produced a decline in conversion of approximately 3.2 vol-%/O.l% change in CRC (57). Upson (Figure 34) showed a somewhat lower response of about 2.4 wt-%/O.l% change in CRC (58).

429

Figure 31

Effect of Hydrocarbon Partial Pressure on C3 Yields

12 -

10 -

70 72 74 76 78 80

Conversion, Vol-96

Reprinted from Schloatmn, et al., Ref. 51.

Figure 32

UOP 2008.31

Effect of Hydrocarbon Partial Pressure on C4 Yields

19

17

15

13

11

9

7

HCPP- minimum

2 x min.

3 x min.

Butenes

70 72 74 76 78 80 Conversion, Vol-%

Reprinted fmm Schlmrrmn. eta]., Ref. 51. UOP Po89

430

Figure 33

Effect of Hydrocarbon Partial Pressure on Octane Number

70 72 74 76 78 80 Conversion, Vol-%

Reprinted from Schlosmnun, eI al., Ref. 51.

UOP m-33

Figure 34

Effect of CRC on Equilibrium Catalyst Activity

~~~~~~ - - 0.0 0.1 0.2 0.3 0.4 0.5

wt-% Coke on Regenerated Catalyst (CRC)

Souce: Upwn, Ref. 58.

W P 2008-34

43 1

Table 10 The Effect of pressure on FCCU Product Yields

11 ALPG

I[ ACoke

11 ARON

11 AMON

Response to a 5 Psi Increase in Partial Pressur

+ 1.2 wt-% FF

- 0.16 wt-%

No Change

+ 0.36 wt-%

+ 0.46 wt-%

- 0.7

- 0.6

The presence of carbon on the cracking catalyst also influences cracking selectivity. Wachtel et al. have reported that increased carbon level results in a loss of gasoline selectivity, a corresponding increase in C, and lighter products, and an increase in catalytic coke formation (57). Schlossman et al. also found that gasoline selectivity decreased as CRC increased (Figure 35) (51).

5. Feed Distribution

Improperly mixing hot catalyst coming from the regenerator with feed at the point of feed inlet to the riser can lead to excessive amounts of thermal cracking and the occurrence of unwanted side reactions resulting from backrnixing. The design of efficient catalyst-feed dispersion and mixing systems was discussed earlier in this chapter.

Mauleon and Courcelle have reported that a good feed distribution system can reduce the adverse dehydrogenation properties of contaminant metals (59). They reported that the installation of such systems has resulted in delta coke reductions ranging from 0.05 up to 0.20 wt-%.

As described earlier, UOP uses a combination of design features in the lower portion of the riser to provide intensive mixing and contacting of catalyst and FCCU feed. The catalyst is first dispersed and preaccelerted up the riser, prior to feed introduction, by injecting lift gas at the base of the riser. The lift gas is either steam, a light hydrocarbon gas (preferably C, and lighter) or a combination of the two. The FCCU feed is mixed with steam and injected into the dispersed catalyst stream through a multiplicity of medium-to-high pressure nozzles located around the circumference of the riser. Cabrera et al. reported on the feedstock properties and operating conditions of a commercial operation before and after such a feed distribution system was installed (Table 11) (39).

432

3 0.110.2 0.110.6

Figure 35

Effect on Gasoline Selectivity of Carbon on Regenerated Catalyst

383 448 5 17

63 t

398 444 524

Table 11 The Effects of Using Lift Gas

Operating Conditions

Feed Rate

I

I I I I I I

70 72 74 76 78

Reprinted from Schlossman, et al., Ref. 51. Conversion, Vol-%

Before Lift Gas

Base Base

CRC

0.0

0.2

0.3

-

Reg. Temp., "F

Feed Temp, "F

11 Feedstock Properties I Before Lift Gas

1344 1362

464 47 1

Specific Gravity

Sulfur, wt-%

Con. Carbon, wt-%

Metals Ni/V, ppm

Distillation, "C 10% 50% 90%

0.9190 I 0.9167

11 Rx Temp., OF I 984 I 986

UOP 2008-55

After Lift Gas

0.9178 I 0.9210

0.110.2 0.110.2

After Lift Gas

Base t 15% I Base + 15%

984 I 986

433

Here it can be seen that the installation of lift gas allowed the refiner to increase throughput by 15% and decrease regeneration temperature 37°F on average. Figure 36, illustrates the lower regenerator temperature and lower gas production achieved after the lift gas system was installed (39).

Figure 36

Effect o f l t f i Gas I

Post-Lift as +I I

1 % 380 370 I

Q 360 I '

- & - $ 350 3 340 $ 330

320 A a I 310 I

[ 4.0

B d 3.5 F 3 $ 3.0 e n

2.5

I I

Post-Lift as +I I 1

I I - A A I 4. I I

- 1 I ! I

1290 1320 1350 1400 Regenerator Temperature, O F

UOP 1mc-10 UOP 2008-36

C. Effect of Raenerstor on Process Yields

1. Regenerator Temperature

Just as poor mixing at the contact point between catalyst and feed has been found to hurt product yields, so too does excessively high regenerator temperatures

434

Test A

hurt product yields. Even with good riser mixing, high regenerator temperatures promote an increase in the ratio of thermal/catalytic cracking at the base of the riser. Schlossman et al. reported for gas oil cracking that gasoline selectivity decreases with increasing regenerator temperatures above 1250°F (51). They reported that for a majority of gas oils, as well as for resid cracking, the optimal regenerator temperature is about 1300°F. Unpublished commercial data from UOP (Table 12) illustrate this further. Test A, which had a 37°F lower regenerator temperature than Test B, had both a higher gasoline yield, a higher gasoline selectivity, and a relative 11% decrease in light gas yield.

Test B

Table 12 Commercial Catalyst Cooler Evaluation

Feed HTD VGO/VAC Bottoms

Feed Temperature, "F

Riser Outlet Temperature, O F

688 63 1

984 985

c, t c,, LV-%

Gasoline, 90% @ 380"F, LV-%

LCO, 90% @ 600"F, LV-%

c o , LV-%

Coke, wt-%

Regen. Temperature, "F 1,305 1,342

Cat/Oil, wt/wt

Cooler Duty MM Btu/hr

Yields:

28.7 28.5

64.4 62.8

11.2 11.6

7.5 8.4

5.1 4.8

Total Liquid Yield, LV-%

Gasoline Selectivity, %

11 Conversion, LV-%

111.7 11 1.2

79.2 78.5

1 81.3- I 80.0

435

2. Regenerator Heat Balance Interaction

In a typical commercial FCCU operating at a regenerator temperature constraint, the effect of raising reactor temperature can be complex. As noted earlier, raising reactor temperature increases conversion, which in turn increases the yield of catalytically produced coke. This situation results in an unacceptable increase in regenerator temperature, which requires corrective action if the increase in riser temperature is to be sustained. A common response is to decrease the air rate to the regenerator, thereby increasing the coke content on the catalyst leaving the regenerator, and decreasing the catalyst activity, as described in Section IV.B.4.

Thus, by decreasing the activity of the catalyst, unit conversion can be decreased, catalytic coke can be decreased, and regenerator temperature brought back to the allowable limit, and riser temperature can be maintained at the higher level. At this new steady-state condition, conversion has been increased, because the catalytic coke-producing reaction has a positive activation energy: -2500 Btu/lb mole, as reported by Wollaston, et al., (60). Thus, at the higher riser temperature, catalytic coke make at constant conversion is lower and consequently, at the overall coke yield that satisfies the heat balance, conversion can increase.

Figures 37 and 38 show this complex relationship when an increase in reactor temperature is countered by a decrease in catalyst activity to hold regenerator temperature constant (61). As seen in Figure 37, although conversion monotonically

Figure 37

Conversion, Gasoline yields at Constant Regenerator Temp. (756OC)

f

Reactor Temp. "C

519 517 515

70.0 70.4 70.8 71.0 71.2 71.6 72.0 MAT Activity 7

" 40.0

Reprinted from Wichers and Upson, Ref. 61. UOP 2041.37

436

increases as temperature increases, gasoline yield goes through a maximum. At the higher temperatures, the negative aspects of poorer gasoline selectivity at the higher temperature drive the gasoline yield down. Figure 38 illustrates the effects on LPG yields. The increasing conversion and increasing temperature produces a steady increase in olefins and saturates, but as temperature increases the olefin increase is proportionately greater than the saturate increase.

Figure 38

Isobutane and Propylene + Butylene yields at Constant Regenerator Temperature

8.5

8.0 -

-

3.0

2.5

2.0

Reactor Temp. "C

524 519 515 510 I I I I

MATActivity ____)

70.0 70.4 70.8 71.0 71.2 71.6 7 .O

Rcprlnbd from Wlchm and Upon. bf. 61.

V. RECENT AND FUTURE TRENDS

The design and operating procedures of the FCCU have changed substantially during the 50 years since the first introduction of the process in 1942. Todays modern FCCU's feature all riser cracking, with only a brief contact period between catalyst and hydrocarbons (-3 seconds) followed by rapid disengagement, to minimize nonselective secondary reactions. Rapid and intimate contact between hot regenerated catalyst and the hydrocarbon feed, through the use of multinozzle feed atomizers at the inlet end of the riser, and high efficiency, multi-stage strippers to recover product vapors from catalyst leaving the riser are additional features of the modern FCCU reactor.

Major design changes have occurred on the regenerator side as well. Two-stage regenerators have proven to be an effective means for controlling regenerator temperatures when highly coke forming feeds are being processed. The development of reliable and versatile catalyst coolers have also been beneficial in processing such

43 7

feeds and have added an extra degree of flexibility to the FCCU. For more conventional VGO type feeds, the development of the fast fluidized combustor style regenerators has resulted in smaller and more efficient units.

On the process side, the trend has been towards higher reactor temperatures and lower reactor pressure. Both of these move the FCCU product yields in the direction of increased production of C, and C, olefins and increased gasoline octane. Regenerator operations have tended towards burning the regenerated catalyst to low levels of carbon (<0.1 wt-%) to maximize catalyst activity and selectivity, but, at the same time, keeping the regenerator temperature below 1350°F to minimize the adverse selectivity affects that result from mixing hot catalyst with ,the colder FCCU feed.

These trends are expected to continue into the future. The U.S. refining industry will be profoundly affected by the demands placed upon it to produce the reformulated fuels mandated by the Clean Air Act (see Chapter 15). Increased quantities of isobutylene and isoamylene will be required to produce methyl tertiary butyl ether (MTBE) and tertiary amyl methyl ether (TAME) to meet the oxygenate needs of reformulated gasoline. The industry will also be looking for increased C, olefin yields in general, to convert to further quantities of iC,= and possibly for increased feed to alkylation units. Since the use of oxygenates will add substantial volume to the gasoline pool, while U.S. gasoline demand is projected to grow only slightly in the next ten year period (64) there will be a companion need for refineries to reduce their production of more conventional type of gasolines. Higher severity FCCU processing (higher temperature, higher conversion) fits these requirements nicely. By operating at these conditions, an increase in olefinic LPG is achieved at the expense of reduced yields of FCCU gasoline. Thus, the trend toward higher temperature FCCU processing is expected to continue in the US., to meet reformulated gasoline needs. Refiners outside the US. will most likely be faced with similar environmental demands in the coming years, and hence FCCU operations around the world are also expected to be moving towards higher severity operations.

The processing of heavier and dirtier feed is also to be anticipated in the future. For refineries which do not need to meet reformulated gasoline type product yields, it is expected that the trend towards the use of two stage regenerators and the liberal use of catalyst coolers will continue, facilitating the processing of high metals, high Conradson carbon feeds in the FCCU.

For refiners who must meet the low sulfur requirements that will be imposed on future reformulated gasoline and diesel fuel, it may be that the FCCU product can not meet these specifications unless the feed to the FCCU is severely pretreated. In such a case, FCCU feed quality will be excellent, producing high yields of light olefins and low sulfur gasoline.

438

M.

1.

2.

3.

4.

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7.

8.

9.

10. 11.

12.

13.

14.

15. 16.

17. 18.

19.

20.

21.

22. 23.

24.

REFERENCES

A. A. Avidan, F. J. Krambeck, H. Owen, and P. H. Schipper, paper AM 90-33 presented at 1990 NPRA Annual Meeting, San Antonio, Mar. 25-27, 1990. C. W. Strother, W. L. Vermilion, and A. J. Conner, paper presented at API 37th Midyear Meeting, New York, May 8-12, 1972. J. D. Heldman, F. Kunreuther, J. A. Marshall, and C. A. Rehbeim, API Proc. Vol36, 1956. C. L. Hemler, paper presented at AIChE Summer National Meeting FCC Tutorial, Philadelphia, Aug. 22, 1989. J. A. Finnegan, J:R. Murphy, and E. L. Whittington, Oil & Gas J., Jan. 14, 1974. W. L. Pierce, D. F. Ryan, R. P. Southern, and T. G. Kaufmann, API 37th Midyear Meeting, New York, May 10, 1972. J. E. Naber, P. H. Barnes, and M. Akbar, Japan Petroleum Institute Petroleum Refining Conference Tokyo, Oct. 19-21, 1988. G. D. Myers, P. W. Walters, and R. L. Cottage, US. Patent 4,066,533, Jan. 3, 1978. C. A. Cabrera and D. Knepper, paper AM 90-39, presented at 1990 NPRA Annual Meeting, San Antonio, Mar. 25-27, 1990. A. L. Saxton and A. C. Worley, Oil and Gas J., May 18, 1970. W. D. Ford, G. J. D'Souza, J. R. Murphy, and A. A. Murcia, API 43rd Midyear Mtg. Proc., Vol. 57, 1978. R. R. Dean, J-L. Mauleon, and R. W. Pfeiffer, U.S. Patent 4,331,533 May 25, 1982. L. E. Busch, W. P. Hettinger, Jr., and R. P. Krock, paper AM 84-50, presented at the NPRA Annual Meeting, San Antonio, Mar. 25-27, 1984. J-L. Mauleon, J-B. Sigaud, and G. Heinrich, paper presented at Japan Petroleum Institute Petroleum Refinery Conference, Tokyo, Oct. 27-28, 1986. D. P. Bunn, Jr., et al., # e m . Engr. Progr., 65, 6, June 1969. M. C. Bryson, G. P. Huling, and W. E. Glausser, Oil and Gas J., May 15, 1972,

L. 0. Sthe, and J. B. Pohlenz, U.S. Patent No. 3,489,673, Jan. 1970. T. Ham, F. Nakashio, and K. Kusunoki, "The Burning Rate of Coke Deposited on Zeolite Catalysts," J. of Chem. Eng. of Japan 8, 2 1975, pp. 127-130. L. L. Upson and H. van der Zwan, "Promoted Combustion Improves FCCU Flexibility," Oil and Gas J., Nov. 23, 1987, pp. 65-70. P. B. Weisz and R. B. Godwin, "Combustion of Carbonaceous Deposits with Porous Catalyst Particles, 11, Intrinsic Burning Rate," J . of Catalysis 6, 1966, pp.

J. R. Arthur, "Reactions between Carbon and Oxygen," Transactions of the Faraday Sociery 47, 1951 pp. 164-178. J. E. Pennick, US. Patent 4,064,039, Jan. 1978. A. B. Schwartz, US. Patents 4,072,600, Feb. 7, 1978, and 4,093,535, June 6, 1978. Y. Guegan, "Use of Combustion Promoter," paper presented at Katalistiks 1st Annual FCC Symposium, Bordeaux, Oct. 1980.

p ~ . 97-101.

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L. L. Upson, "Catalytically Promoted Combustion Improves FCC Operations," paper No. AM 79-39 presented at the 1979 NPRA Annual Meeting, San Antonio, Mar. 1979. J. Yerushahni and N. T. Cankurt, "Further Studies of the Regimes of Fluidization," Powder Technology 24, 1979, pp. 187-205. D. Geldart. "Types of Gas Fluidization," Powder Technology 7 , 1973, pp. 285. J. Yerushalmi and N. T. Cankurt, "High Velocity Fluidized Beds," Chem Tech 8, 1978, pp. 564. N. L. Giuricich and B. Kalen, "Dominant Criteria in FCC Cyclone Design," paper presented at Katalistiks 3rd Annual FCC Symposium, Amsterdam, May 1982. "Maximum Dilute Phase Entrainment in Vertical Gas-Solids Flow," API Publication 931, May 1975, Chapter I1 - Cyclone Separators. A. S. Krishna and E. S. Parkin, "Modeling the Regenerator in Commercial FCC Units," Chem. Eng. Prog. Apr. 1985. E. Singer, D. B. Todd, and V. P. Guinn, "Catalyst Mixing Patterns in Commercial Catalytic Cracking Units," I@. C., 49, 1, Jan. 1957. K-E. Wirth, "Axial Pressure Profile in Circulating Beds," Chem Eng Tech 1988,

D. Kunii, K. Yoshida, and 0. Levenspiel, "Axial Movement of Solids in Bubbling Fluidized Beds," paper presented at I. Chem Eng. Tripartie Chem Eng. Conference - Fluidization Series, Montreal Sept. 1968. C. A. Cabrera and R. W. Mott, "High Efficiency Regenerator, Theory of Operation and Commercial Experience," paper presented at Katalistiks 3rd Annual FCC Symposium, Amsterdam May 1982. J-L. Mauleon, J-B. Sigaud, and G. Heinrich, "FCC-Heat Balance Management with Heavy Feeds: Mixed Temperature Control Approach," paper presented at Japan Pet. Inst., Petroleum Refining Conference, Tokyo, Oct. 1986. C. Wen and M. Leva, "Fluidized Bed Heat Transfer," AIChE Journal 2, 4, 1956, pp. 482. W. J. Danziger, "Heat Transfer to Fluidized Gas-Solids Mixtures in Vertical Transport," I&E.C. Proc. Des. and Dev. 2, 4, Oct. 1963, pp. 269-276. C. A. Cabrera, C. L. Hemler, and S. P. Davis, "Improve Refinery Economics via Enhanced FCC Operations," paper presented at Katalistiks 8th Annual FCC Symposium Budapest, June 1987. L. C. Yen, R. E. Wrench, and C. M. Kuo, T h e Important Role of Regenerator Temperature on Catalytic Cracking," paper presented at Katalistiks 6th Annual FCC Symposium, Munich, May 1985. J. R. Murphy and Y. L. Cheng, paper presented at Katalistiks 5th Annual FCC Symposium Vienna, May 1984, pp. 11.1-11.12. H. Owen, P. W. Snyder, and P. B. Venuto, Proceedings of Sixth Int. Congress on Chtalysis 2, 1977, p. 1071. L. L. Upson and R. Sikkar, Applied Chtalysis 2, 1982, pp. 87-105. L. L. Upson, E. E. Winfree, and E. L. Leuenberger, paper AM 78-50 presented at the March 1978 NPRA Annual Meeting, San Antonio. J. S. Magee, R. E. fitter, D. N. Wallace, and J. J. Blazek, Oil and Gas J. , 78,

pp. 11-17.

32, 1980, pp. 63-67.

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46. F. Valeri, paper presented at Katalistiks Symposium on Octane, Amsterdam, Jan. 1986.

47. H. U. Andreasson and L. L. Upson, Oil and Gas J., 83, 32, 1985, pp. 91-96. 48. W. J. Reagan, G. M. Woltermann, and S. M. Brown, ACS Symposium on

Advances in Catalytic Cracking, Washington, DC, Sept. 1983. 49. A. Witoshkin, G. S. Koermer, and R. J. Madon, paper AM-88-46 given at the

March 1988 NPRA Annual Meeting, San Antonio. 50. P. B. Venuto and E. T, Habib, Fluid catalytic Cracking With Zeolite Catalysts,

Marcel Dekker Inc., New York, NY, 1979, p. 126. 51. M. Schlossman, S. E. Ronczy, R. E. Wrench, and L. C. Yen, paper presented at

Katalistiks 7th Annual FCC Symposium May 1986, Venice, Italy. 52. E. F. Schwarzenbek, C. E. Slynstad, P. T. Attridge, and J. W. Jewel, Proc. 3rd

World Petroleum Congress, Sec. IV, 1951, p. 166. 53. C. F. LeRoy, 1986 NPRA Q and A Session on Refinery and Petrochemical

Technology Gerard L. Ferrar & Assoc. Tulsa, 1986, p. 41. 54. B. W. Wojciechowski and A. Coma catalytic Cracking, Catalysts, Chemistry,

and Kinetics Marcel Dekker Inc., NY, 1986, p. 183. 55. B. E. Langner, Ind. Eng. Chem.. Process Des. and Dev. 20, 326, 1981. 56. J. A. Montgomery, Oil and Gas J . 70, 50, 1972, p. 81. 57. S. J. Wachtel, L. A. Baillie, R. L. Foster, and H. E. Jacobs, Prepr. Div.

Petroleum Chem., Amer. Chem. Society, 16, 3, 1971, A-55. 58. L. L. Upson, paper No. AM 79-39 presented at the March 1979 NPRA Annual

Meeting, San Antonio, Texas. 59. J-L. Mauleon and J. C. Courcelle, paper presented at Katalistiks 6th Annual

FCC Symposium, Munich, Germany, May 1985, p. 9.5. 60. E. G. Wollaston, W. J. Hafin, W. D. Ford, and G. J. DSouza, Hydrocarbon

Processing 54, 19, 1975, p. 93. 61. W. R. Wickers and L. L. Upson, Oil and Gas J., 82, 12, 1984, pp. 157-164. 62. H. U. Andreasson and L. L. Upson, paper presented at Katalistiks 6th Annual

FCC Symposium, Munich Germany, May 1985, pp. 3.9-3.11. 63. I. B. Cetinkaya and R. P. Culler, U.S. Patent 4,927,606, May 26, 1990.


CHAPTER 12

THE EFFECT OF FEEDSTOCK ON YIELDS AND PRODUCT QUALITY

WARREN S. LETZSCIT AND ANTHONY G. ASHTONb

a Refining Process Services, Ellicott City, Maryland BP International Limited, Sunbury-on-Thames, Middlesex, England

I. INTRODUCTION

The composition of the feed to an FCCU is the largest single influence on the resulting yields and product qualities. Recognition of this fact has compelled unit designers and operators to search for methods to predict unit performance on the basis of feedstock analysis.

The most basic approach is to crack pure hydrocarbons to obtain information on the products produced. Because the FCCU's are normally fed wide-boiling-range gas oils that contain hundreds of molecules, the single component tests have limited applicability.

Gross feedstock analyses have traditionally been used to characterize feeds. These analyses are normally simple to perform and can be correlated to the families of molecular species found in crude oils. In the most elaborate analyses, the larger groups of molecules are broken down into even smaller components, for example, aromatic molecules are separated into mono, di, tri, and so forth, aromatic species for more exact predictions. All these methods are examined in this chapter.

Residual feeds present a different set of problems. Not only are the hydrocarbons more difficult to characterize in the heavier oils, but also the heteroatoms frequently act as poisons and must be accounted for in any analysis of the feed. These topics are also discussed in this chapter.

Finally, product properties are also receiving more attention because motor fuel specifications are becoming more stringent throughout the world. The starting feed composition is related to many of the commonly measured product parameters.

11. BASIC CRACKING REACTIONS

44 1

The four major classes of hydrocarbon compounds are paraffins, olefins, naphthenes, and aromatics. Olefins are not found in crude oils but are a product of cracking reactions. The pioneering work of Greensfelder, Voge, and Good (1-6) elucidated the main chemical reactions of each class of hydrocarbon, which were summarized in Table 1 by Greensfelder (7). The primary features of the cracking reactions are:

442

Smallest products are Cis and C,'s. Largest molecules crack the fastest. Primary product is an olefin. Pure aromatics are largely unreactive. Alkyl aromatics crack at the benzene ring (chains of three or more carbons). Reactivity increases from primary to tertiary carbons but is inhibited by quarterary configuration.

of secondary reactions take place in an FCCU. The most significant ones are enumerated in Table 2. The result is an exponential increase in the number of FCCU products. Although both catalyst selection and unit operating conditions are used to manipulate these reactions, the basic character of the feed is what determines the initial concentration of each of the hydrocarbon types.

The extent of any reaction can be estimated from thermodynamics. Equation 1 relates the equilibrium constant to the driving force for the reaction, the change in free energy, or

-AF = RT In Keq (1)

If the equilibrium constant is large, the reaction is favored and vice versa. The free energy property allows the calculation of the equilibrium conditions when temperature and pressure are held constant:

F = H - T S (2)

Combining Equation 2 with Equation 1 and making these assumptions gives:

The heat of reaction, AH, is assumed to be constant over the temperature range considered. Equation 3 indicates that if AH is negative (an exothermic reaction), the equilibrium constant decreases as the temperature increases. Analogously, & increases with temperature for an endothermic reaction, which is the type of reaction associated with catalytic cracking. Table 3 Uustrates the principal reactions (both primary and secondary), gives the logarithm of the equilibrium constant, the change of the logarithm of the equilibrium constant with temperature, and the heat of reaction at 950'F (510°C) (8).

443

Table 1 Catalytic Cracklng Characteristics of Pure Hydrocarbons

n-Paraffins Extensive breakdown to C, and larger fragments. Product largely in C, to C, range and contains many branched aliphatics. Few normal 01 - olefins above C..

Isoparaffins Cracking rate relative to n-paraffins increased considerably by presence of I tertiarv carbon atoms.

1 Naphthenes Crack at about Same rate as those

paraffins with similar numbers of tertiary carbon atoms. Aromatics produced with much hydrogen transfer to unsaturates.

Little reaction; some condensation to biarvls.

Unsubstituted aromatics

Alkyl aromatics Entke alkyl group cracked next to ring

n-Olefiis Product similar to that from n-paraffins but more olefinic.

Table 2 Secondary Reactions in Catalytic Cracklng

Comwund I Reaction

Olefins Isomerhation (geometric and skeletal) Hydrogen transfer Polymerization 1 Aromatization

Aromatics Condensation

Paraffins Dehydrogenation of

Table 3 Some Thermodynamic Data for Idealized Reactions of Importance in Catalytic Cracking

2.04 1.68 12.44 11.22 0.32

0.33 1.00 0.65 2.11 0.41

-0.20

-2.21 _ _

Reaction Class

2.46 2.10 11.09 10.35 0.25

0.30 1.09 0.65 1.54 0.88

-0.23

- 1.52 -_

Cracking

Hydrogen transfer

Isomerization

Transalkylation Cyclization Dealkylation Dehydrogenation Polymerization Paraffin ablation

Specific Reaction

Log I<E (Equilibrium Constant)

Gq-GF (454°C) (5 l0T)

-- -- I

980°F (527°C)

-_ 2.23

-_ _ _

0.09 -0.36

1.10 0.65

1.05

-1.2 -3.3

--

_- --

Heat of reaction, Btu/mole, 950°F

+ 32,050 + 33,663 -109,68 1 -73,249 -4,874 -3,420 -1,310 + 6,264

-221 -37,980 + 40,602 + 56,008 _ _

_ _

Source: Reference No. 8

445

In catalytic cracking, equilibrium conditions do not necessarily occur. Although both operating conditions and hardware design affect the various reactions, some general observations about the chemistry can be made. Cracking, deakylation, and dehydrogenation are strongly endothermic. Because the first two reactions predominate at low conversions, the overall heat of reaction is also endothermic. The secondary reactions, such as hydrogen transfer, cyckation, and condensation, become more significant at higher conversions; and because they are exothermic, the net heat of reaction usually decreases. These facts may explain the large differences in LW cracking found in the literature. Also, certain catalysts may promote some of the reactions but have little impact on others.

The effect of increasing temperature is to increase the rate of cracking and dehydrogenation (olefin-producing reactions) and to decrease the exothermic hydrogen transfer reactions (olefin-consuming reactions) on a relative basis. The exothermic hydrogen transfer reactions are associated with coke formation. Thus, more olefins are formed and preserved at high reactor temperatures (RON also increases), and the regenerator temperature may not increase as much as the reactor temperature.

Figure 1 (18) shows the reactivity of several of the major classes of hydrocarbons as a function of molecular size. Although these data are for

Figure 1

Muior Reucfivify Trends of Representufive Clusses of Hydrocurbons

Total Conversion, %

*7 A D h t h e n e s I 1 00

80

60 n-Paraffins

40

20

0 4 8 12 16 20 24

. -- Monoalkyl benzenes

80

60 n-Paraffins

40

20

0 4 8 12 16 20 24

Number of Carbon Atoms

Reprinted from Ref. 8, Pg 108 by Courtesy of Marcel Dekker, Inc.

446

amorphous catalysts, the same trend is seen with zeolites. In catalytic cracking, the largest molecules crack the fastest unless their accessibility to the active sites is inhibited by steric considerations. Branched molecules react much faster than straight-chain or normal paraffins. Smaller molecules ((2,'s or smaller) are virtually inert.

In pure catalytic cracking, the smallest fragments produced are C,'s and C4)s because tertiary C-H bonds are 10 times as reactive as secondary bonds and 20 times as reactive as primary bonds. A primary (1") carbon atom is attached to only one other carbon atom, a secondary (2") is attached to two others, and a tertiary (3") to three others:

1" 2" 1" 1" 3" 1" I" 2" 2" 1"

1"

Although dealkylation of alkyl aromatics (with hydrocarbon chains of three or more carbons) occurs fairly readily, the pure aromatic molecules, such as benzene and naphthalene, are virtually inert.

Thermal cracking also occurs in an FCCU. These reactions occur by a free radical mechanism in which Gs are the main product and some C,'s and C3's occur in minor amounts. Venuto and Habib quote reactor temperatures of 1022°F (S50°C) to 1292°F (700°C) for noncatalytic thermal reactions (8). These temperatures are present at the base of the riser, where the hot, regenerated catalyst is mixed with the oil.

111. GAS OIL CHARACTERIZATIONS AND CORRELATIONS

A further breakdown of the principal hydrocarbon types was reported by White (10). The paraffins were separated into four main classes: normal, iso, monocyclic, and polycyclic. The aromatics were separated by the number of rings, ranging from one to five or more. Coke and gasoline yields are shown in Figures 2 and 3 respectively. The coke yield from the paraffins was very low.

Although White's results were for amorphous catalysts, the conclusions reached are generally valid for modern FCC operations. Gasoline yield is maximized with naphthenes (cyclic paraffins) and monoaromatics feed components. Straight-chain

447

Figure 2 Production of Coke from P olynuclear Aromatics

Yield, Wt% of Component Tetacyclics

Tricyclics

Dicydics

Severity

Source: White, P.J., Oil B Gas Joumal, 5/20/68

Figure 3 Production of C5- 430" F Gasoline

Polycyclic Paraffins Monccyclic Aromatics

Monocyclic Paraffins Yield, Wt% of Component

Normal Pataffins

Dicyclic Aromatics

Severity

Source White, PJ., Oil 8 Gas Joumal, 5/20/68

448

paraffins and isoparaffins easily overcrack to liquefied petroleum gas (LPG), and polycyclic aromatics do not produce much gasoline.

Detailed analyses of feedstocks are not always available. Therefore, the refiner must rely on simpler gross analyses to determine trends. A number of these techniques are described here.

A. SDecWc Gravity

Specific gravity is the easiest parameter to measure on any feed. It is related to degrees API by the following formulas:

"API = 141.5 - 131.5 SG = 141.5 SG API + 131.5

(4)

Gary and Hartwerk provide curves showing the effect of feed gravity on both yields and product quality as a function of conversion (1 1). The coke and gasoline correlations are presented in Figures 4 and 5. The feeds with lower API gravity and more aromatics clearly produce more coke and make much less gasoline at typical FCC conversions. Letzsch and Ashton have correlated the data from about 20 individual FCC operations (also shown in Figure 5) and found the Gary and Hartwerk lines for gasoline yield to be slightly low (19). This observation is not surprising because newer zeolitic catalysts and unit designs have improved gasoline selectivity. Specific gravity by itself may not be a reliable parameter (Figure 6). The data on a variety of feeds run at various conditions show a weak correlation between conversion and feed gravity (26). If the feed distillation varies,.the specific gravity changes, and yet the molecular character of the hydrocarbon stream may remain essentially the same. If several sources of feed are used, then densities by themselves can mask rather significant compositional changes. The UOP K factor was derived to overcome some of these shortcomings.

B. UOP K Factor

The UOP K factor is defined as:

( W P ) UOP K = SG

Where: CABP = Cubic average boiling point SG = Specific gravity at 60°F

The Watson K factor also is used to evaluate hydrocarbon streams. It uses the mean average boiling point (MeABP), which is the average of the molal average boiling point (MABP) and the CABP. This definition has changed over the years. Originally, the MABP was used but was later changed to the CABP. Then the UOP and Watson K factors were identical. Now the practice is to include the MeABP in the numerator although the K factor can be calculated directly by using correlations of viscosity, specific gravity, ASTM distillations, and molecular weight (15).

449

Figure 4

Cutulyfic Crucking Yields Zeolite Cufulyst (Coke)

Coke, Wt%

40 50 60 70 80 90

Conversion, Vol YO


Figure 5

Gusoline ws Conversion us u Function of "API Zeolite Cafalyst

Gasoline (C, - 430'

28' 2 7 24" 23' * 19" 19""

60 70 80 90

Conversion, Vol%

*Reprinted from Ref. 11, Pg 389 by Courtesy of Marcel Dekker, Inc.

450

a4

80

76 Conversion Wt% 72

68

64-

60

Figure 6

- 0

- - o m B

- n r p 0

I I 1 I I I I I I

Effect of API on Conversion

Source: Kreider, KR, e t al. AIQO Calalyst Symposium Amsterdam, 1991

A UOP K factor of 11.2 indicates a more aromatic stock. A K factor of 12.5 indicates a highly paraffinic stock. Table 4 lists the K factors for some of the common hydrocarbons found in petroleum (15).

The average boiling point is an indication of the boiling range and types of molecules. Figure 7 illustrates that large differences in the boiling point can exist for compounds having the same mass but different chemical structures (12). As the boiling point increases, the molar mass range rapidly widens. Above 1000"F, the concentration of heteroatoms such as nitrogen, sulfur, oxygen, and metals (Ni, V, Fe) increases dramatically. Consequently, small increases in gas oil end point can significantly alter the crackability of a feed.

Cracking yields as a function of K factor were first provided by Nelson for amorphous catalysts and were upgraded later for zeolite-containing products (13,14). Figure 8 is plotted from the data for zeolitic catalysts. Yields are presented at essentially constant coke yield, which represents a commercial mode of operation. Operating conditions are chosen to maximize gasoline yields, and so reactor temperatures must be moderate (925 to 950°F) for the more paraffinic feedstocks.

45 1

Table 4 Characterization Factors for Selected Hydrocarbons

Hydrocarbon

n-Octane n-Decane n-Hexadecane n-Octadecane

Feeds

Straight-chain paraffins

Naphthenes

Aromatics

Source: Reference 15.

~ ~~ ~

Cyclohexane Methylcyclohexane n-Butylcyclohexane n-Decylcyclohexane

Benzene Isobutylbenzene Biphenyl Naphthalene 2-Methylnaphthalene

K Factor

12.68 12.64 12.90 13.05

11.00 11.31 11.64 12.27

9.58 10.84 9.58 9.32 9.75

In actual operations, some overcracking usually occurs because refiners operate above the conversion level for maximum gasoline or at a higher than optimum reactor temperature because of their desire to increase light olefins and FCCU gasoline octanes.

C. Midboiline Point

The midboiling point of a feed has not been a reliable index for making yield predictions for FCC feed streams coming from several sources because the molecular character of the various streams can differ widely. For a single crude, however, the gasoline yield seems to correlate well with the mean boiling p ~ i a (Figure 9) (10). Figure 9 also indicates that cycle stocks give lower gasoline yields as the mean boiling point increases. This result is expected because the concentration of polycyclic aromatics increases considerably for this feed with higher boiling points. When more than one refinery stream is fed to the catalytic cracker or if recycle is practiced to any degree, the plot of gasoline yield versus midboiling point can go in either direction.

Figure 7

Effect of Molecular Structure on Boiling Point Molar Mass - Carbon Acyclic Alkanes C,H,, Number

422 30 28 26 24 22 20 18 16

282 0 14 12

142 10

0 t 2L 0 0 200 400 600 800 1000 1200 1400

-18 93 204 315 427 538 649 760 Reprinted with Pennissbn from Energy a Fuels, Vol, No.1 G0~1!?87Amer.chemlcalsociety

"F "C

At m osp her ic Equivalent Boiling Point

Figure 8

Cracking Yields as a Function of Characterization Factor Approximutely 5.3 Wt% Coke (After Nelson 14)

Volume YO (FF Basis)

- C, Gasoline

20 Alkylate

(Outside iC, Used)

I I I I I I I 11.2 11.4 11.6 11.8 12.0 12.2

K Factor

454

Figure 9

Effect of Midboiling Point on Gusoline Yield

Yield @ I 50% Conversion

500 600 700 800 980 1000

Midboiling Point, O F

Source: White, P.J., Oil and Gas Journal, May 20,1968

Higher gasoline yields are generally observed as feed boiling point increases (10,16,21). However, extrapolating this trend past 1000 to 1050°F end-point feeds is dangerous, because coke precursors in the heavy end of the feed will adversely affect both the yields and heat balance of the commercial unit.

The lighter feeds do not crack as readily as the heavier feeds and make less gasoline over low activity catalysts (16). Present-day high-activity zeolite catalysts (with zeolite contents of 35 wt-% or more) can crack these small molecules effectively and comvert them to gasoline and LPG. Propylene yields are usually enhanced when material boiling below 650°F is added to the feed, but gasoline octanes usually decline.

D. Molecular Weight

Molecular weight is a useful number to know for a hydrocarbon feedstock but has the same limitations by itself as specific gravity and K factor as a correlating parameter for feeds. The molecular weight is estimated from the viscosity of the feed and its specific gravity using the following equation (15):

455

(-1.2435 t 1.1228SG) M = 223.56 KV,,

(3.4758 - 3.038SG) -0.6665 x KVm x SG

Where: M = Molecular weight KV,, = Kinematic viscosity at 100"F, centistokes KV,,, = Kinematic viscosity at 210"F, centistokes SG = Specific gravity, 60°F

The molecular weight is needed to calculate molar flow rates, which in turn allow the volume of vaporized oil to be calculated at the base of the riser. Cracking models and gas-phase oil volumes are used to set the sizes of the reactor and stripper and the upper sections of the riser to accommodate molar expansion as the feed is converted to lower molecular weight products (17).

E. Hvdroge n Content of Hvdroca rbons

Another feed parameter used to predict yield performance of the catalytic cracking unit is the hydrogen content of the feed. This parameter is more fundamental to the feed molecular structure than specific gravity or K factor and has been applied to general refining processing as well as to fluid catalytic cracking.

An overall hydrogen balance gives insight into the types of reactions needed to produce a desired product slate from known feedstocks (crude oils). One of the more illustrative examples of the carbon-hydrogen approach is the Strangeland Chart (Figure 10) (27,32). Hydrogen-to-carbon ratios for pure paraffins, olefins, naphthenes, and aromatics are shown along with a number of petroleum fractions. Lines drawn for Arabian Heavy and the syncrudes from shale and coal point out the hydrogen deficiencies of the latter stocks. Carbon numbers and approximate boiling points help put the chart in perspective.

A complementary chart (Figure 11) illustrates the differences between the primary resid conversion processes (coke excluded). The amounts of the respective products from each process are not shown, but the basic character of each product, as related to hydrogen content, is readily seen. Refinery processes have been frequently categorized into those that add hydrogen and those that reject carbon. Carbon rejection is the removal of hydrogen-deficient hydrocarbons containing from 4 to 6 wt-% hydrogen. Either route can be used to produce products with the desired hydrogen contents.

A more comprehensive review of residual upgrading processes separates the processes into hydrogen or nonhydrogen, catalytic or thermal, cracking or treating and reactor type (36). The conditions used in these refining steps, shown in Figure 12, illustrate the unique position catalytic cracking holds in the upgrading schemes.

456

Figure I0

The Strangeland Chart for Fuels

HiC Atomic Ratio

0 10 20 30 40 50

Carbon Number

Source: Bridge, AG., et al., Oil and Gas Journal, Jan. 19,1981

Figure I I

Comparison of Primary Resid Conversion Products on the Strangeland Chart

HiC Atomic Ratio

Paraffins

,+ (Approximate Boiling Point)

0 10 20 30 40 50

Carbon Number Sourca Bridge, AG., et el., Oil and Gas Journal, Jan. 19,1981

457

Figure 12

Conditions Used in Upgrading Processes Pressure, Bars -

t T E M P E R A T U R E *C

1

r Without Hydrogen ~lennpmc.-~ari

& Hydro- 100 200 300

Dsv Hydrovisbmak

With Hydrogen

N 0 N C A T Y L I T I C

\

Reprinted by Permission from Hydrocarbon Processing, Feb., 1984 All Rights Reserved

Hydrogen contents of the major classes of hydrocarbons are given in Tables 5 and 6. Paraffins, naphthenes, aromatics, and combinations of these are the normal FCC feed components. These same groups plus the olefins are also the products. Hydrogen content calculations for coke frequently give values of 5 to 7 wt-%. From Table 5 these values correspond to condensed aromatics with 4 or 5 rings and explain why White found that coke was the main product of cracking such compounds (10).

The hydrogen content of petroleum products has been estimated by Hinds and is shown in Table 7 (33). Also listed are the maximum theoretical yield of each product from a straight-run residue containing 12 wt-% hydrogen. High yields of all components except LPG are possible. Catalytic cracking was found to be more efficient (85 to 90%) in using the available hydrogen than was thermal cracking or coking (65 to 75%).

The absolute amounts of the various hydrocarbon compounds actually found in crude oils and their various fractions can vary over a wide range (Figure 13). An approximate carbon number and molecular weight are given for the cuts from six different crude oils. A closer examination of midcontinent crude (38) indicates large amounts of saturates throughout the gasoline and distillate fractions, but a

Table 5 Hydrocarbon Ratio of Aromatics, Mixed Molecules, and Olefins

Mixed Molecules ll

Olefins:

c, wt -

72

156

264

360

120

168

12N

H, wt -

6

10

14

18

12

14

2N

H, w t 4

7.7

6.0

5.0

4.8

9.1

7.7

14.3

H/C, Qtomic Ratio

1.000

0.769

0.636

0.600

1.200

1.000

2.000

459

Table 6 Hydrocarbon Ratio of Parafins and Naphthenes

Paraffins:

Naphthenes:

0

c , wt

-

12 24 60 288 576

72

168

360

H, wt

-

4 6 12 50 98

12

24

48

H, wt-%

25.0 20.0 16.7 14.8 14.5

14.3

12.5

11.8

WCl Atomic Ratio

4.000 3.000 2.400 2.083 2.042

2.000

1.714

1.600

Table 7 Hydrogen Content of Petroleum Products

Min., wt-% Avg., wt-% Max. Pot. Yield, wt-%*

Petrochem. olefins Gasoline 14 - 15 Aircraft turbine fuel 13.8 14.0 - 14.4 Distillate fuel oil 12.5 12.5 - 14.0 Bunker fuel 7 - 10 100

Source: W. P. Hinds, Jr., Proc. of 8th World Petroleum Congress, Vol. 4,

* Yield is from 12% H-residue. Residue is assumed to be converted into 1971.

indicated product and coke containing 4 wt-% H.

Figure 13

Moleculur Weight Distribution of Petroleum Components as u Function of the AEBP

cdu&Qll ~ ~ u . ! u % ~ m ~ m Almont 39.2 79.1 91.1 97.9 96.4 96.5 98.6 Arablan H e m 38.3 68.6 03.2 92.7 97.7 99.6 100 Maw 36.0 65.7 77.2 86.7 91.0 99.4 100

Weight' Number Bman 14.5 36.4 57.2 77.8 89.0 96.1 100

Molecular Carbon OlfshoreCal. 36.1 60.8 76.4 86.4 92.5 99.3 100 Kern River 18.0 56.7 79.1 92.3 97.0 99.5 100

2522

2242 1962 1682 1402 1122

562 282

0

a42

180

160 140 120 100 80 60 40 20 0 Atmospheric 0 300 650 1000 1300 1700 O F Equivalent

-1 8 149 343 538 704 927 "C Boiling Point 'Acyclic Alkanes C,H,, 'Cumulative WP/. from Crude Oil Reprinted with Permission from Joumel of Energy and

Fuels, 1987,1,2 Copyright 1987 American Chem. SOC.

Figure 14a

Distribution of Various Petroleum Compound Classes in a T'picalMedium Crude Oil After Bestougeff, 1967

Light Oil Lub. Oil Distllhta Dktlilata Gasdll-lO Kerosene Gas011

I I I

100 150 200 250 300 400 500 M . + T ~ ~ ' C

Cyclo-Al kanes

~

Source: Tissot and Welte, "Petroleum Formation and Occurance" Spnnger-Verlag Pg 281 (1 978)

0

25

50

75

100%

Figure J4b

Compound-Cluss Distribution in Kern River Crude

100 90 80 70 60 50 40 30 20 10 0

- - -

10 30 50 70 90

Concentration, Wt%

Reprinted with Permission from Joumel of Energy and Fuels 2,1988 Pg 597 Copyright 1988 American Chemical Society

463

heavy (13.6 "API) Kern River crude (34) has little gasoline and is considerably more aromatic than the paraffinic-naphthenic crude from the U.S. midcontinent. Figures 14a and 14b (pages 461 and 462, respectively) clearly differentiate these crudes. The 1050°F bottoms from the heavy crude contain more than 85% aromatics and polar compounds and would make a rather poor FCC feedstock.

The principal yields from the FCCU as a function of feed hydrogen content are shown in Figure 15. These yields are optimized for maximum gasoline yields from the gas oil feed. Most commercial operations running virgin feedstocks have hydrogen contents ranging from 11.8 to 12.8 wt-%. Blending feedstocks makes this type of correlation less accurate, particularly if the feedstocks are precracked or are extracts.

A major advantage of the hydrogen approach is that a hydrogen balance can be done on the feed and products (39). Base yields are used to establish the balance for a particular feedstock and unit. Predictions are then made where the hydrogen in the products matches the base case. Thus, predictions are much more realistic and prevent estimates that show slurry oil being converted into gasoline.

Figure 15

Yields (% Fresh Feed' YS Hydrogen Content Feed

Vol % wt Yo

Hydrogen, Weight YO Fresh Feed

464

Authors

ASTM D 3343-74

Hydrogen content or carbon-hydrogen ratios are parameters that were first used to estimate the heat of combustion for petroleum fuels. They are measured in the laboratory (ASTM DlOl8) or by nuclear magnetic resonance (NMR) or can be determined by correlation using such feed parameters as specific gravity, viscosity, aniline point, sulfur, and MeABP. A summary of the major hydrogen correlations, presented in Table 8 (39), can be used to estimate the hydrogen concentration in FCC feeds and products.

Input Parameters Remarks

SG, AP, ASTM D86 Developed mainly for aviation fuels

Table 8 Hydrogen Content Correlations

Fein-Wilson-Sherman SG, AP All petroleum fuels

Dhulesia SG, RI, MW, S, Vis Developed for various FCC feedstocks

Hougen-Watson- Ragatz

K, MeABP Very inaccurate for highly aromatic, low-boiling materials

Jenkins-Walsh SG, AP Jet fuels only (A and Al )

Linden SG, AP, MeABP Calculates C/H ratio for distillate petroleum fractions

SG, MeABP Nomogram

F. Correlations of Hvdroca rbon TvDes

Better correlations of cracking yields are possible if the hydrocarbon types can be quantified. The aromatics are the most important because they have the greatest effect on coke yield and overall crackability of the feed. The more rigorous methods divide the larger groups into small subclasses. The literature describes two methods that calculate hydrocarbon types and then use these calculations to determine FCC yields.

The first method by Castiglione uses VABP, specific gravity, aniline point, and sulfur content to measure both yields and product properties (24). A correlating parameter, a, which is a measure of the aromatic content of the feed, is determined.

465

= 75 - 0.065 (VABP) - 0.9 (S) + 0.6 (AP) - 0.26 (AF'/SG)

Where: VABP = Volume average boiling point S = Sulfur,wt-% AP = Aniline point, O F

SG = Specific gravity

(7)

This method was for an FCCU running with a reactor pressure of 15 psig and 0.2 wt-% carbon on catalyst in a riser reactor operated to maximize 400°F end-point gasoline. A conversion level is assumed for the operation, and the yield breakdown follows from the correlating parameter, a. The conversion level, which is not shown, is assumed from pilot plant correlations over catalysts of varying activity. Equipment vendors usually specify the MAT activity of the catalyst to use in their guarantees. The correlations for gasoline, LPG, coke, and dry gas are shown in Figures 16a-d.

Figure 16a

Yield of C3 400' F Frnction

100, 1

Parameter:

Factor, a

C,-400" F Fraction, Vol Yo

40 50 60 70 80 90 100

Gas Oil Conversion, Vol YO

Reprinted by Permission from Hydrocarbon Processing, 62(2) 1983 All Rights Reserved

466

Figure 16b

Rutio of Gusoline in C3 4OO'F Fraction

0.9

0.8

C,-400" F1 OS7 C,-400' F EP, Vol YO Ratio 0.6

0.5

0.4 40 50 60 70 80 90 100

Correlating Factor, cx


Figure I6c

Yield of Coke und light Ends

Coke + C, and Lighter, Wt%

40 50 60 70 80 90 100

Gas Oil Conversion, Vol YO


467

Wt Ratio

0.65

figure 16d

9 0

Relationship of Coke und light Ends

Reprinted by Permission from Hydrocarbon Processing, 62(2) 1983 All Rights Resewed

A second method of using hydrocarbon types involves the calculation of aromatic (C,), naphthenic (C,), and paraffinic (C,) hydrocarbons. The N.D.M. method (ASTM D-3238-74) uses the refractive index, specific gravity, and molecular weight of the feed to calculate C, and C,. The C, value is found by difference. The test method includes equations to calculate the molecular weight and correct the density and viscosity for temperature should they be measured at something other than 20°C (68 OF).

Reif and coworkers did extensive work on correlating FCCU yields to feed properties (21). Their work, which was done over amorphous catalyst at low conversions (<60% wt-%) and cannot be used for modern FCC operations, was the first to use the hydrocarbon families along with parameters such as MABP, basic nitrogen, and total sulfur to define the role of the feedstock.

Correlations of FCCU yields against the major hydrocarbon families usually gives limited results. Workers at Akzo presented pilot plant data showing conversion, gasoline, and coke versus C, (26). Figure 17 shows the considerable scatter and indicates a number of other variables are at work. Trial and error evaluation of different regression equations is needed to find a satisfactory fit for a particular unit and is normally not directly applicable to other units.

468

Figure 17

Effect of Aromatks on Conversion, Gasone, und Coke

Conversion

0 Gasoline

40

301 20

10

0 10 12 14 16 18 20 22

c a

Source: Kreider, et. al.. Akzo Catalyst Symposium, Amsterdam, 1991

Raw data showing gasoline yield vs. conversion is plotted in Figure 18a for a commercial unit that had tested three different catalysts (41). The gasoline yields numbers were corrected for feed quality using both C, and C, / C,. Although the fit is not perfect, the data shown in Figure 18b make much more sense than that in Figure Ma, where the Agasohe/Aconversion ratio appears to be greater than 1.0 between 61 and 64 wt-% conversion.

Andreasson has plotted the hydrocarbon types against the UOP K factor (25). Although the K factor generally is thought to correlate to aromatics, his data (Figures 19a and 19b) show that C, t C, gives a much better fit for the refinery studied.

Whittington and coworkers presented the yield curves of Figure 20 for the feeds shown in Table 9 (23). Both feedstock analyses and product distribution are given for the feeds representing the three main hydrocarbon classes. The shape of the curves is important in predicting yields because overcracking of a feed has a major impact on product yields and quality. The gasoline yield reaches a maximum once the paraffins and naphthenes a te cracked.

469

52- 51 50 49 48

Gasoline 47

45 Wt %

44 43 42

41

c,-2oO4c 46

0

+ 0 0

4 + ! 0 Catalyst A

4 4 + +Catalyst B 400 0 0 0 4 Catalyst C

o+ + &+

Souroe Katalistiks Tech Service DeDartment

Figure J8b

Garsoline Selectid y Commercid FCCU

Gasoline

Wt Yo c,- 200' c

""1 Corrected for Feed Quality 49 I

48

4

o Catalyst A +Catalyst B 4 Catalyst C

45

43 Regression:

421 I # I I I I I I I 60 61 62 63 64 65 66 67 68 69 70

Conversion 200" C Wt YO

Source: Katalisbiks Tech Service Department

470

28

26

24-

22

20

18.

Feed Ca (wr Yo)

- @ @

NO Correlation

@ @

- 0 0 -

I I I

Source: Katalistiks Annual FCC Symposium, 1985

Figure 19b

Feed CA +CN (Wt%) ws UOP K Fuctor Dufu from Refinery C

11.5 11.6 11.7 11.8 11.9 12.0 12.1

Feed UOP K Factor - Source: Katalistiks WI Annual FCC Symposium, 1985

471

Figure 20 Feed Stock Quulity & Conversion Determine Totul liguid Recovery Shgle Puss Cruckhg with Zeolite Cufdyst

110 - Total Liquid

Recovery, 105- Vol Yo

(CJ

100 0 20 40 60 80 100

Conversion, Vol%

Source: Whinington. e t 4.. ACS Div. Petro. Chern. N.Y. Mtg. Aug. 29,1972

Table 9 Effect of Feedstock QluUty 011 FCC Yields

Yields: Conversion, vol-% I 70.0 I 85.0

34 230 0.13 0.04

650 810 980

93.0 0.1 2.5

34.5 73.0 5.0 2.0 4.8

Source: ACS Div. Petrol. Chem., NY Meeting, Aug. 29, 1972 (Whittington, et. al.)

P

472

(Short Resid)

IV. RESIDUAL FEED CHARACTERIZATION

Vacuum Dist. (D-1160)

Residual stocks cannot be analyzed as easily as gas oils. Specific gravities vary widely with boiling range, and regular distillations cannot be run without cracking the feed. Therefore, the K factor calculation cannot be obtained by the usual methods. The refractive index is often a key analysis in the correlations for hydrocarbon type. However, resids are opaque, and a reliable refractive index measurement of resid is not possible. To overcome these problems, researchers have obtained the residual fractions of crude oils by successive distillations (Figure 21) (42). The vacuum bottoms are separated by solvent extraction into various fractions, which are further characterized by such tests as mass spectrometry, gas chromatography, infrared spectrophotometry, and NMR.

Asphaltenes

Figure 21

n-Pentane (40 Parts Solvent) Extraction (Deasphalted Oil)

Crude Oil Andyses

Crude Oil

Usually None with Petro. Residues

011s n Resins

Reprinted by Permission from "The Concept of Asphaltenes" ACS Adv. in Chem. Series 195, Copyright 1981 American Chem. Society

473

The paraffins and naphthenes with high molecular weight should crack readily during processing even though some of the molecules are measured as Conradson carbon in the laboratory. Vaporization of the feed is important with the heavier feeds. The same basic reactions occur with the larger molecules as with their smaller counterparts, but the complex structure of the asphaltenes makes characterizing them and predicting yields difficult.

The asphaltenes are defined by the solvent used to precipitate them. Corbett and Petrossi graphically illustrate (Figure 22) the effect of the molecular weight of the solvent on the amount of precipitate for Arabian Light atmospheric residuum (28). Although many recent articles specifically discuss asphaltenes; care must be taken in relating these discussions because the type and amount of solvent used will greatly alter the asphaltene obtained. The ASTM test method D2006 calls for a minimum of 40 volumes of n-pentane. Speight and Moschopedis state that isolating asphaltenes with much lower proportions of solvent leads to errors both in the amount of precipitate and in its composition (29). When the ASTM procedure is followed, asphaltenes are found to contain 82 f 3% carbon and 8.1 f 0.7% hydrogen. This analysis corresponds to a H/C ratio of 1.18 f 0.05.

Figure 22

Effect of Solvent Curbon Number on Inso/ub/es

wt Yo lnsolubles @ 10Ol1 Ratio

10 9 8 7 6 5 4 3 2

Carbon Number of n-Paraffin Solvent

Reprinted by Permission from Ind. Eng Chem Prod. Ren Dev. 1978.17,342 Copyright 1981 American Chem. Society

474

Yen used X-ray analysis to characterize the microstructure of asphaltenes (Figure 23) (30). The micelle is made up of a number of aromatic sheets with alkyl side chains. Weak attractive forces hold several of these sheets together to form a particle of micelle. Figure 24 illustrates the approximate size of the asphaltenes (31). The macrostructure shown for asphaltenes represents a delicate balance between the oils, resins, and asphaltenes.

A number of complex structures have been proposed for various asphaltene molecules (Figures 25a and 25b) (43). However, recent work suggests that the structure of some of these heavy oil molecules (Figure 26) may be much less condensed than previously thought (44). Thus, new processing schemes may be effective in converting the small multi-ring fragments into liquid products. The size of asphaltenes are illustrated in Figure 27 along with the micelle in shale oil and normal hydrocarbon structures (38). The pore structures of catalysts will influence the processing of these molecules. Molecular sieves with openings of 9 8, or less are not likely to be effective in cracking such large structures. As the molecular weight of the asphaltene increases, aromaticity and heteroelements increase. The resins keep the asphaltenes suspended in the paraffinic portion of the crude oil. A typical analysis of residual feeds includes the following:

0 Gravity, "MI 0 Sulfur, wt-% 0 Nitrogen, wt-% 0 Sodium,ppm

Nickel, ppm 0 Vanadium, ppm 0 Simulated distillation 0 Yield, LV-% crude

Pentane, hexane, and/or heptane insolubles, wt-% Ramsbottom or Conradson carbon, wt-%

0 Iron, ppm

As with gas oils, measuring the hydrogen content of the residual feed is desirable. Carbon-hydrogen analyzers are now available, and these results should be added to the normal analyses listed previously. The value of C/H ratios for resid is somewhat limited because sulfur and nitrogen can be substituted for carbon in the hydrocarbon structures. Knowing the distribution of these heteroatoms throughout the boiling range and molecular structures is important because they affect the reactivity of the hydrocarbons and the product qualities. A more rigorous breakdown can be made by characterizing each of the fractions derived from separating the vacuum bottoms. Crude oil analysis has received a lot of attention by the major refiners (45-50). Each of these oil companies has extensive libraries on the properties of the world's crude oils to assist them in petroleum operations. The addition of state-of-the-art analytical equipment, such as low-pressure (0.02 mm Hg) distillation, high-resolution chromatography, and mass spectroscopy, provides more detailed and reliable analyses. Computers are used to extract and refine the information and put it into the most useful format.

475

The General Features Related Substances

Figure 23

of the Macrostructure of Asphaltenes and

Macro Structure of AsDhaltics A. Crystallite B. Chain Bundle C. Particle D. Micelle E. Weak Link F. Gapand Hole . G. lntercluster H. lntercluster 1. Resin J. Single Layer K. Petroporphyrin L. Metal

Figure 24

Diagrammatic Representation of an Asphaltene

Aromatic Sheets r Aliphatic 4

T 3.6 - 3.8 A (d2)

t

N, = L(d2 = 3-5

476

Figure 25u

Proposed Structure of Asphultenes

I

Venezuelun Crude Cu1;forn;un Crude

Reprinted by Permission from Advances in Chemistty Series 195 The Chemistry of Asphaltenes Pg 1 -1 5 Copyright 1981 American Chem. Society

Figure 25b

Proposed Structure of Asphulfenes CH,CH,CH, \

CH, CHCH,CH,CH,

CH,CH,CH,CH,

CH,CH,CH, CH,CH,CH,CH,CH,

CH,CH2CH,

CHJCn2vr m 1

00

CHCH,

h q u i Crude Reprinted by Permission +om Advances in Chemistry Series 195 The Chemistry of Asphaltenes Pg 1-1 5 Copyright 1981 American Chem. Society

477

Figure 26

Proposed Asphalfene Sfrucfures

CH

Previous Concept Current Concept Sourn: Malhla end Emmatein. Chem. Engr. Pmg.. Vol 80 (12) Pg 22-28 1984 Aapmduced by Penn~slan Amedcan Institute 01 Chemlcal Engineom 0 1984 AlChE

Figure 27

Pore Sizes in Shde Oils MoIecularDmeters 0 25 50 li -

Micelle 0 -lo Asphaltene

Complex Ring Structure

n-Alkane

I l l MolecularDiameters 0 25 50 1

3A

OA

Source: Tissot and Welte, "Petroleum Formation end Occurrence," Pg 270 (1978)

478

The heterogeneous atoms in the bottom of the barrel can have a profound impact on the yields from the cracking process. Each has a different effect on the FCC process. Along with sulfur, nitrogen, and the contaminant metals, Conradson carbon is also included as an atypical feed component even though it is strictly a hydrocarbon. Each of these elements is discussed in more detail.

Metals in crude oil are found predominantly in the vacuum bottoms. Porphyrins are the most commonly cited source of metals in crude oil although more than half of the metals may be in other structures. The nucleus of the chlorophyll molecule is virtually identical to the porphyrin structure except that magnesium is tetrahedrally bonded to the nitrogens in place of the vanadium or nickel. This fact leads scientists to the conclusion that plant life is one source of crude oil. The highest concentration of metals is in the asphaltenes, although the resins may contain considerable amounts as well. Nickel is, in general, more concentrated in the asphaltenes than is vanadium.

Both nickel and vanadium adversely affect the performance of most catalysts, and cracking catalysts are no exception. The value of crude oils is inversely related to their metals content. The Oil and Gas Journal continually runs assays on the world’s export crudes. These analyses were used to construct Tables 10 and 11. Similar figures were first constructed by Mauleon at Total. The best feeds (only considering metals content) are those in Table 10. Table 11 lists the higher metals content feeds. Total metals can be used to rank these feeds (Table 12). Recently, Nieskens, et al., reviewed resid processing and ranked the quality of the resids as a function of metals and carbon residue (88).

The relationship of nickel and vanadium can vary widely with different oils, but individual reserves have known compositions. Figure 28 from Tissot and Welte shows Ni vs. V for 175 crude oils from around the world (38).

V. RESIDUAL FEED CORRELATIONS

Quantifying the effect of each metal is difficult because unit performance is affected in two ways. First, the metal can reduce catalyst activity, which lowers conversion and yields. More subtly, the metals act as dehydrogenation catalysts when deposited on the FCC catalyst surface. The magnitude of these nonselective secondary reactions affects the coke selectivity of the catalyst and add to the delta coke yield. Because the delta coke directly affects the heat balance by Equation 8, the overall impact of the metals depends on the capability of the FCC unit to adjust the heat balance:

Wt-% coke (FF) = A Coke X C/O (8)

Where: A Coke = The difference in coke entering and leaving the regenerator.

C/O = Catalyst-to-oil ratio (wt/wt)

<5 PPm

5-20 ppm

LZ-i

Table 10 World's Export Crudes (650 + OF) with Low Metals Content

<5 PPm 5-20 ppm

Montrose Cinta Taching Bonny Lt. Sarir ' Statfiord Labuan Blend Tapis Bonny Med. Zanaitine

Berri Bombay High Handil Escravos Ardjuna Mubarek Miri Lite Pulai Espoir Arimbi Murban Zakum Sanga Sanga Qua Iboe Duri (Sumatra) Bekapai Argyll Rainbow Lt. & Med. Amna Minas Bekok Beatrce Rangeland - South L & M Es Sider Ekofisk Champion Brent Federated L & M

Magnus Dukan Brent Blend Maureen Qatar Marine Dulin Murchison Umm Shaif Forties Ninjan Blend Auk Fulmar Tartan Beryl

Brass River Brega Gulf Alberta L & M

I Thistle Brae

il & Gas Journal

2 M o PPm

Cabinda Zire Cormoraut South

Soyo Blend

Table 11 World’s Export Crudes (650 + OF) with High Metals Content

1 V/Ni

20-50 pprn

1 50-100 ppm I ~

I

100-200 ppm

520 ppa

Abu Al Bu Khoosh Buchan Arabian Lt. Cormorant Khursaniya Gorm Basrah Lt. Loreto Dorrood West Sak Hout Kuparuk

Ashtart Kuwait Exp. A N S

20-50 pprn

Oriente

Burgan

s0-1(10 ppa

Kole Marine Blend

Arabian Hvy. Iranian Lt. Arabian Med. Khafji Basrah Med. Kuwait Exp. Dubai S k i Eocune Soviet Exp. Bld. Foroozan Bow River Hvy. Gulf of Suez Mix

Aboozar Bahrgansar/Nowruz Basrah Hvy. Soroosh Leona

>1m PpDl

> 200 pprn

Mandji Wainwright-Kinsella

Belayim Iranian Hvy. Souedie Lloydminster

Merey Maya Tia Juana Hvy. Bachaquero

Tia Juana Lt. La Rosa Med. BCF 24

Boscan

Source: Oil & Gas Journal

Figure 28

Vanodium und Nickel Contents of 175 Crude Oils

ACanada *USA =Venezuela

Vanadium (PPm)

Shaded Area Contains 75 Crudes of Paraffinic, Naphthenic or Paraffinic-Naphthenic Types Found Around the World.

1 2 / 5 10 20 50 100 200 500

L.--- -Nickel (ppm) - W. Mrk. - Source: Tissot and Welte, "Petroleum Formation and Occurance" Springer - Verlag (1 978)

482

Table 12 Pmcessibility OF Residual Feeds

Metals Content

Typical II 50 - 75 Marginal

75 - 100 Outer limits + 100 plus Not economic

A. Contaminant Metal$

The most prominant contaminant metals in the feed are nickel, vanadium, and iron. Their effect on catalyst activity depends on the age of catalyst, the regenerator environment, the catalyst used, and the concentration of total metals on the catalyst. Research done at Ashland and confirmed by others indicates nickel has little impact on catalyst activity (up to 5,000 ppm on catalyst), but vanadium can be quite detrimental (Figure 29) (51-54). Organic iron is more like nickel in character, but metallic iron or scale acts more like a CO promoter than a dehydrogenation catalyst. For every loss of a catalyst MAT number, actual unit conversion decreases 0.3 to 0.7%.

Figure 29

6kfalysf Metal Torture Test Iron, Nickel Sodium, Vanadum Effects

Relative Activity 20

Iron

+ Nickel Sodium

100 A Vanadium

483

Dehydrogenation reactions are best monitored by the hydrogen production from the FCCU. Fresh catalysts produce 0.04 to 0.08 wt-% hydrogen in commercial operations on clean gas oils or about 20 to 50 SCF of hydrogen per barrel of fresh feed. The hydrogen produced from metal-catalyzed reactions can send this figure soaring (Figure 30). Nickel is a unique metal in that it is a strong dehydrogenation catalyst yet can be passivated effectively by compounds of antimony and bismuth or by modifications of the catalyst matrix (55,5739).

Vanadium is much more difficult to passivate because of its mobility on the catalyst surface and the difficulty in forming suitable stable vanadium compounds. Although it is not as active a dehydrogenation catalyst as nickel, vanadium concentrations on catalyst are usually twice as high. Tin has been reported to be an effective passivator as have compounds of barium, calcium, magnesium, and the rare earths (56,59). Passivation is usually considered to be effective if hydrogen is below 100 SCFB with 3,000 ppm or more vanadium.

No passivators for iron have been reported. Iron in moderate concentrations (Figure 29) does not affect catalyst activity or unit conversions. Its effect is about one-third to one-half that of vanadium on hydrogen production. Copper is a strong dehydrogenation metal, at least equal to nickel, and has no known passivators. Fortunately, the concentration of copper on the catalyst is usually very low (< 300 ppm).

The effect of contaminant metals on coke yield is conventionally determined from the coke factor of the standard MAT test (60). Coke formed as a result of the contaminant metals typically amounts to about 15 to 30% of the total coke yield in actual operations, depending on the reactor design and catalyst used. The standard ASTM MAT test is no longer a reliable indicator of contaminant coke because of the long catalyst residence time of the MAT. An estimate of the delta coke should be made from the hydrogen produced in commercial operation. This way, both operating parameters and unit design get consideration in the delta coke correlations.

B. Sodium

Sodium is not a dehydrogenation catalyst like the other contaminant metals. Its sources and effects along with the other alkali metals were discussed by Letzsch and Wallace (61). Sodium has two main effects: it reduces catalyst activity, and it can permanently destroy the catalyst structure. The former effect occurs when sodium concentrations are low. The latter effect occurs as added sodium levels exceed 0.5 wt-% on catalyst. Elevated regenerator temperature and the presenue of vanadium accelerate the structural decline attributed to sodium (51,53,62). Incorporating sodium traps can significantly improve the catalyst’s stability ‘(89). Figure 31 shows the improvement in activity maintenance possible with the use of sodium traps compared to catalysts without such technology.

Figure 30

licensee Hydrogen YieM Decrease

250

Hydrogen Yield, SCF/B FF

Unpassivated

- rn Passivated Unpassivated

200

150

100

50

50 0 2 4 6 8 10 12 14 16 18

Metals Concentration on Catalyst, 4 Ni + V, ppm in Thousands

Source: Katalistiks 3d Annual FCC Symposium (1 982) Phillips Petro.

485

MAT Activity

70 c New Technology

0.0 0.4 0.8 1.2 1.6

Sodium Concentration (Na, Wt% on Catalyst)

Source: Katalktiks' 1991 Auotrali FCC Conference

As activity declines as a result of increasing sodium, conversion and gasoline yields decline, and dry gas and diolefin yields increase. These reaction trends are symptomatic of more thermal cracking and less catalytic influence.

C. Nitrogen

Nitrogen is recognized as a temporary poison of FCC catalysts (65-69). The nitrogen compounds are strongly adsorbed on the active sites, thereby reducing the number of possible reactions. Higher reactor temperatures reduce their effect (65). Once the catalyst passes into the regenerator, the nitrogen compounds are burned and leave primarily as molecular nitrogen in the flue gas. Catalyst activity is fully restored.

Research with a wide variety of nitrogen compounds indicates a good correlation between the molecule gas-phase proton affinity and its poisoning effect (65). As nitrogen concentration increases, conversion and gasoline yields decline, the cycle oils increase, and coke selectivity worsens. The concentration of nitrogen in the liquid products declines with conversion, and the nitrogen content of the spent catalyst increases (67).

Catalyst type plays a large role in the actual effect of nitrogen (Figure 32).

486

Figure 32

Effect of feed Nitrogen on Commerciul Catulysts Performance Catalyst Pretreatment: Steam 788" C'5 Hours

Conversion, Vol Yo

Gasoline, V YO -232' C

LCO, v Yo 232 - 355' C

60

50

0.3 0.5 0.75 Wt YO Nitrogen in Feed

catalvsts

Source: Scherzer & McArthur, Katalistiks 7th Ann. FCC Catalyst Symp. 1986

487

Data on six catalysts show wide variations in nitrogen tolerance, but none of the catalysts is immune to the poison. For many gas oils, about one-third of the total nitrogen is basic nitrogen, the type thought to react with the catalyst's active sites. Commercial data indicates 1 vol-% loss in conversion for increases in total nitrogen of 100 to 300 ppm (68,70).

D. Sulfur

Sulfur is more of a nuisance than a poison. High sulfur levels in the feed directly translate into high sulfur levels in all the liquid products and the gas streams leaving the reactor and regenerator.

The actual effect of sulfur on conversion is small and probably has a deactivation mechanism that involves blockage of the active sites rather than a chemical reaction (69). Higher sulfur content in the feed usually results in less gasoline because the H,S yield increases significantly (71-73). Most of the aliphatic sulfur compounds in the FCC feed would have ended up as gasoline had they not been removed in the cracking process.

E. Conradson Carbon

Conradson carbon is a destructive distillation test (ASTM D 189) that is used to predict coke yields in delayed coking operations. A similar test is the Ramsbottom carbon (ASTM D 524). A correlation between the two tests has been prepared by the ASTM. Both residual carbon tests have been related to additive carbon. Researchers found that for a constant processing period in the lab, the coke yield equals (74):

Carbon on Catalyst = a8" t b / (C/O) (9)

Where: a, n = Constants e = Catalyst residence time b = Additive coke C/O = Catalyst-to-oil ratio (wt/wt)

Multiplying by the C/O ratio gives Equation 10:

Wt-% coke (fresh feed) = (C/O) a 8" + b (10)

The importance of the coke yield is paramount in both the design of the unit and its proper operation. Subsequently, additive carbon has been found to be a function of feed vaporization (53), the reactor-regenerator operating temperatures, and the type of feedstock. Heptane insolubles have also been used as a measure of the coke-making tendency of a feed.

488

The effect of Conradson carbon (or additive coke) is to raise the regenerator temperature and lower conversion as a consequence of lower catalyst-to-oil ratios or an increase in the coke yields in units equipped with heat removal. A hotter regenerator usually means more thermal cracking and an increase in dry gas. When Conradson carbon increases, gasoline yield declines because the hydrogen content of the feed is lower and more coke is formed.

Because a high level of Conradson carbon is associated with the vacuum resid, increasing levels of Conradson carbon in a feed is tantamount to adding vacuum resid to the gas oil. If an FCCU has spare capacity, processing at least some resid is usually economical unless the bottoms has a value approaching that of the processed crude oil.

F. Overall Effect of Resid Addition

When resid feed components are added to the cracking process, all of the above-mentioned heteroelements increase. Workers at Phillip's Petroleum quantified the relative contributions of these feed parameters to cracking yields and product quality (78). Their results indicated that of all the feed parameters examined, basic nitrogen had the largest impact on conversion and selectivities. Metals on a catalyst seem to affect the relative contribution of each of the feed parameters. Following are their equations developed for gasoline yield on the same catalyst containing 9,000 and 12,900 ppm equivalent nickel (4 Ni t V).

52.7911 + 0.5168 (API) t 1.4694 (CRES) - 0.00630 (BASN) - 0.1179 (ARO) - 0.1544 (Ni) - 0.6843 (V) (9000 PPM)

93.3637 - 0.01072 (BASN) - 0.3154 (SAT) - 0.3111 (ARO) - 0.4122 (Ni) (12,900 PPM)

Where: BASN = Basic nitrogen, ppm CRES = Carbon residue, wt-% Ni = Nickel, ppm V = Vanadium, ppm ARO = Aromatics in feed, wt-% SAT = Saturates in feed, wt-% API = Feed density, "API

The complex and changing nature of the relations shows that empirical correlations have limited value in quantifying the feed variables. Both equations were for a specific catalyst. If that variable is changed, presumably the relationships shown no longer hold. Because every resid operation uses a different catalyst, unit design, and operating parameters, the empirical correlations from other studies should be used to indicate which feed variables may be examined to explain changes in conversion and product selectivity rather than to predict specific yields.

489

VI. EFFECT OF FEEDSTOCK PROPERTIES ON PRODUCT QUALITY

A.

The effect of feed specific gravity on LPG olefinicity can be seen in Figures 33 and 34 (11). The amount of C,’s decrease with increasing density (lower “API), and they become less olefinic (Figure 33) while the total yield on C l s is about constant (Figure 34). Higher feed specific gravity reduces the branching and increases the butylenes. Overall, the volume of LPG goes down at constant conversion as the feed gets heavier. From Figure 8, the effect of lowering the K factor, or from Figure 15, the effect of decreasing the feed hydrogen (the feed becomes more aromatic) is also to lower the volume of LPG produced.

B. Gaso line Properties

The feedstock probably has the largest effect on gasoline octanes. Feedstock density was found to affect motor octanes (Figure 35) (79,80). Research octane would move in the same direction. As feeds get heavier, they become more naphthenic and then more aromatic. These molecules produce higher octanes.

The UOP K factor shows the same trends. Magee, et al., showed the variation in pilot plant octanes with K factor (Figure 36) (81). Going from 12.1 to 11.4 raised the RON and MON four and three numbers, respectively. The relatively low absolute values of the octane numbers probably accentuates the effect of the K factor.

Clear RON’S increased in both the pilot plant and commercial operation with an increasing CN/CP ratio of saturated hydrocarbons (82). Octanes usually go up as CA increases.

The ratio of feed hydrogen to carbon influences octanes and the composition of the gasoline. Figures 37a and 3% show the effect of H/C on light and heavy gasoline octanes (83). The cutpoint was 265°F. Vislocky also presents data on the feed and product compositions as a function of hydrocarbon families.

A number of other gasoline properties are a function of the fresh feed quality. Gasoline density and aromatic concentration increases with the aromatic content of the feed. Benzene concentrations in the gasoline usually increase with feed aromatics. Bromine numbers or olefinicity of the gasoline is higher for more paraffinic feeds (81).

The introduction of resid into the feed increases the concentration of aromatics. As a result, a more aromatic gasoline is produced with generally higher octanes and specific gravities. The exceptions to these general observations occur when the inclusion of resid significantly lowers unit conversion. Octanes will drop in this instance.

490

12

10

8 - C, Yields, Vol Yo

4 -

0

figure 33

-

API -

6 -

i-C, API

2- n-C, -2'

I I I I

Cutulytic Cmckiog Yields Zeolite Cutdyst [CJ

C, Yields, Vol Yo

4 i 2

01 I I I I I 40 50 60 70 80 90

Conversion, Vol YO



49 1

86

82

Octane Number 78

Figure 35

Feed Gruvity vs MONC Commerciul Dutu Anulysis

82

81- I

Gasoline Octane, MONC

79 - 0.1 5 MONCPAPI

78 I I I I I I 20.0 20.5 21.0 21.5 22.0 22.5 23.0

Feed Gravity, 'API

RON

r -

- MON

Sourca: Witoshkln, ei aL, Paper AM4646 NPRA Annual MU. 1988

Source: Magee ei al. (81)

492

Figure 37a

Averuge fight Gusoline RON und MON us u Function of feed H/C Rutio

RON

MON

Feed H/C Ratio, W W t

Source: Wslocky, ASC Symp. MV. of Petro. Chem. Sept., 1989. (83)

Figure 37b

Average Heuvy Gusoline Octunes us u Function of Feed H/C Rutio

Octane Number

Feed WC Ratio, W W t

Source: Vislocky, ASC Symp. Div. of Pebp. Chem. Sept. 1 S8S. (83)

493

c. Lieht cv cle Oil

The quality of the LCO produced in the FCC is a direct function of the aromatic content in the LCO. A correlation of cetane number to aromatics exhibits a good fit (Figure 38) (84). Although a positive correlation between feedstock and LCO densities and aromatic contents is weak, the overall conversion level for the particular feedstock is the prime consideration in the resulting LCO cetane index. Even paraffinic feeds give low cetane number LCO's at high conversions. In presenting data on a variety of feedstocks, Unzelman shows the calculated cetane numbers at 60 and 80% conversion on the FCCU (89). Other data imply cetane index increases with K factor and that cetane number and aniline point are closely related.

Figure 38

Aromutic Content ws Clew Cetune Number light Cycle Oils

Aromatics, Vol Yo

80 "j 40 30

20

' O I 0 15 20 25 30 35 40 45 50

Clear Cetane Number

Source: Collins B Unzelman API 47a M i d y e a r Meeting 5/11/62

Higher nitrogen concentraticm in the feed increase nitrogen in the LCO. These in turn can adversely affect I he distillate fuel stability, probably through reactions with mercaptans and oltbl,ns (85,86).

A portion of the aromatic feed sulfur will end up in the LCO. A rule of thumb is that the LCO sulfur content will equal the feed sulfur for many FCC units. Raising conversion will increase LCO sulfur to the point that it becomes a product liability.

494

VII. SUMMARY

Feedstock composition, which is the major variable affecting fluid catalytic cracking yields, is even more important than operating variables or catalyst selection. Any dynamic model that adjusts operating conditions to maximize unit performance is only as good as its feedstock characterization model. More important, the value of a particular crude oil is directly dependent on the yields and product qualities actually obtained although all refinery LP’s use a model to make these predictions. The recognition of these facts drives refiners and modelers to better characterize FCC feeds.

In the future, improved characterization will occur as a result of improved instrumentation and new test methods. Better predictions for residual feeds will be available, and on-line feed forward control will become more prevalent.

495

VIII. REFERENCES

1. B. S. Greensfelder and H. H. Voge, Ind Eng. Chem, 37,514 (1945). 2. B. S. Greensfelder and H. H. Voge, 983. 3. B. S. Greensfelder and H. H. Voge, 1038. 4. B. S. Greensfelder, H. H. Voge, and G. M. Good, Ind Eng. Chem, 37, 1168

(1945). 5. H. H. Voge, G. M. Good, and B. S . Greensfelder, Znd Eng. Chem., 38, 1033

(1946). 6. G. M. Good, H. H. Voge, and B. S. Greensfelder, Ind Eng. Chem., 39, 1032

(1947). 7. B. S . Greensfelder, The Chemistry of Petroleum Hydrocarbons, Vol. 2, Reinhold

Publishing, 1955, pp. 137-141. 8. P. Venuto and T. Habib, Fluid Catalytic Cracking with Zeolite Catalysts,

Marcel Dekker, 1979. 9. D. M. Nace, I n d Eng. Chem. Res. Develop., 9(2), 203-209, (1970). 10. P. J. White, Oil and Gas Journal, 112-116 (May 20, 1968). 11. J. H. Gary and G. E. Handwerk, Petroleum Refining Technology and Economics,

Marcel Dekker, 1984. 12. M. M. Boduszynski, Journal of Energy and Fuels, 1,2, (1987). 13. W. L. Nelson, oil and Gas Journal, 161 (June 11, 1962). 14. W. L. Nelson, Oil and Gas Journal, 107 (Sept. 3, 1979). 15. Technical Data Book, Petroleum Refining, American Petroleum Institute,

New York, 1966. 16. R. V. Shankland, Advances in Catalysis, 6,309-398 (1954). 17. J. B. Pohlenz, Oil and Gas Journal, 61(13), 124 (1963). 18. H. H. Voge, Catalysis, P. Emmett, ed., Reinhold Publishing, New York City,

1958, Vol. 6, Chapter 5. 19. W. S. Letzsch and A. G. Ashton, Private communication Katalistiks. 20. F. W. Whn, Petroleum Refiner, 36(2), 157 (1957). 21. H. E. Reif, R. F. Kress, and J. S. Smith, Petroleum Refiner, 237-244

(May 1961). 22. M. R. Riazi and T. E. Daubert, Oil and Gas Journal, 110-112 (Dec. 28, 1987). 23. E. L. Whittington, J. R. Murphy, and L. H. Lutz, paper at Katalistiks paper to

ACS, Div. Petro. Chem., NY meeting (Aug. 29, 1972). 24. B. P. Castiglioni, Hydrocarbon Processing, 62(2), 35-38 (Feb. 1983). 25. H. U. Andreasson and L. L. Upson, What Makes Octane, paper at Katalistiks

6" Annual FCC Symposium, May 22-23, 1985. 26. K. R. Kreider, D. A. Keyworth, and T. A. Reid, Modeling for Feed Effects in

FCC, AKZO Catalyst Symposium, Amsterdam, 1991. 27. A. G. Bridge, G. D. Gould, and J. F. Berkman, Oil and Gas Journal, 85-91

(Jan. 19, 1981). 28. L. W. Corbett and U. Petrossi, Ind Eng. Chem. Prod Res. Dev., 17, 342 (1978). 29. J. G. Speight and S. E. Moschopedis, "On the Molecular Nature of Petroleum

Asphaltenes," ACS A&. in Chem, Series 195, pp. 1-15, J. W. Bunger and N. C. Li eds., pp.1-15.

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30. T. F. Yen, "Structural Differences between Asphaltenes Isolated from Petroleum and from Coal Liquid," ACS Adv. in Chem., Series 195, J. W. Bungen and N. C. Li, eds., pp. 39-51.

31. T. F. Yen, J. G. Erdman, and S . S . Pollack, Analytical Chemistry, 1961, 33 (1987). 32. J. W. Rosenthal, S. Beret, and D. C. Green, "Hydrogen Utilization in

Residuum Conversion," paper at 48" mid-year refining meeting, API (May 10, 1983).

33. W. P. Hinds, Jr., Proceedings sh World Petroleum Congress, Vol. 4, 1971,

34. M. M. Boduszynski, Journal of Energy and Fuels, 2,597, (1988). 35. M. M. Boduszynski, "Characterization of Heavy Crude Components," paper at

36. B. Schuetz and H. Hofmann, Hydrocarbon Processing, 75-82 (Feb. 1984). 37. M. A. Bestougeff, Fundamental Aspects of Petroleum Geochemistry, B. Nazy and

38. B. P. Tissot and D. H. Welte, Petroleum Formation and Occurrence, Springer-

39. F. Valeri, "New Methods for Evaluating your FCC," paper at Katalistiks 8"

40. G. P. Hinds, Jr., "Hydrogen Conversion in Petroleum Refining," in

pp. 235-244.

Div. Petro. Chem., ACS, Chicago, Sept. 8-13, 1985.

U. Colombu, eds., Elsevier, 77-108 (1967).

Verlag, 1978.

Annual FCC Symposium, Budapest, June 1987.

J. J. McKetta, ed., Advances in Petroleum Chemistry and Refining, Vol. 10, Interscience, NY, 1965.

41. P. Gnass, Katalistiks Technical Service Report. 42. R. B. Long, "The Concept of Asphaltenes," ACS Adv. in Chem. Series 195,

pp. 17-27, J. W. Bunger and N. C. Li, eds. 43. J. G. Speight and S. E. Moschopedis, paper, Symposium on the Chemistry of

Asphaltenes (ACS), Washington, Sept. 9-14, 1979, Reprint pages 910-923. 44. J. F. Mathis and A. M. Brownstein, Chemical Engr. Prog., 80(12), 22-28 (1984). 45. C. Csoklich, B. Ebner, and R. Schenz, Oil and Gas Journal, 86 (Mar. 21, 1983). 46. R. J. O'Donnell, Oil and Gas Journal, 94 (Mar. 21, 1983). 47. F. P. McNelis, Oil and Gas Journal, 94 (Mar. 21, 1983). 48. R. J. Wampler and E. L. Kirk, Oil and Gas Journal, 98, (Mar. 21, 1983). 49. G. V. Nelson, G. R. Schierberg, and A. Sequeua, Oil and Gas Journal,

50. G. McClesky and B. L. Joffe, Oil and Gas Journal, 124 (Mar. 21, 1983). 51. W. P. Hettinger, Jr., "Development of a Reduced Crude Cracking Catalyst,"

108 (Mar. 21, 1983).

paper at ACS Symposium Fluid Catalytic Cracking: Role in Modern Refining, M. L. Occelli, ed., pp. 308-340, 1988.

NPRA Annual meeting, AM-81-44, Mar. 1981. 52. R. E. Ritter, et al., "Recent Developments in Heavy Oil Cracking,"

53, J. L. Mauleon and J. B. Sigaud, Oil and Gas Journal, 52-55 (Feb. 23, 1987). 54. P. F. Schubert and C. A. Altomare, "Effects of Ni and V in Catalysts on

Contaminant Coke and Hydrogen Yields," paper at ACS Symposium Fluid Catalytic Cracking: Role in Modern Refining, Chapter 11, 1988.

Benefits from the Phillips Metals Passivation Process," paper at Katalistiks 3d Annual FCC Symposium, Amsterdam, 1982.

55. W. C. McCarthy, T. Hutson, Jr., and J. W. Mann, "How to Estimate the

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56. A. R. English and D. C. Kowalczyk, Oil and Gas Journal, 127-128 (July 16, 1984).

57. P. Ramamoorthy, et al., "A New Metals Passivator in Fluid Catalytic Cracking," paper at the Annual NPRA meeting, AM-88-50, Mar. 1988.

58. R. C. Barlow, "Commercial Application of Vanadium Passivation Technology," paper at the NPRA Annual meeting, AM-86-57, Mar. 1986.

59. F. W. Denison, J. F. Hohnholt, and A. R. English, "Metals Passivation of Sodium and Vanadium on FCC Catalyst," paper at the NPRA Annual meeting, AM-86-51, Mar. 1986.

60. R. N. Cimbalo, R. L. Foster, and S. J. Wachtel, Oil and Gas Journal, 112-117 (May 15, 1972).

61. W. S. Letzsch and D. N. Wallace, Oil and Gas Journal, 58-68 (Nov. 29, 1982). 62. J. R. Murphy, Jr., "Designs for Heat Removal in HOC Operations," paper at the

Petroleum Refining Conference, Japan Petroleum Institute, Tokyo,

63.

64. 65.

66. 67.

68. 69.

70. 71. 72.

73.

74.

75. 76.

77.

78.

79.

Oct. 22-29, 1986. W. S. Letzsch, "New FCC Technology from Katalistiks," paper at the 1991 Katalistiks Austrialian FCC Conference, Apr. 1991. J. A. Montgomery, Davison Catalagram, No. 38, pp. 10-11. C. M. Fu and A. M. Schaffer, Ind. Eng. chem. Prod. Res. Dev., 24(1), 68-75, (1985). G. W. Young, J. Phys. G e m . , 90(20), 4894-4900 (1986). J. Schener and D. F. McArthur, "Nitrogen Resistance of FCC Catalysts," paper, at the Katalistiks Th Annual FCC Symposium, Venice, May 1986. J. D. Pollock, Katalistiks Tech Service Report, Jan. 19, 1987. R. F. Schwab and K. Baron, "Fluid Catalytic Cracking of High Metal Content Feedstocks," paper at the Katalistiks 2d Annual FCC Symposium, May 1981. NPRA Q & A transcript, "Fluid Catalytic Cracking Effect of Nitrogen." L. L. Upson and R. Sikkar, Applied Catalysis, 2, 87-105 (1982). H. C. Kliesch and T. Normand, "Verification Test Results on a FCC Model," paper at the Katalistiks 4" Annual FCC Symposium, May 1983. L. L. Upson, "Effect of Feed Quality Upon FCCU," paper at the Katalistiks 1" Annual FCC Symposium, October 1980. P. H. Johnson, C. R. Eberline, and R. V. Denton, "Catalytic Cracking of Petroleum Residuum," paper at the ACS, Division of Petroleum Chemistry, Dallas, April 1956. I. P. Fisher, Applied Catalysis, 65, 189-210 (1990). L. C. Yen, R. E. Wrench, and A. S. Ong, "Reaction Kinetic Correlation for Predicting Coke Yield in Fluid Catalytic Cracking," paper at the Katalistiks 8" Annual FCC Symposium, 1987. K. V. Krikorian and E.K. Johnson, "How to Maximize Cat Cracker Revenues," paper at the Katalistiks 8" Annual FCC Symposium, 1987. R. W. Wenig, M. G. White, and D. L. McKay, "The Effects of Feed Properties on Resid Cracking Yields," paper at the ACS, Advances in Petroleum Chemistry Meeting, Washington, DC, Aug. 28-Sept. 3, 1983. W. S. Letzsch, J. S. Magee, L. L. Upson, and F. Valeri, Oil and Gas Journal, "Advance Zeolites Used in FCC Catalysts Boost Motor Octane Number," Oct. 31, 1988.

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80. A. Witoshkin, G. S. Koermer, and R. J. Madon, "Means of Increasing Motor Octanes in the FCC Unit," paper at the AM-88-46, NPRA Annual Meeting, San Antonio, Mar. 1988.

81. J. S. Magee, R. E. Ritter, D. N. Wallace, and J. J. Blazek, "How Catalytic Cracker Feed Composition Affects Octane Catalyst Performance," paper at the NPRA Annual Meeting, Mar. 23, 1980.

82. H. F. Henz, V. M. de Marco Meniconi, and J. M. Fusco, "Petrobras Experience With Octane Enhancement in Resid Cat Cracking," paper at the Ketjen Catalyst Symposium, 1986.

83. J. M. Vislocky, "The Effect of Feed Type and Operating Conditions on FCC Gasoline Properties: A Pilot Plant Study," paper at the ACS Symposium, Div. of Petro. Chem., Sept. 10-15, 1989.

84. J. M. Collins and G. H. Unzelman, "Diesel Trends Emphasize Cetane Economics, Quality and Prediction," paper at the 4? Midyear Refining Meeting, New York City, May 11, 1982.

85. Dupont Chemical Technical Memorandum, FO-5001, Dec. 1963. 86. M. W. Schrepfer, R. 1. Arnold, and C. A. Stansky, Oil and Gas Journal, 79-84,

Jan. 16, 1984. 87. G. H. Unzelman, Oil and Gas Journal, 178-201, Nov. 14, 1983. 88. M. J. P. C. Nieskens, F. H. H. Khouw, M. J. H. Borley, and K. H. W.

Roebschlaeger, Oil and Gas Journal, 37-44, June 11, 1990.


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CHAPTER 13

SHAPE SELECTIVITY IN CATALYTIC CRACKING

FRANCIS G. DWYER AND THOMAS F. DEGNAN

Mobil Research and Development Corporation Paulsboro Research Laboratory, Paulsboro, NJ 08066-0480

1. INTRODUCTION

Shape selective catalysis is a term normally reserved to describe reactions that take place over restricted pore molecular sieves. First proposed by Weisz and Frilette in 1960 [l], the concept of zeolitic shape selective catalysis has been the basis for at least 10 commercial zeolite catalyzed processes including:

Distillate dewaxing Lube dewaxing Xylene isomerization Toluene disproportionation Et hylbenzene synthesis Methanol to gasoline conversion Paraethyltoluene synthesis

The discovery of ZSM-5 in the late 1960’s and the rapid development of this zeolite for several of the applications listed above provided the impetus for evaluating this new zeolite in Fluid Catalytic Cracking (FCC) applications. Attempts to use ZSM-5 as a single component FCC catalyst quickly showed that it could not convert the heavier components present in conventional gas oils as effectively as zeolite Y. The smaller pores of ZSM-5 (- 5.5 A vs 7.2 A ) limit access to only the linear or near linear molecules in the gas oil.

Primarily on the strength of ZSM-5’s ability to improve gasoline octane and produce light olefins for alkylation, when used in conjunction with REY and USY in catalytic cracking, i t has found widespread acceptance among refiners. Today, ZSM-5 has been used in over 20% of the commercial Thermofor Catalytic Cracking (TCC) and Fluid Catalytic Cracking (FCC) units worldwide.

The knowledge gained from evaluating ZSM-5 in diverse cracking applications will form the basis for the discussion in this chapter of the broader topic of shape selectivity in catalytic cracking. A significant portion of the chapter will examine the chemistry involved in shape selective cracking by zeolites in general and ZSM-5 in particular. This will be followed by an analysis of results from several tr ials of ZSM-5 in commercial cracking units. ZSM-5’s resistance to poisoning and attack by metals,

500

another of its unique attributes, will then be examined. Finally, laboratory cracking results with other shape selective zeolites and molecular sieves will be reviewed. Additional discussion about the application of ZSM-5 and other molecular sieve types in catalytic cracking can be found in Chapter 3 of this monograph.

2. SHAPE SELECTIVE CRACKING CHEMISTRY

Shape selectivity in zeolite catalysis is characterized by one or any combination of three primary mechanisms [2]:

(1) Reactant shape selectivity whereby molecules are sterically discriminated based upon their ability or inability to enter the restricted pores of the zeolite.

(2) Product shape selectivity whereby bulkier molecules are sterically hindered from leaving the zeolite.

(3) Spatioselectivity whereby the formation of molecular transition states is restricted by the confines of the zeolite channels, intersections, or cages.

These characteristics of shape selective catalysis have been traditionally applied to restricted pore zeolites such as ZSM-5. However, if we would extend them to Y zeolite based cracking catalysts one could argue that these also exhibit reactant shape selectivity. Indeed, many if not most of the hydrocarbon components in a typical catalytic cracking gas oil feedstock are too large in molecular diameter to diffuse through the -7.2 A pore opening of the Y zeolite that is considered the most active component in today’s cracking catalysts.

There is general agreement that the higher molecular weight components a re cracked either thermally or on the surface of the catalyst’s matrix component. The smaller molecular size products can then pass into the faujasite zeolite for further molecular weight reduction. It is a moot point whether this type of shape selectivity is responsible for the high activity and desirable selectivity associated with today’s zeolite cracking catalysts. Until active molecular sieve catalysts with pores large enough to admit gas oil feed components are discovered and developed, the validity of this speculation will not be resolved.

The octane enhancing ability of ZSM-5 in catalytic cracking is practiced in a manner different from the way it is used in other shape selective petroleum o r petrochemical processes. Normally, ZSM-5 is used as a stand alone catalyst or in sequence with other catalysts. In cracking processes, it is used as a co-catalyst. Evidence indicates that the ZSMd selectively converts components produced by the primary cracking catalyst.

Initial laboratory studies combining ZSM-5 and Y zeolite cracking catalysts showed the octane improvement benefits of ZSMJ addition that were later borne out in larger scale tests. However, it was not until the first commercial trial of ZSM-5 in the Neste Oy TCC unit [3] that the true complexity of the ZSMJ chemistry was appreciated. The surprising results from this first commercial trial formed the basis for numerous

501

controlled or semi-controlled studies that have subsequently been conducted to probe further into this chemistry.

2.1. How ZSM-5 Works The most simplistic picture of how ZSM-5 operates in a cracking regime is based

solely on its characteristic shape selectivity for cracking aliphatics (i.e., olefins and paraffins). The intermediate size pores of ZSM-5 restrict the access of highly branched and cyclic hydrocarbons to the interior of the zeolite where the active sites are located. Lower octane normal and monomethyl aliphatics enter and are preferentially cracked to lighter products. The higher octane branched paraffins and olefins and aromatics stay in the gasoline boiling range. Indeed, Chen et al. [2] have shown that the reactivity patterns of ZSM-5 favor the conversion of the lowest octane C, to C7 paraffins. For example, as shown in Figure 1 below, the relative cracking rates of normal paraffins over ZSMJ are n-heptane > n-hexane > n-pentane. However, the RON values of these paraffins follows the reverse ranking. Similarly, the paraffins with the highest degree of branching (e.g, 2,2- and 2,3-dimethylbutane, 2,3- and 2,4-dimethylpentane) have the lowest cracking rates in ZSM-5, but the highest RON values.

Figure 1. ZSM-5 Catalyzed C, to C, Paraffin Relative Cracking Rate vs Paraffin Research Octane Number [2]

Because ZSM-5 works as a co-catalyst, converting the primary products from :racking reactions over larger pore components, the reaction chemistry is not straight- Forward. As ZSM-5’s larger pore counterpart, the Y zeolite also competes with it for

502

conversion of the primary cracked products. Olefinic intermediates formed by cracking heavier hydrocarbons are among the principal molecules upon which both zeolites operate. Hydrogen transfer catalyzed by the larger pore zeolite frequently dictates the olefidparaffin ratio of the cracked intermediates and therefore their reactivity. Furthermore, ZSM-5 is much less susceptible to deactivation via coking than the Y zeolite [2]. Thus, the relative activities of the two zeolites will change as the catalysts transit the cracking zone of the unit.

Attempts to decouple the effects of the two zeolites by sequentially converting gas oil over Y-type zeolite FCC catalysts followed by gasoline conversion over ZSM-5 have often produced results quite distinct from the combined case. This interdependence of different cracking components in determining the product selectivity has led to controversy over the particular mechanism by which zeolites like ZSM-5 work. This section will summarize the areas of agreement and then consider the specific areas where interpretations differ. Table 1 is an attempt to summarize the published mechanistic studies of ZSM-5 effects in cracking heavier hydrocarbons (e.g., gas oils) or probe molecules (e.g., decane) under FCC conditions.

Several of these studies [4,5,6] have examined the effects on gas oil cracking of adding fresh, calcined ZSM-5 to USY or REY FCC base catalysts. These have conclusively shown paraffin selective cracking under FCC conditions. For example, Rajagopalan and Young [6] observed a selective 20% reduction in paraffins after adding 1 wt% of a thermally treated ZSM-5 component (704OC in air for 3 hr ) to a steamed REY catalyst in cracking a commercial gas oil. The gasoline product was enriched in olefins and C, to C, aromatics. Paraffin conversion diminished when the ZSMJ catalyst was severely steamed.

In typical FCC operations, fresh catalyst is added to the regenerator. Therefore, the catalyst always undergoes some degree of hydrothermal deactivation before contacting the hydrocarbon feedstock. For this reason the majority of studies aimed at elucidating the ZSM-5 mechanism in FCC have used samples of ZSM-5 catalyst in which the zeolite has first been placed in a conventional FCC matrix and then exposed to high temperature steam for a period of hours. The steamed catalysts have either been examined alone or physically combined with REY or USY based FCC catalysts.

Various steaming conditions have been used to simulate the equilibration of the additive under typical FCC regenerator conditions. Normally, the additives a re steamed at temperatures above 76OOC for periods ranging from 3 to 10 hours. Additive levels a r e typically 0.5 to 3 wt% ZSM-5 (as crystal) to simulate the commercial applications.

From both laboratory investigations and commercial trials there has evolved a broad consensus that ZSM-5 increases research and motor octane number by selectively upgrading low octane components in the gasoline boiling range to higher octane, lower molecular weight compounds. ZSMJ accomplishes this selectively, without increasing methane, ethane, hydrogen, or coke [71. Principal gas components a re propylene, butylenes, and isobutane with the increase in propylene being typically twice that of the butylenes. According to one investigation ethylene yields also increase marginally [S].

The resulting gasoline can be slightly more aromatic. However, this is a result of a concentration effect due to the selective cracking and removal of linear and slightly

Table 1 Summary of Mechanistic Investigations

Base Catalyst

Run. Temp. Ref. Investigators ZSM-5 Catalyst Feedstock Reactor

[71 Donnelly et al. Equil. combined REYIZSM-5 and separate particle

Thermally trt'd and steamed sep. particle

Steamed sep. particle



Equil. sep. particle

Unstearned sep. particle

Unsteamed and steamed

Unsteamed REY/ZSM-5 composite



Unit equil. sep. particle

Unsteamed sep. particle

REY Commercial gas oils

Commercial > 450OC TCC and FCC

Ragopalan and Young

Biswas and Maxwell

Pappal and Schipper

Pappal and Schipper

Schipper et al.

Anders et al.

REY and Si0,-Al,O,-clay

USY

Commercial gas oil 5oooc Fixed bed

Arab Hvy. Dist. Riser ( r e3 sec) 520OC HCpp = 0.05-0.5 bar

Fixed fluid bed USY and REY Arab Lt. Dist.

USY and REY Nigerian gas oil Riser pilot unit 538OC

REY

ZSM-5 alone and REY

USY

Commercial gas oil

Hydrotreated vacuum dist.

Model hydrocarbons and gas oil

Light fuel oil

Commercial FCC > 5oooc

Fived bed 5oooc

Miller and Hsieh

Grzechowiak and Masalska

Rawlence and Dwyer

Buchanan

Fixed bed and 0.5 BPD circ. pilot plant

516OC

Fived bed 220-33OOC [51 REY

Recirculating folded not specif. riser pilot unit

Fixed bed 538OC

USY and REY Commercial gas oil

USY and REY Pure C,-CI, olefins and paraffins

Commercial gas oil

Commercial gas oil

Madon USY Arc0 type 521OC pilot FCC unit

Davison continuous 505OC and riser 520'C cn

Elia REY

504

branched-components from the gasoline. There is no evidence for dehydrocyclization or aromatic alkylation [9,101.

There is similar agreement that steamed or equilibrated ZSM-5 does not increase boiling point conversion and that it does not change light cycle oil (LCO) and main column bottom (MCB) selectivities. Due to reactant shape selectivity, the molecules in the LCO fraction (approx C,, to Cz0) and the MCB fraction (> Czo) a re normally excluded from the pores of ZSM-5. While there may be a significant change in the distribution and concentration of olefins, there is no net increase in the level of dienes above that measured with the Y zeolite base catalyst [11,12,13].

A number of studies have shown that the increase in gasoline octane results from a decrease in the concentration of linear or singly branched C,+ paraffins and olefins and an increase in the concentration of C, and C, olefins and isoparaffins [8,9,14]. The olefins in the cracked products from the ZSM-5 containing catalysts are more highly branched than those from the base Y zeolite catalyst. This confirms that olefin isomerization is a significant reaction. Table 2 shows some typical incremental olefin yield shifts produced by adding ZSM-5 to Y zeolite cracking catalysts.

Table 2 Typical ZSM-5 Incremental Olefin Yields 171

C, Isomer Isobutene 1-butene cis-2-butene trans-2-butene

C, Isomer 2-methyl-1-butene 2-methyl-2-butene Other isomers

- C, Isomer 2-met hyl- 1 -pentene 2-methyl-2-pentene + 4-methylcyclopentene cis-3-methyl-2-pentene t rans-3-methyl-2-pen tene trans-4-methyl-2-pentene Other isomers

Yield Shift ?6 of Incremental Butene

40 20 20 20

Yield Shift % of Incremental Pentene

40 80 -20

Yield Shift % of Incremental Hexene

17 33 22 22 17 -11

505

2.2. Evidence for Shape Selectivity Olefin and paraffin cracking and isomerization reactions can be catalyzed by any

acid catalyst via conventional carbenium ion chemistry and hydrogen transfer [ 151. The evidence for zeolite shape selectivity is obtained primarily by examining the relative rates of conversion of isoparaffins to normal paraffins. Since isoparaffins have intrinsically higher cracking rate constants, evidence of selective normal paraffin conversion or higher iso-/normal ratios in the product is usually evidence for shape selectivity.

In their study of the effects of adding ZSM-5 to REY catalyst, Rajagopalan and Young [6] observed that both monomethyl and straight chain paraffins have equivalent cracking activity. They speculated that monomethylparaffins are produced at a much higher rate than n-paraffins by the REY catalyst and that, at higher temperatures, the ability of ZSM-5 to distinguish between monomethylparaffins and normal paraffins diminishes. Thus, these investigators saw little evidence for shape selective paraffin cracking.

To address the issue of shape selectivity over ZSM-5, Pappal and Schipper [16] conducted a detailed analysis of the gasoline components produced by mixtures of separate particle ZSM-5 and low and high rare earth Y catalysts. They compared these yields with the expected equilibrium values at reaction temperature and pressure. With ZSM-5 present, Pappal and Schipper observed an increase in the ratio of iso- to normal paraffins in excess of the equilibrium iso-/normal ratio for each carbon number irrespective of the rare earth level in the Y zeolite.

In the C,+ range, all of the paraffins (both branched and straight chain) decreased when ZSMS was added. With ZSM-5 combined with either the low or high rare earth Y zeolite catalysts, the normal paraffins decreased at a faster rate despite having an initial concentration (from cracking over the base case Y catalyst) that was a factor of 8 to 10 lower than that of the branched paraffins. Table 3 shows the detailed gasoline paraffin composition expressed as a percent of fresh feed.

Evidence for selective paraffin removal was obtained in a comparison of gasoline composition obtained with the two different Y zeolites. Without ZSM-5 addition, the higher rare earth, higher unit cell size catalyst produced more paraffins because of its higher hydrogen transfer activity. When ZSM-5 was added, the decrease in paraffins was greater than that observed in the analogous trial with the low rare earth Y zeolite.

Despite the consistently higher isohorma1 paraffin ratio, there was virtually no change in the overall paraffin yield with Z S M J present. However, there was a change in the carbon number distribution of the paraffins; C, and C, paraffins decreased and C, and c6 paraffins increased. The olefins showed similar increases in the iso- to normal ratio for each carbon number with ZSM-5. The yield of C,' olefins decreased while C, and C, olefins increased. The increase in the isohorma1 paraffin and olefin ratios confirmed that there is a shape selective mechanism operating. This also explained the higher research and motor octane numbers in gasoline produced with the

In a separate study, Pappal and Schipper 1171 examined the c6+ olefin isomer distribution produced by cracking gas oil over a combination of ZSM-5 and equilibrium REY catalysts. They noted that ZSM-5 produced a much larger increase in

ZSM-5.

506

singly branched olefins than it did in multi-branched olefins. This was also interpreted as a manifestation of ZSM-5's shape selectivity.

Table 3 FCC Gasoline Paraffin Composition [ 161

Catalyst A = low RE Y, low unit cell size Catalyst B = high RE Y, high unit cell size

Catalyst A A + B B + ZSM-5 z5m-5

Paraffins, wt% FF G iC, I/N

nC6

dmC6 mmC,

I/N

nC7 mmC7 dmC, I/N

nC8 mmC8

I/N dmC8

nC9 mmC9 dmC9 I/N

0.40 3.02 7.6

0.29 2.40 0.30 9.3

0.22 1.44 0.29 7.4

0.20 0.95 0.10 5.3

0.10 0.60 0.38 9.8

0.40 3.33 8.3

0.27 2.45 0.35 10.4

0.19 1.48 0.31 9.4

0.18 0.90 0.12 5.7

0.09 0.55 0.30 9.4

0.29 1.76 6.1

0.16 0.90 0.15 6.6

0.17 0.70 0.18 5.2

0.10 0.60 0.08 6.8

0.09 0.35 0.38 8.1

0.29 1.95 6.7

0.14 1 .oo 0.17 8.4

0.13 0.60 0.20 6.2

0.08 0.50 0.07 7.1

0.07 0.25 0.35 8.6

2.3. Paraffin Conversion There is much disagreement on the subject of gasoline range paraffin cracking as

catalyzed by ZSM-5 in the FCC unit. Miller and Hsieh [18] for example spiked a gas oil feed with normal nonane and saw less than 5% conversion of the normal paraffin in cracking studies (516OC) over an equilibrium Y FCC catalyst blended with 10% ZSM-5

507

as an additive. From this they concluded that paraffin cracking plays a negligible role in octane enhancement.

Rajagopalan and Young 161, examined the products generated by passing an REY cracked gasoline over a steamed 1% ZSM-5 catalyst at 500°C and determined that it had negligible paraffin conversion activity. However, when ZSM-5 was combined with an REY catalyst, the paraffin yield as percent of fresh feed decreased markedly when compared to the products from the REY catalyst alone. The largest decrease was in the C, and C, paraffins. They noted that at equivalent boiling point conversion, the thermally treated (unsteamed) ZSM-5 catalyst reduced gasoline paraffin content twice as much as did the steamed catalyst.

From their data, Rajagopalan and Young postulated that the reduction in gasoline range paraffins was not a result of cracking catalyzed by ZSM-5. They believed that ZSM-5 was preventing the formation of paraffins by impeding secondary reactions (e.g., hydrogen transfer and chain transfer) which led to the formation of paraffins from olefins. Rajagopalan and Young suggested that ZSM-5 prevented these bimolecular reactions by catalyzing the monomolecular cracking of larger carbenium ions produced by the Y zeolite to C, and C, products.

Buchanan 1101 examined the cracking characteristics of pure component normal paraffins and olefins over severely steamed ZSM-5 alone and ZSM-5 admixed with REY and USY. Reaction temperatures and residence times were selected to approximate typical FCC conditions (53SoC, 7 = 6 sec). The paraffins exhibited little reactivity; conversions were less than 1%. Cracking of the olefins ranged from 26% for hexene to 80% for decene.

Buchanan noted that olefin isomerization was faster than olefin cracking which, in turn, was much faster than paraffin cracking. He concluded that equilibrated ZSM-5 in the FCC unit does little paraffin cracking and that the reduction in paraffins resulting from ZSM-5 addition to the FCC is due to the cracking of gasoline-range olefins (primarily C,'). This leaves fewer of these olefins available for conversion to paraffins via hydrogen transfer over the Y zeolite base cracking catalyst, Hexene cracking experiments over sequential beds of ZSM-5 and Y zeolite base catalysts demonstrated that olefin cracking by ZSM-5 can reduce the formation of paraffins over REY and USY catalysts. Placing ZSM-5 behind the base Y zeolite catalysts had little effect on hexane yields, indicating that little cracking of hexane was accomplished by

Madon reached similar conclusions in his pilot scale riser studies [8]. By comparing the product distributions from cracking vacuum gas oil with USY and REY versus combinations of severely steamed Z S M J and either USY or REY, he concluded that the decrease in paraffin yield is due to the removal of olefins which would otherwise undergo secondary hydrogen addition.

Biswas and Maxwell [91 noted a significant loss in C,' paraffins when ZSM-5 was added and attributed this to selective paraffin cracking by ZSMS. They also observed that paraffin and normal olefin yields expressed as percent of fresh feed, diminished at the same rate when plotted against ZSM-5 content. This suggested a correlation between the normal olefin and paraffin concentrations.

ZSM-5.

508

Cracking studies by Anders [4] conclusively showed selective paraffin cracking over fresh, unsteamed ZSM-5 catalyst under simulated FCC conditions.

The explanation for these diverse observations may lie in the acidity of the zeolite used in the study. As Rajagopalan and Young have shown, ZSM-5 catalyzed paraffin conversion activity diminishes markedly when the zeolite is steamed. Whether selective paraffin cracking is noted or not appears to depend on the degree of hydrothermal deactivation (i.e., dealumination) of the ZSM-5. Because they are several orders of magnitude more reactive, olefins can still undergo extensive isomerization and cracking over fewer and/or weaker acid sites.

2.4. Effect of ZSM-5 Additive Level Biswas and Maxwell [9], examined the effects of adding ZSMJ as a separate particle

to a U S Y catalyst over a range of ZSM-5 (crystal) loadings ranging from 0 to 3 wt%. They showed that gasoline yield decreases in direct proportion to the amount of ZSMJ that is added. Conversely, they showed that butylene and propylene increase at the same rate in proportion to the amount of ZSM-5 that is added.

In another study which examined the effects of ZSM-5 used in combination with USY, Madon [S] noted that doubling the ZSM-5 content (from 1.1 to 2.2 wt%) doubled the RON and MON increases but did not double the gasoline yield loss. He attributed this to changes in the reactivity of the gasoline fraction with compositional changes produced by ZSM-5.

Elia et al. [19] varied the added Z S M J level from 0 to 3 wt% in their study of the effects of cracking temperature in a pilot scale riser reactor. They found that RON gains above the equilibrium RE-USY base cases were relatively insensitive to the cracking temperature in the 500 to 52OoC range. However, over this same range, the response of MON to ZSM-5 addition decreased as the temperature increased. In an attempt to explain this, they carried out an extensive compositional (i.e., PIONA) analysis. Their analysis showed that, at the lower temperatures, there was an increase in light, predominantly branched, olefinic and paraffinic compounds in the C5-C6 range. However, there was almost no increase in aromatics or naphthenes. At the higher temperatures, they observed an increase in C7-CS aromatics and naphthenes and a much smaller increase in C5 and C6 compounds. Because MON is very sensitive to the concentration of light branched aliphatics, they reasoned that this was the reason for the greater MON sensitivity at the lower temperatures.

In FCC applications, ZSM-5 content has ranged up to 3 wt% (expressed as wt% ZSM-5 crystal) of the total inventory. The targeted ZSM-5 loading is dictated not only by the desired octane increase, or light olefin production levels, but also by the base octane, gasoline cut point, regenerator temperature and base catalyst makeup rate.

2.5. Effect of Zeolite SiO,/AI,O, Ratio Virtually all of the published studies listed in Table 1 were based on catalysts

prepared with ZSMS catalysts with SiO,/AI,O, ratios less than 100, the range in which ZSM-5 is typically prepared. Varying the zeolite SiO,/AI,O, ratio changes the number of acid sites but has little, if any, effect on the pore dimensions. By examining the

509

performance of a series of zeolites with different SiO2/AI20, ratios, the effects of acid site density and concentration can be elicited.

In a series of gas oil cracking studies with a separate particle ZSM-5 and equilibrium FCC catalysts, Miller and Hsieh [18] examined the effects of varying ZSM-5 Si02/A1203 ratios over the range of 40 to 1000. They found that the Si02/A1203 ratio of the ZSM-5 crystal had a significant effect on gasoline yield loss, but had much less effect on gasoline octane and little or no effect on C4 olefinicity.

The gasoline yield obtained with a 525 SiO2/AI2O3 ratio ZSM-5 was still lower than the base cracking catalyst, but was higher than with a ZSM-5 catalyst with a Si02/A1203 ratio of 40. The octane gain/gasoline yield loss ratio was 0.6 for the additive catalyst prepared with the lower (40:l) SiO2/AI2O3 ZSM-5 vs 1.0 for the catalyst prepared with the higher (525: 1) Si02/A1203 ZSM-5.

Analysis of the reaction products showed that there were definite shifts in selectivity. With the lower Si02/A120, ratio ZSM-5, C, olefins and paraffins were increased primarily at the expense of C, to C, olefins and paraffins. Addition of the higher Si02/A1203 ratio ZSM-5 did not change the C, to C, distribution but reduced the C, fraction.

The ZSM-5 additives increased the i sohorma1 paraffin ratio of the C4 to C, fractions irrespective of the Si02/A120, ratio. Both additives also produced more C3 to C, olefins and higher iso-/normal olefin ratios than the base cracking catalyst. Most notably, there was a significant reduction in the amount of linear a-olefins

Miller and Hsieh [18] concluded that the octane increase with the lower Si02/A1203 additive primarily came from the enrichment in C, olefins. With the higher Si02/A1203 ratio ZSM-5, at least half of the octane increase came from the higher iso-/normal paraffin ratio of the gasoline, even though overall gasoline olefin content increased. Within the gasoline fraction, both ZSM-5 additives selectively reduced the yield of the heavier gasoline fraction (130 - 22OOC). Miller and Hsieh proposed that this yield loss resulted principally from paraffin conversion as denoted by the increase in the aromatics to paraffin ratio throughout the gasoline boiling range.

2.6. Summarv - Octane Enhancement Chemistry

These include: Shape selective zeolite catalysts increase FCC gasoline octane by several means.

Selective cracking of the lower octane linear and monomethylparaffins and olefins.

Decrease in the overall molecular weight of the gasoline.

Net decrease in the amount of low octane C, olefins and paraffins and a corresponding increase in the amount of high octane C, components.

Concentration of the aromatics.

Increase in the iso-/normal paraffin ratio.

Increase in the iso-/normal olefin ratio.

510

(7)

Taken together, the data from the studies listed in Table 1 convincingly show that ZSM-5's shape selective cracking, isomerization, and reduced hydrogen transfer characteristics all contribute to increasing gasoline RON and MON. However, the studies also suggest that the relative contributions of each of these catalytic characteristics does change over the lifetime of the zeolite in the cracking unit. It would appear that ZSM-5 works in three sequential mechanistic regimes as shown in Figure 2 below.

Increase in the overall olefin content.

Dominant Reactions Catalyzed by ZSM-5 FCC Additives

Fresh

1 ZSM-5 Catalyst Age

Olefin

Paraffin Cracking

Olefin Cracking

I

Deactivated I

Increasing Gasoline Yield Loss - Increasing Octane + C3= + C4= Gain -

Figure 2. Dominant Reactions Catalyzed by ZSM-5 FCC Additives

In its fresh state, soon after its addition to the FCC regenerator, ZSM-5 has sufficient acidity to crack C,' paraffins, isomerize and crack C,+ olefins, and catalyze some hydrogen transfer reactions to create isoparaffins from the isoolefins. Gasoline yield loss is significant because of the cracking of paraffins and olefins via conventional carbenium ion chemistry to C, and C, components. Octane is increased as a result of the removal of linear or near linear paraffins as well as olefin isomerization and aromatics concentration.

After many regenerations, or following severe steaming to simulate an equilibrium catalyst, the paraffin cracking activity of ZSM-5 diminishes due to the reduction in

511

zeolite acid sites. The zeolite still has sufficient activity to catalyze olefin cracking since olefin cracking rates are two to three orders of magnitude faster than paraffin cracking rates under FCC conditions [10,20]. Gasoline yield loss is still significant because olefin cracking predominates. Octane is increased as a result of the increase in the concentration of isoolefins, particularly the lighter olefins such as isopentene.

After ZSM-5 has been through several hundred regeneration and cracking cycles, perhaps after several weeks of being exposed to regenerator temperatures above 7OOOC in 6 to 8 psia of steam, the zeolite loses a significant fraction of its acid activity due to dealumination. Its framework SiO,/AI,O, ratio increases significantly and it exhibits the catalytic characteristics of the higher SiO,/AI,O, ZSM-5 described by Miller and Hsieh [MI. Olefin cracking activity diminishes with a concomitant drop in C, and C4 production.

Despite being highly dealuminated, the ZSM-5 catalyst still remains active for olefin isomerimtion (primarily from normal to monomethylolefins) since the rate constant for this reaction over acid sites is again several orders of magnitude greater than that of olefin cracking [10,20]. Because the density and the number of acid sites are reduced, the isoolefin has a greater chance of diffusing within or out of the ZSM-5 crystal without undergoing cracking. The isoolefin can therefore undergo bimolecular hydrogen transfer over the Y zeolite base catalyst to produce the corresponding isoparaffin.

3. COMMERCIAL EXPERIENCE

In analyzing the impact of new FCC catalysts, commercial selectivity and activity data are the most significant, but also the most difficult to obtain. Day to day changes in feedstock properties and operating parameters in response to market demands often preclude a precise comparison with the performance of previous catalysts. Moreover, in typical FCC operation, the changeout of one catalyst for another is usually accomplished over the course of many weeks while the unit continues to operate. For competitive reasons, refiners usually wish to limit dissemination of information regarding the base catalysts they a re using, their operating conditions, and the properties of the feedstocks they are running.

For all of these reasons, the number of detailed published studies of commercial FCC catalyst performance is usually small. Such has been the case with ZSM-5. The five published commercial evaluations of the effects of adding ZSM-5 to either FCC or TCC units are summarized in Table 4.

In each of these cases, the ZSM-5 catalyst was added incrementally and allowed to build up to a value of no more than approximately 3 wt% (expressed as wt% ZSM-5 crystal) of the catalyst inventory.

3.1. NesteOy The first full scale commercial trial of ZSM-5 in cracking occurred in the TCC unit

at the Neste Oy refinery in Naantali, Finland in 1983 [3]. The catalyst was a composite bead catalyst containing both REY and ZSM-5. The ZSM-5 containing catalyst was

Table 4 Summary of Commercial Trials

Gasoline Loss Octane Gain Feedstock or (Gain) vol% ARON AMON Notable Features

Neste Oy, Gas Oil Naantali, Fin. (1983) [Ref. 31

2 4

Oklahoma Refining Paraffinic VGO 2 3 1.7 Cyril, OK

[Ref. 221

ENICHEM, ANIC Vacuum Gas Oil 0.9 1.9 Gela, Italy (1985) Atm Resid [Ref. 11,12,131

(1984)

HDT Hvy Coker GO

Indian Oil Corp. Heavy Vacuum GO (4.9) 2.4 India (paraffinic) (gain) [Ref. 231

2 Maximize octane . Composite REYIZSM-5 Catalyst - RON/vol loss improved with TOS Selectivation demonstrated . TCC Unit

0.6 . Maximize octane and alkylate Composite REYIZSM-5 catalyst Target octane obtained 1-7DOS First FCC trial

0.7 . Maximize octane Separate part. ZSMJ additive . Metals tolerance demonstrated . Trials with REY and USY Two addition strategies tried

not . Maximize LPG - extend LCO cut pt.

. Trials in two FCC units - Low severity (35% conv.)

. Increase in gasoline volume . Distillate pour pt. reduced

reported - Separate part. ZSMJ additive

Unocal Corp. Atm. Gas Oil 1.4 (R+M)/2 = 0.7 - Maximize alkylate - inc. octane Coker Gas Oil . Separate part. ZSMd additives Lemont, IL, USA Vacuum Gas Oil

[Ref. 241 Light Distillates Evaluation of different additives (1987-1989)

513

added as make-up over a 92 day period averaging slightly below 2 tons per day. The TCC unit had a fresh feed rate of 13,400 BPD and a 347 ton catalyst inventory. Catalyst circulation rate was -400 tons per hour. Reactor vapor outlet temperature was 484OC. In the publication no description of the feedstock properties or catalyst composition was provided.

The gasoline showed a steady increase in both RON and MON and a corresponding decrease in yield. As expected, the principal cracked products were C3 and C4 olefins. At the end of 92 days of ZSM-S/REY addition, the product analysis showed a RON gain of 4.5, a MON gain of 2.1, a gasoline loss of 2.2 vol%, and a C3 + C4 olefin increase of 1.9 ~01%. Figure 3 shows the performance of the unit as a function of time over the 92 day addition period as well as for the subsequent 210 days.

0 Reactor Effluent Survey 0 Refinery Monitoring Q End of Catalyst Addition Gasoline, 42

% Vol 4,

40 - D b - 4

i T - 4 -

I l l 1 1 I I I

- = 6 - - - Motor Octane 78

77

0 20 40 60 80 100 120 140 160 180 200 220 240 260 280 300 Time on Stream, Days

Figure 3. Neste Oy ZSM-5 Additive Trial - Commercial Data 131

Based on earlier, shorter duration laboratory evaluations it was expected, as the catalyst deactivated in the unit, that the unit performance would rapidly return to its pre-test state. Surprisingly, the gasoline yield gradually rose to the pre-test level, but

514

90 89 88 87

85 84 83 82 81 80

Research 86 Octane

the MON and RON both showed significant residual improvement above pre-test levels. These observations were confirmed with laboratory evaluations of catalyst samples taken 14 and 133 days after the end of the catalyst addition. Laboratory fixed bed catalytic cracking results obtained using samples of the Neste Oy feed taken at the start of the test run and catalyst samples taken from the unit at various times are shown in Figure 4.

- - -

p - - k - -

- -

1 1 1 1 1 1 1 1 1 1 1 1 1

43 -

0 Initial TCC Feed A Corresponding Time

TCC Feed

Figure 4. Laboratory Evaluation of Neste Oy Catalysts [3]

To explain the residual octane enhancement, Anderson et al. proposed that ZSM-5 could be increasing octane by two separate mechanisms. The first they interpreted to be the cracking of low octane normal olefins and paraffins in the gasoline fraction to lighter hydrocarbons. The second was the isomerization of straight chain olefins to higher octane branched olefins. They proposed that as the catalyst aged in the unit, the higher acid activity required for paraffin and olefin cracking was destroyed, but the lower acid activity remained. This was sufficient for olefin isomerization.

515

t2 Vol% + 1

Base

The Neste Oy study was significant in that it demonstrated a “residual” gasoline octane improvement to gasoline volume loss ratio that was much larger than observed in earlier laboratory studies.

The sustained residual octane improvement noted following the end of ZSM-5 addition was later demonstrated by Takano in his trials of a commercial FCC ZSM-5 additive [21]. The response of C, and C, olefin production, RON and gasoline yield, to starting and stopping ZSM-5 addition in Takano’s study are shown in Figure 5.

Cg -

.

Addition Start Addition Stop

+ 2 Vol% + 1

Base

CZ - - -

+ 2

+ 1

Base -1

Vol% -2 -3

-

-==- RON -

1 -

One Week M

- Gasoline .I= - I

Time on Stream

Figure 5. Effects of Starting and Stopping ZSM-5 Addition on C, and C, Olefin Production, RON, and Gasoline Yield [21]

3.2. Oklahoma Refining Co. The second full scale commercial demonstration of the use of ZSM-5 to increase

octane in catalytic cracking took place in 1984 in a fluidized catalytic cracking unit of the Oklahoma Refining Co, in Cyril, Oklahoma [221. The test was conducted by the Filtrol Corporation. A ZSM-5 component was introduced as a fluid additive containing REY. No information was provided regarding the ZSM-5 component level in the catalyst or the catalyst addition rates. The feedstock was highly paraffinic (API gravity = 28.4), and the refinery had excess alkylation capacity and available isobutane.

The goals of this test were to quantify ZSM-5 catalyst performance in an FCC unit where the operating severity and base octane of the cracked gasoline were higher than

516

in the Neste Oy TCC operation. Catalyst inventory and catalyst turnover were also considerably higher in this unit. This trial attempted to determine to what extent the results of the Neste Oy test would translate to a typical FCC operation.

The Oklahoma Refining Co. FCC study demonstrated three important effects 1221:

The degree of paraffin cracking was directly related to the amount of ZSMJ in the catalyst.

. Octane benefits were realized even when a paraffinic stock was cracked at temperatures as low as 493OC.

Total gasoline plus alkylate yields increased above levels obtained without ZSM-5.

The commercial data showed that adding ZSM-5 resulted in a gasoline loss of 2.3 vol% and produced 2.4 vol% more C, and C, without significantly increasing light-gas yields or coke yield. Because the refinery had excess alkylation capacity and isobutane available, C, and C, olefins were alkylated. Total gasoline plus alkylate increased by 1.1 ~01%. Gasoline RON increased by 1.7 and MON increased by 0.6 numbers.

No attempt was made to monitor the performance after the addition of the ZSM-5 containing catalyst was ended. With the more rapid turnover of the catalyst in the FCC unit, the residual octane enhancement noted in the Neste Oy unit most probably would not have been as prolonged.

3.3. ENICHEM ANIC ZSMS was added as a separate particle catalyst for the first time in a commercial

unit at ENICHEM ANIC’s Gela Refinery FCC unit in 1984 [11,12,13]. This is a UOP side-by-side unit with approximately 120 metric tons of catalyst in inventory. Two trials with ZSM-5 were conducted; one with an REY base cracking catalyst and one with a USY base cracking catalyst. The feedstock for ENICHEM ANIC’s unit during these trials was either a mixture of vacuum gas oil and hydrotreated heavy coker gas oil or a blend of these two gas oils with an atmospheric resid. Concentrations of the high Ni and V resid ranged up to 50 vol% of fresh feed.

Prior to the introduction of ZSM-5, the unit was operating with a conventional REY catalyst with a make-up of 1.6 to 1.7 metric tons per day. Two different ZSM-5 strategies were tried. In the first, the separate particle ZSM-5 additive was blended with fresh REY catalyst and introduced at the normal make up rate. This afforded a relatively modest build-up in ZSM-5 catalyst inventory and ensured that the resulting increases in C, and C, would not overload the wet gas compressor.

In the second strategy, an additive with a high concentration of ZSM-5 was introduced into the unit at nearly twice the rate as the first strategy. The addition rate was then reduced to maintain the desired octane level.

In the first trial RON and MON increased by 1.9 and 0.7 numbers respectively. Consistent with laboratory studies, there were no changes in coke level, decant oil, H, or C2- make. Significant increases in propylene and butene were noted with -1.5 wt%

517

ZSM-5 in the inventory. The changes in the C3 and C4 components which resulted from the addition of ZSM-5 are shown in Table 5.

Table 5 ZSMS Performance in ENICHEM ANIC’s Gela FCC Unit [ll] Changes in the composition of C3 and C4 Components

Base Case + ZSM-5

Component, vol% Propane Propylene Iso bu t ane n-Butane 1-Butene + Isobutene 2-Trans-Butene 2-Cis-Butene Butadiene

14.0 35.4 17.2 5.0

15.7 7.6 5.0 0.2

A -1.4 +1.5 +o. 1 -0.3 +0.4 +0.3 +0.4 nil

-

Analysis of the gasoline composition showed a net decrease in paraffins and an increase in olefins, Table 6.

Table 6 ZSMJ ENICHEM ANIC’s Gela FCC Unit Test [ 111 FCC Gasoline Analysis

______ ~~~ ~~

Base Case + ZSM-5

Composition, wt% Paraffins Oletins Naphthenes Aromatics

Physical Final Boiling Pt, OC Specific Gravity

RON MON

22.0 31.0 20.0 27.0

165 0.730

92.6 80.1

21.0 33.0 20.0 26.0

165 0.730 A

+1.9 +0.7

-

518

0 0 Feedstock A (92.6 Base Octane) 0 W Feedstock B (90.9 Base Octane)

-

-

W

Up to a t least the 2.5 wt% ZSM-5 level, MON and RON increased in direct proportion to the amount of ZSM-5 in the unit inventory. This is shown in Figure 6. Here, the rate of ZSM-5 addition was constant.

A Octane (Clear)

2

1.6

1.2

0.8

0.4

0 ~

0 1 2 3 4 S 6 7 8 9 1 0 1 1 1 2 1 3

Time on Stream, Days

Figure 6. Octane Increase vs Days-On-Stream for the ENICHEM ANIC Trial [ll]

3.4. Indian Oil Corporation, Ltd. Das et al. 1231 have recently reported the results of commercial trials of ZSM-5 in

two different FCC units operated by the Indian Oil Corporation. The objective of adding ZSM-5 in this trial was to increase the amount of LPG produced in a low severity (30 to 45 vol% conversion to 216OC3 operating mode. LPG is valued highly in India. However, the distillate mode of operation of the majority of Indian FCC units does not permit cracking at higher temperatures to increase LPG.

The feedstock was a heavy vacuum gas oil (pour point = 4SoC, API gravity = 27.3) derived from a paraffinic crude (Bombay High). In both units, ZSM-5 was added as a separate additive to an REY base catalyst. No information about the design, capacity, or operation of commercial units was provided except that one operated a t a higher severity than the other.

Unexpectedly, Das et al. found that the yield of gasoline (FBP = 150OC) actually increased at equivalent catalyst/oil ratios when ZSM-5 was added. Yield improvements were as high as 3 vol% at 35% conversion. They attributed this to a selective shift in the boiling range out of the 216' - 371OC range and into the light gasoline range due to the low conversion operation. They also saw a net increase in gasoline octane, and C,

519

and C4 yields, with a significant gain in propylene make. Dry gas and coke make were both reduced.

Most significantly, from a shape selectivity standpoint, Das et al. observed a drop of as much as 5OC in the pour point of the distillate after ZSM-5 was added. This suggested that ZSM-5 was selectively cracking the normal or slightly branched paraffins as it does in Mobil’s distillate dewaxing (MDDW) process [2]. The Indian Oil Company was able to take advantage of this by increasing the end point of its distillate to increase distillate yield.

3.5. Unocal Kowalczyk et al. [24] have described the commercial performance with ZSM-5

additives in a 58,000 BPD FCC unit at Unocal’s Lemont refinery over an 18 month period. The shape selective additive was used in the UOP side-by-side unit to produce additional olefins for alkylate production. The feedstock comprised atmospheric gas oils, vacuum gas oils, coker gas oils, and lighter distillates. The trial was discontinued because of a major revamp in the FCC unit and the reduction in excess alkylation capacity.

Over the course of the trial, Unocal obtained an average increase of 0.7 road octane (R+M)/2 with a gasoline loss of 1.4 ~01%. The ZSM-5 additive level was 4.8 wt% of unit inventory (1.2 wt% ZSM-5 crystal of a 25% additive). Including alkylation of the light olefins, the estimated net increase in gasoline yield was approximately 2 ~01%. There was no discernible effect on coke yield and no increase in C2- products. Upon discontinuing ZSM-5 addition to the unit, olefin yields returned to their previous values after about five days.

3.6. Experience in Other FCC and TCC Units In addition to the trials described above, at least 60 FCC and TCC units, ranging in

size from 6,000 to 90,000 BPD have used ZSM-5 additives. Table 7, taken from a paper by Schipper et al. [14], provides a summary of the RON and MON data from twenty of these units. ZSM-5 content in the units ranged from a minimum of 0.2 wt% to a maximum of 3.0 wt% with an average of 2.0 wt%.

Table 7 Summary of Commercial Performance of ZSM-5 in Catalytic Cracking [14]

A 91.5 +1.5 79.0 +0.6 2.8 B 93.4 +0.9 80.5 +0.3 2.0 C 91.8 +1.0 79.8 +0.3 0.2 D 86.0 +4.5 77.4 +2.2 2.2 E 87.3 +1.7 78.2 +0.6 0.9

520

Table 7 (continued) Summary of Commercial Performance of ZSM-5 in Catalytic Cracking [14]

F G H I J K L M N 0 P Q R S T

87.6 89.8 88.5 91.5 92.6 93.4 92.5 92.5 92.5 92.0 92.5 92.7 91.4 88.6 91.0

+1.6 +2.2 +1.5 +1.2 +1.9 +0.7 +0.9 +1.4 +1.0 +1.0 +0.7 +1.2 +0.8 +1.6 +1.2

77.3 N/A N/A 81.5 80.3 80.3 80.5 80.6 81.5 79.9 80.2 80.3 N/A N/A 79.3

+1.2 N/A N/A +o.s +0.7 +0.3 +1.1 +0.9 +0.4 +0.4 +0.5 +0.5 N/A N/A +0.7

0.3 1.2 0.8 0.5 1.5 2.4 0.2 2.2 2.4 0.2 0.2 3.0 1.5 2.2 2.2

In all twenty of these cases, addition of the shape selective zeolite increased both RON and MON and produced C, and C, olefins at the expense of gasoline. Variations in the improvement in RON and MON for a given concentration of ZSM-5 are due to variations in the base octane of the gasoline (i.e, octane without ZSMJ), base catalyst makeup rate, gasoline cut point, and regenerator temperature [14].

Experience has shown that the higher the concentration of C,' paraffins and olefins in the base gasoline, the lower will be the base Research and Motor octane. With a fair degree of accuracy, commercial experience has shown that the base octane can be used to characterize the gasoline upgrading potential. When both the Research and Motor octanes are lower, the gasoline is more amenable to being reacted over ZSM-5, and less ZSM-5 is needed to obtain a given octane increase [14]. This is shown in Figure 7.

52 1

A RON

Relative Yield Shift

2.5 - 2 -

0 .25 .5 .75 1

Relative ZSM-5 Activity

FCC Gasoline Plus Alkylate

i

I I I

Base Octane 89 91 93

Figure 7. Impact of Base Octane on ZSMS Catalyzed Research Octane and Gasoline + Alkylate Improvement [14]

Commercial and laboratory data have also shown that ZSM-5 increases the octane across the entire gasoline boiling range. The composition of the base gasoline and the cut point of the product gasoline define the overall octane improvement. This is illustrated by an example taken from Schipper [14] which uses gasoline fractions taken from a commercial trial (Figure 8).

522

84

82

80

78

76

74

72

A RON = + 1.5 104

100 - Base + ZSN-5

+2.5 RON

+ 1.5 RON

+ 1.0 RON

96

92 - 88

84

80

- Research Octane

- -

Base +' ZSM-5 A MON = + 1.0 -

-

- + 1.0 MON -

- +2.5 MON +0.6 MON

I I I I I I I I I

Motor Octane

Figure 8. Commercial Example of RON and MON vs Gasoline Cut Point [I41

Addition of ZSM-5 improved the octane of the 12O'OC gasoline fraction by 2 to 2.5 MON and 2 to 3 RON and the 12O-OC fraction was improved by 0.5 to 1.0 MON and 1 to 2.5 RON. The overall octane improvement is 1.0 MON and 1.5 RON. Consistent with the explanation given above, the octane of the heavier gasoline fraction increases because the concentration of low octane C,' paraffins and olefins decreases. The lighter gasoline fraction has a higher octane due to the concentration of monomethyl C, and C, olefins and paraffins.

This experience has enabled refiners to obtain the necessary data to determine how to best integrate the use of ZSM-5 catalyst in cracking into their overall refining operation.

4. ZSM-5 RESISTANCE TO METALS AND POISONS

Nickel and vanadium are well known for their adverse effects on FCC catalysts. Nickel catalyzes dehydrogenation reactions in the riser thereby increasing coke and hydrogen production. Vanadium is known to reduce the crystallinity of faujasite. Vanadic acid and low melting point vanadium salts formed in the regenerator, attack

523

the zeolite [25]. Sodium is known to increase the potency of vanadium through the formation of sodium vanadate.

Both laboratory and commercial trials have demonstrated that the effects of nickel and vanadium are not as pronounced on ZSM-5 as they are on USY and REY [7]. As mentioned previously, the shape selective additive has been used commercially at ENICHEM ANIC to process feeds containing up to 50% atmospheric resid [11,12,13,14]. No adverse effect on the additive’s yield-octane response was noted even at catalyst metals levels up to 10,000 ppm nickel plus vanadium and 6,000 ppm sodium.

The effects of nickel on ZSM-5 and REY are compared in Table 8. Here, the catalysts were doped with nickel naphthenate then calcined at high temperatures to simulate regenerator conditions. Nickel has a negligible effect on the ZSM-5. However, hydrogen, coke, and light gas all increased significantly when the nickel has been deposited on the REY catalyst. Octane increased as a result of the formation of aromatics and olefins via dehydrogenation over nickel.

Table 8 Effect of Nickel Addition on ZSM-5 Product Selectivity [7]

Case 1 Case 2 Case 3

Nickel, ppm On Base REY On ZSMJ Additive

A Yields, % FF at constant conversion relative to base REY without ZSM-5 or nickel

Gasoline, vol% Cq’s, vol% CS’S, vol% c,+cz, wt% Hydrogen, wt % Coke, wt%

A C,’ Gasoline, RON

0 0

-2 +1 +2

+1.2

0 2000

-4 +2 +4

+1.8

2000 2000

-11 +1 +3

+0.4 +0.7 +6.5

+3.6

The tolerance of ZSM-5 to vanadium is shown in Table 9. Doping the ZSM-5 additive with 10,OOO ppm vanadium produces a 22% loss in surface area, but no loss in cracking activity. At comparable vanadium loadings, an REY catalyst lost at least 50% of its surface area and cracking activity.

524

Table 9 Effect of Vanadium on ZSM-5 Product Selectivity [71

Vanadium on ZSM-5 additive, ppm ZSM-5 additive surface area, mZ/g

A Yields, % FF at constant conversion relative to base without ZSM-5.

Gasoline, vol% C4%, vol% Cg’s, vol% CI+CZ, wt% Coke, wt%

A C,+ Gasoline, RON

0 87

-3 +1 +3

+1.4

10,000 68

-3 +2 +2

+1.7

5. OTHER SHAPE SELECTIVE MOLECULAR SIEVES Bench scale catalytic cracking studies have shown that several other shape selective

zeolites and molecular sieves including crystalline silicoaluminophosphates (SAPO) have potential for increasing octane and light olefins at the expense of gasoline yield 115, 26-28]. Aluminophosphates (AIPO,), another interesting class of molecular sieves, have no framework acid sites and therefore appear to have little utility in catalytic cracking applications.

Most of the aforementioned catalytic scoping studies were undertaken with the assumption that small changes in the diameter or shape of the pore might produce significant changes in selectivity. To date, the commercial application of crystalline materials other than ZSM-5 as octane cracking catalysts appears to be limited. Vaughan, in Chapter 3 of this monograph, discusses the structural features and catalytic implications of 8-, lo-, and 12-membered ring molecular sieves including ALPO, and SAPO materials in greater detail.

The scientific and patent literature includes references to the evaluation of at least four other shape selective aluminosilicate zeolites as FCC additives. These are offretite, ZSM-23,ZSM-35, and ZSM-57. Among the non-zeolitic molecular sieves, patents have been issued teaching the use of SAPOJ, SAPO-11, and SAPO-37 in FCC applications.

Offretite Offretite is an intersecting 12- and 8-membered ring system with effective pore

dimensions of 6.4 A (12-MR) and 3.6 X 5.2 (8-MR). Data published by Marcilly et al. [27] show that this zeolite increases gasoline RON by approximately one number for every one volume percent gasoline loss. Like ZSMS, its principal cracked product is

525

propylene. The effects of steaming the offretite catalyst on gasoline yield and octane were not examined.

ZSM-23 This zeolite is characterized by a unidimensional pore structure with a 10-membered

ring tear drop shaped pore. The effective diameter of ZSM-23 is less than that of ZSM- 5. Laboratory studies have indicated that ZSM-23 has fresh activity sufficient for octane enhancement, but loses this activity upon steaming.

ZSM-57 This is an intersecting two-dimensional 8 and 10-membered ring zeolite with pore

dimensions of 5.1 x 5.8 13 (10 MR) and 3.3 x 4.8 8, (8 MR). In its fresh state it can improve octane, but, like ZSM-23 it lacks hydrothermal stability under FCC regenerator conditions.

ZSM-35 ZSM-35, which also has a structure related to ferrierite, exhibited no octane

improvement activity either fresh or following steaming. The gasoline and wet gas (C3=, C4=, and i-CJ yield shifts and accompanying research

octane number improvements observed with addition of equal amounts of ZSM-23, ZSM-35, or ZSM-57 to a commercial equilibrium REY catalyst are compared with ZSM-5 in Table 10. The superior steam stability of ZSM-5 makes it the preferred shape selective zeolite for catalytic cracking applications.

Table 10 Comparison of the Yield-Octane Performance of Z S M J with ZSM-23, -35, and -57 [from ref. 261

REY +ZSM-5 +ZSM-23 +ZSM-35 +ZSM-57 Base Stm’d Calc. Stm’d Calc. Stm’d Calc. Stm’d

Yields, vol%

CJ=+C4=+iC4 19.8 +5.7 +9.1 +0.2 +2.5 +0.4 +16.0 +25 C,+ Gasoline 50.6 -4.8 -73 +os -2.7 -0- -155 -1.4

Octane RON 89.7 +1.1 +1.6 +0.6 +0.4 +0.4 +2.9 +0.1

526

Table 10 (continued) Comparison of the Yield-Octane Performance of ZSM-5 with ZSM-23, -35, and -57 [from ref. 261

Notes: Fixed fluidized bed cracking of heavy vacuum gas oil. All additive catalysts blended to 2 wt% zeolite. Base catalyst is a commercial equilibrium REY. Calcined samples were heated in N2/air to 538OC for 6 hrs. Steamed samples were treated with 6 psi H20, 788OC for 10 hrs.

SAPO-5 This is a hexagonal unidimensional 12-MR silicoaluminophosphate molecular sieve

that has no structural analog among aluminosilicate zeolites. It has a pore size of approximately 8 A and therefore, like USY, might not be considered a shape selective catalyst. However, i t appears to act like other shape selective fluid catalysts in selectively converting gasoline range materials to lighter components. Microactivity test (MAT) comparisons with conventional USY zeolite fluid cracking catalysts show that it produces more C,s and light gas at the expense of gasoline when used as either an additive to the Y zeolite catalyst [29,30] or as a stand-alone cracking catalyst [31]. In these evaluations SAPO-5 appeared to be stable to exposure to 100% steam at 760°C for two hours.

SAPO-11 This silicoaluminophosphate molecular sieve has unidimensional elliptical 10-MR

channels. Microactivity test (MAT) evaluations with SAPO-11 as an additive to a USY catalyst has shown that it produces higher iso-/normal ratios among Cs - C8 paraffin products 1291. It also produces more olefinic C3 and C, products (Table 11). Like SAPOJ, it appeared to be stable to exposure to 100% steam at 760OF for two hours.

SAPO-37 This 12-MR crystalline material is isostructural with faujasite. Most of the

evaluations with SAPO-37 have been as stand-alone cracking catalysts in comparisons with REY and USY [32,331. These comparisons have shown that it can be a more active cracking component than either REY or USY and that it has the potential for producing higher MON and larger quantities of C4's than either REY or USY at equivalent gas oil conversions [33]. However, the patent literature indicates that SAPO- 37 should not be exposed to moisture at room temperature if it is to retain its activity P31.

527

Table 11 Effect of Adding SAPO-11 as an Additive to USY [30]

Catalyst Zeolite

Matrix

Conversion, wt % Gasoline, wt% Gasoline + Alkylate, wt%

Wt. Paraffin Analysis' Isohorma1 C, Isohorma1 C, Isohorma1 C, Isohorma1 C,

Wt. Gas Product Analysis' Olefin/paraffin C, Olefidparaffin C,

Coke, wt%

C 15% USY

A1203 + Kaolin Clay

61.9 46.5 55.3

4.8 13.6 15.6 7.1

2.8 0.55

4.1

D 15% USY + 10% SAPO-11

A1203 + Kaolin Clay

62.9 47.6 56.7

5.7 17.1 16.8 8.1

3.8 0.66

4.0

'Isoparaffin to normal paraffin ratios of compounds having the indicated number of carbon atoms.

'Olefin to paraftin ratio of compounds having the indicated number of carbon atoms.

6. CONCLUSIONS

With the passage of the 1990 Clean Air Act, which mandates minimum levels of oxygenates in gasoline, the use of ZSM-5 in cracking operations becomes a prime contender for the preferred technology to produce light olefin feedstocks for ethers (e.g., MTBE, TAME) or other oxygenate manufacture. Indeed, the ability of ZSM-5 to selectively produce large quantities of light olefins for alkylation and oxygenate manufacture may supersede its octane improvement potential as its most attractive feature. Like many technologies that have preceded it, the application of ZSM-5 in cracking may have its most significant benefit in areas not originally contemplated during its initial development.

528

While the commercial potential of ZSM-5 in FCC may not yet be completely realized, the search for even better shape selective zeolites continues. Inevitably this search will take the same route as did ZSM-5. The development of ZSM-5 is a good example of a common occurrence in catalytic science. Discovery and exploratory experimentation a re closely followed by commercial demonstration. However, a satisfactory understanding of the underlying chemistry of the reaction and the catalyst does not come about until well after the technology is commercialized. Without the experimentation that accompanies this search for knowledge, the opportunity to observe the new and unexplained phenomena that advance the science of catalysis would not and could not be realized.

7. REFERENCES

1.

2.

3.

4.

5.

6.

7.

8.

9.

P. B. Weisz and V. J. Frilette, J. Phys. Chem., 64,382 (1960).

N. Y. Chen, W. E. Garwood, and F. G. Dwyer, Shape Selective Catalysis in Industrial Applications, Marcel Dekker (1989).

C. D. Anderson, F. G. Dwyer, G. Koch, and P. Niiranen, "A New Cracking Catalyst fo r Higher Octane Using ZSM-5", Proceedings of t he Ninth Iberoamerican Symposium on Catalysis, Lisbon, Portugal, July (1984).

G. Anders, I. Burkhardt, U. Illgen, I. W. Schultz and J. Scheve, "The Influence of HZSM-5 Zeolite on the Product Composition After Cracking of High Boiling Hydrocarbon Fractions", Appl. Catal., 62,271 (1990).

J. R. Grzechowiak and A. Masalska, React. Kinet. Catal. Lett., 29,275 (1985).

K. Rajagopalan and G. W. Young, "Hydrocarbon Cracking Selectivities with a Dual Zeolite Fluid Cracking Catalyst Containing REY and ZSM-5", Div. of Petroleum Chem., American Chemical Society Meeting, New Orleans, Aug. 30- Sept. 4, 1987.

S. P. Donnelly, S. Mizrahi, P. T. Sparrell, A. HUSS, P. H. Schipper, and J. A. Herbst, "How ZSM-5 Works in FCC", Div. of Petroleum Chem., American Chemical Society Meeting, New Orleans, Aug. 30-Sept. 4,1987.

R. J. Madon, "Role of ZSM-5 and Ultrastable Y Zeolites for Increasing Gasoline Octane Number", J. Catal., 129,275 (1991).

J. Biswas and I. E. Maxwell, "Octane Enhancement in Fluid Catalytic Cracking. I. Role of ZSM-5 Addition and Reactor Temperature", Applied Catalysis, 58,l-18 (1990).

529

10.

11.

12.

13.

14.

15.

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J. S. Buchanan, "Reactions of Model Compounds over Steamed ZSM-5 a t Simulated FCC Reaction Conditions", Appl. Catal. 74,83 (1991).

F. G. Dwyer, P. H. Schipper, and F. Gorra, "Octane Enhancement in FCC via ZSM-5", presented at the 1987 Natl. Petroleum Refiners Assoc. Annual Mtg., San Antonio, TX (AM-87-63), March 30,1987.

F. G. Dwyer, F. Gorra, and J. A. Herbst, "ZSM-5 in FCC - Potential Impact on Refinery Operations", Fourth CCIC Technical Meeting, June 9, 1986, Tokyo, Japan.

F. G . Dwyer, N. L. Economides, J. A. Herbst, and F. Gorra, "ZSMJ in FCC - Potential Impact on Refinery Operations", Pet. Rev., 41(486), 48, July 1987.

P. H. Schipper, F. G. Dwyer, P. T. Sparrell, S. Mizrahi, and J. A. Herbst, "Zeolite ZSM-5 in Fluid Catalytic Cracking: Performance, Benefits, and Applications", Chapter 5 , American Chemical Society Symp. Ser. 375, p. 64, 1988.

J. Scherzer, Octane Enhancing, Zeolitic FCC Catalysts, Marcel Dekker, pp. 41-109 (1990).

D. A. Pappal and P. H. Schipper, "Comparison of the Performance of Ultrastable Y and Rare Earth Y with or without ZSM-5 in a Fixed Fluidized Bed", Div. of Petroleum Chem., American Chemical Society Meeting, Miami, FL, September 10-15,1989.

D. A. Pappal and P. H. Schipper, "ZSM-5 in Catalytic Cracking: Riser Pilot Plant Gasoline Composition Analyses", Div. of Petroleum Chem., American Chemical Society Meeting, Washington D.C., August 26-31,1990.

S. J. Miller and C. R. Hsieh, "Octane Enhancement in Catalytic Cracking Using High Silica Zeolites", Div. of Petroleum Chem., American Chemical Society Meeting, Washington D.C., August 26-31,1990.

M. F. Elia, E. Iglesias, A. Martinez, and M. A. Perez Pascual, "Effect of Operation Conditions on the Behavior of ZSM-5 Addition to a RE-USY FCC Catalyst", Appl. Catal. 73,195 (1991).

W. 0. Haag, R. M. Lago, and P. B. Weisz, "Transport and Reactivity of Hydrocarbon Molecules in a Shape-selective Zeolite" Faraday Discuss. Chem. SOC., 22 317 (1982).

I. Takano, "Application of FCC Catalysts in Japan", Fifth CCIC Technical Meeting, September 12,1988, Tokyo, Japan.

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22.

23.

24.

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S. J. Yanik, E. J. Demmel, A. P. Humphries, and R. J. Campagna, "FCC Catalysts Containing Shape-Selective Zeolites Boost Octane Number and Yield", Oil Gas J., 83(19), 108 (1985).

A. K. Das, Y. V. Kumar, V. R. Lenin, and S. Ghosh, "Performance of Z S M J Additive in Distillate FCC Units", paper presented at AKZO's Symposium on FCC Catalysts, Scheveningen, Netherlands, June 1991.

M. Kowalczyk, R. B. Miller, L. R. Howard, and E. J. Demmel, "Commercial Experience with ZSM-5 Octane Additives a t Unocal's Lemont Refinery", Symposium on New FCC Technology: Additives, 1990 Spring National AIChE Meeting, Orlando, FL, March 18-22,1990.

R. F. Wormsbecher, A. W. Peters, and J. M. Maselli, "Vanadium Poisoning of Cracking Catalysts: Mechanism of Poisoning and Design of Vanadium Tolerant Catalyst System", J. Catal. 100,130 (1986).

F. G. Dwyer, "The Evolution of Cracking Catalysts and the Challenges of the Future", Gordon Research Conference on Catalysis, June 29,1990.

C. Marcilly, J. M. Deves, and F. Raatz, US Patent 4,992,400, February 12,1991 and US Patent 5,008,000, April 16,1991.

A. S. Krishna, C. R. Hseih, A. R. English, T. A. Pecoraro and C. W. Kuehler, "Additives Improve FCC Process", Hydrocarbon Processing, November 1991, p. 59.

R. J. Pellet, P. K. Coughlin, M. T. Staniulis, G. N. Long, and J. A. Rabo, "Catalytic Cracking Catalysts Comprising Non-Zeolitic Molecular Sieves", US 4,791,083, Dec. 13,1988 to Union Carbide Corporation.

R. J. Pellet, P. K. Coughlin, M. T. Staniulis, G. N. Long, and J. A. Rabo, "Cracking Catalysts and Process Using Non-Zeolitic Molecular Sieves", European Patent 0 202 304 B1, October 16,1991 to UOP.

R. J. Pellet, P. K. Coughlin, M. T. Staniulis, G. N. Long, and J. A. Rabo, "Catalytic Cracking Catalysts Comprising Non-Zeolitic Molecular Sieves", US 4,666,875, May 19, 1987 to Union Carbide Corporation.

R. J. Pellet, P. K. Coughlin, M. T. Staniulis, G. N. Long, and J. A. Rabo, "Catalytic Cracking Process Using Silicoaluminophosphate Molecular Sieves", US 4,842,714, June 27,1989 to UOP.

G. C. Edwards, J.-P. Gilson, and C. V. McDaniel, "Cracking Catalyst", US 4,681,864, July 21,1987 to W. R. Grace & Co.


53 1

CHAPTER 14

ADDITIVES FOR THE CATALYTIC REMOVAL OF FLUID CATALYTIC CRACKING UNIT FLUE GAS POLLUTANTS

ALAK BFIATTACHARYYA AND JIN S. YO0

Amoco Chemical Company, Amoco Research Center Naperville, Illinois 60566

1. INTRODUCTION

Fluid Catalytic Cracking (FCC) is the major method of producing gasoline and middle distillates in the petroleum industry. However, the process is undergoing many changes due to rapidly changing feedstock, product requirements, and environmental regulations. Federal and local authorities are enacting legislation that requires refiners to limit the emissions of pollutants from FCC units. These pollutants are the oxides of sulfur, carbon, and nitrogen.

Of chief concern at this time is control of the concentration of FCC flue gas SQ and SO,. The amount of SO, emitted from an FCC unit regenerator is a function of the quantity of sulfur in the feed, coke yield, and conversion. Depending on the feedstock aromaticity, 4555% of feed sulfur is converted to H,S in the reactor, 3545% remains in the liquid product, and about 10% is deposited on the catalyst in the coke. It is this sulfur in the coke which is oxidized to SO2 (90%) and SO, (10%) in the FCC regenerator. Flue gas scrubbing and feedstock desulfurization are effective means of SO, control but are laborious and capital intensive. Amoco introduced the Ultracat process [l] to reduce SO, levels in FCC units. Based on pilot-plant data, up to 88% SO, reduction appeared possible at that time with a catalyst designed to lower coke make and lower sulfur on coke. This process was not considered a commercial success mainly because of increased FCCU NO, emissions. It is now recognized that the least costly and most convenient alternative is the use of a SO,-reduction catalyst as an additive to the cracking catalyst inventory [2-51. Recent advances in the research, development, and commercialization of effective SO, reduction catalysts will be discussed in this chapter.

532

Two other FCC unit flue gas pollutants are NO and CO. Carbon monoxide control was studied extensively in the 1970s and the problem was effectively solved by using an additive catalyst which oxidized toxic CO to nontoxic CO,. FCC unit NO, (mainly NO) emissions are not presently controlled, though regulations in the future are sure to establish their limits. The challenge for NO, reduction is the fact that unlike SO, and CO, where oxidation of these pollutants is the most effective first step for their removal, NO, should be reduced to N2 under the regenerator conditions. Several approaches for NO, removal will be discussed in this chapter.

2. RESULTS AND DISCUSSION

2.1. Chemistry of Refinery SO, Reduction

The sulfur in the coke is mainly oxidized to SOz (eq 1). Sulfur dioxide should be further oxidized to SO3 (eq 2) so that it can be reactive [6] toward metal oxides to form sulfate (eq 3).

S (in coke) + 0, = SO, (1)

2S0, + 0, = 2S03 (2)

SO3 + MO = MSO, (3)

As the operational temperature of the regenerator is increased, the formation of SO3 is less favored [7]. The regenerator temperature of an FCC unit is usually in the range 675-730 "C. Gibbs free energy change of reaction 2 is -9.5 kJ/mol at 675 "C and -4.4 kJ/mol at 730 "C. At these temperatures, under FCC conditions, the equilibrium ratio of SO, and SO3 concentrations is about 9:l. Catalyzing reaction 2 is one of the major functions of a SO, catalyst. Equation 3 represents the capture of SO, in the regenerator by the catalyst. The sulfated catalyst then moves to the FCCU reactor where the sulfate is reduced by hydrogen and other reducing gases to metal oxide and H,S (eq 4) or metal sulfide (eq 5).

MS04 + 4H2 = MO + H2S + 3H20 (4)

MS + H2O = MO + H2S (6)

533

Metal sulfide can be hydrolyzed in the steam stripper to form the original metal oxide (eq 6). Most refineries are equipped to handle this excess H,S by converting it to elemental sulfur using the Claus reaction (2H,S + SO, = 2H,O + 3s). This is the generally accepted mechanism [6-111 of FCC SO, reduction.

A schematic diagram of an FCCU is shown in Figure 1. A SO, reduction catalyst, thus, should have three functions: oxidation, chemisorption, and reductive decomposition. A material having inactivity in any one of these areas would not be a good SO, reduction catalyst.

Fluegas I

Regenerator

L

Reactor effluent

r

Air from blower -

Regenerator: Stripper:

s+o2 =SO2

2SO2 + 0 2 = 2so3

MS + H20 = MO + HzS

Reactor: SO3 + MO = MS04

MSO4 + 4H2 = MO + H2S + 3H20

MSO4 + 4H2 = MS + 4H20

Figure 1. Schematic diagram of a typical FCC unit and SO, reduction chemistry. Reprinted with permission from reference 16. Copyright 1988 Am. Chem. SOC.

534

2.2. SO, Reduction Catalysts

An effective SO, reduction catalyst must oxidize the SO, to SO3 and form a sulfate. This sulfate has to be stable under the regenerator conditions and be able to release the sulfur as sulfide in the reactor.

In the seventies, many researchers [12-151 studied the absorption of SO, by various metal oxides and also studied the reduction of sulfates by H,. Several of these oxides are not applicable to FCC systems as they will be poison to zeolite. Lowell et al. [13] evaluated 47 metal oxides to find possible use of these metal oxides to absorb SO, from flue gas. The authors selected a group of 16 potentially useful metal oxide absorbants, which include the oxides of aluminum, cerium, and titanium. Interestingly, magnesium oxide was eliminated from the group of potentially useful oxides because of an unfavorable sulfate decomposition temperature. This evaluation was based on the assumption that the absorbants would be regenerated thermally and does not consider the possibility of regeneration under reducing conditions. More work [2,3,5,16] in this area has proven that MgO, A1203, MgAl,O,, La203, and Ce0,-based catalysts are more suitable for FCC operations.

The role of CeO, is to oxidize the SO, to SO3 (eqs 7,8) so that the sulfate formation reaction (eq 3) is more facile. An FCC regenerator contains 1-3% oxygen which regenerates CeO,. Vanadium pentoxide is an excellent oxidation catalyst and is especially useful for the oxidation of SO, to SO3. However, the amount of Vz05 required to carry out the oxidation is too high (5 to 10%). Vanadium pentoxide in such high concentrations may be poisonous to zeolite present in an FCC catalyst. A lower amount (< 3%) of vanadium may be used in combination with CeO, (5-10%) to catalyze the oxidation (eq 2) as well as the reduction (eqs 4 3 ) reactions. Platinum [2,3], a good oxidant for CO -- > CO,, can also be used for this purpose, but it is expensive, may not be very effective under actual FCC regenerator conditions [4], and is also a coke producer. Cerium dioxide, a relatively poor CO combustion catalyst, was thus recognized as the most suitable oxidation component of a SO, reduction catalyst.

2Ce0, + SOz = SO3 + CqO, (7)

CqO, + 950, = 2Ce0, (8)

Several Ce0,-containing catalysts were tested using a thermal analysis technique. The experiment was divided into four zones. Zone A: Under nitrogen, the catalyst sample was heated to 700 "C. Zone B: Nitrogen was replaced by a gas containing 0.32% SO,, 2.0% 02, and balance N,. The flow rate was 200 mL/minute. The temperature was kept constant at 700 "C. This condition was maintained for 30 minutes. In this zone the SO2 was oxidized to SO3 and chemisorbed by the catalyst (eqs 2,3). Zone C: Passage of SO,-containing gas was ceased and replaced by N,. The temperature was reduced to 650 "C. The thermal stability of the sulfate was tested in this 15-minute time zone.

535

Zone D: Nitrogen was replaced by pure hydrogen keeping the temperature constant at 650 "C. Ease of reductive decomposition of the catalyst (eqs 4 3 ) was examined here. This condition was maintained for 20 minutes.

One of the catalysts that was tested by several groups of researchers is CeQ on y-Al,03 [5,16]. This can be conveniently prepared by impregnating y-alumina with Ce(N0,),.6H20 solution followed by drying and calcining at 730 "C for three hours. The amount of CeQ was 12.3%. The y-alumina chemisorbs the SO3 produced by the oxidation of SQ and forms A1,(SO&. This sulfate starts to decompose at 580 "C [17]. Hence one disadvantage of using this catalyst is the fact that any FCC regenerator operating at a temperature higher than 600 "C would have some decomposition of AI,(SO,), (back reaction of equation 3). A thermogravimetric analysis (TGA) of this catalyst is shown in Figure 2. The catalyst was first preheated to 700 "C under N2 (zone A). Then it was exposed to a gas containing 0.32% SO,, 2% O,, and balance Nz at a flow rate of 200 mL/min (zone B). The weight gain of 5.5% indicated in Figure 2 is the amount of SQ formed and absorbed by the alumina to form AI,(SO,),. The TGA indicates that only 2.5% of the alumina is involved in picking up SO3 during the first 15-minute period. This number is called the SO, oxidation and absorption index (SOAI; defined as the percentage of absorbent that is involved in picking up SO3 which is produced in-situ by the oxidation of SO2 in the presence of the catalyst in 15 minutes at a standard TGA condition). This SOAI of 2.5 is considerably lower than other catalysts tested as described later in this chapter. In addition, the activity decreases drastically during the second 15-minute period. Zone C is when the passage of SO,-containing gas was stopped and replaced by pure N2. At this point the temperature was dropped to 650 "C. The TGA clearly shows that at this temperature the AI,(SO,), is thermally unstable and releases some of the SO, it absorbed in zone B. However, the sulfate is reduced very efficiently to regenerate the oxide in the presence of H, (Figure 2, zone D). The low SOAI and the thermal instability of the sulfate under FCC conditions clearly indicate that CeO, on y-alumina is not a very effective SO, catalyst.

0 20 40 60 80 100 Time (min.)

Figure 2. TGA test of a Ce02/A1,03 catalyst. Reprinted with permission from reference 16. Copyright 1988 Am. Chem. SOC.

536

I26

122

118 N - 114- h

v

S 0) ._ ; 110..

106

102

A Ce0,-containing MgO catalyst was prepared by impregnating MgO with an aqueous Ce(N03)3 solution. The composition of the final calcined catalyst was 12.3% CeO,/MgO. Since MgO is much more basic than alumina, it was hoped that it will be much more reactive towards SO3. A TGA analysis is shown in Figure 3. The catalyst gains 28.5% weight in 30 minutes due to SO3 absorption (zone B). This is 5.2 times greater than the CeO2/A1,O3 catalyst. The SOAI of this material is 8.7 which indicates that the SO3 absorptivity of CeO,/MgO is 3.5 times higher than the corresponding alumina catalyst during the first 15 minutes. Linearity of the absorption plot (zone B) indicates that the absorption during the second 15 minutes is as efficient as the first 15 minutes. When the passage of the SQ-containing gas was stopped and replaced by Nz (zone C), unlike the CeO,/Al,O, catalyst, this material did not loose any weight, indicating the thermal stability of MgSO,. Magnesium sulfate is not expected to decompose below 780 "C [17]. Under H,, the sulfate formed reduces at 650 "C; however, the MgO cannot be regenerated as efficiently as the alumina catalyst. About 27% of the absorbed material still remains with the catalyst even after 20 minutes of H, reduction, possibly as MgS or MgSO,,. Fast deactivation of this catalyst is one of the major reasons why CeO,/MgO was not considered as a potential SO, reduction catalyst for FCC units.

..

-~

..

.'

- -

- -

+Zone A* +Zone B * +Zone C +Zone O +

98 3 0 20 40 60 80 100

Time (min.)

Figure 3. TGA test of a CeO,/MgO catalyst. Reprinted with permission from reference 16. Copyright 1988 Am. Chem. SOC.

One of the most promising catalysts tested is the CeQ-containing magnesium aluminate spinel, CeO,/MgAl,O,, and its solid solutions with both alumina (MgAl,O,.xAl,O,) and MgO (MgAl,O.,-yMgO). The solid state chemistry of these materials and their SO, reduction capabilities are discussed in detail in the next section. A Mg and Al double hydroxide, formed usually by the reaction of Mg(NQ), and NaAlO, at pH 8.5-9.5, is calcined to prepare the spinel [2,3,18,19]. The spinel structure is based on a cubic closed-pack array of oxide ions. Typically, the crystallographic unit cell contains 32 oxide ions; one-eighth of the tetrahedral holes (of which there are two per anion) are occupied by the divalent metal ion (Mgz'), and one-half of the octahedral holes (of which there is one per anion) are occupied by the trivalent metal ion (AP+).

537

A TGA analysis of 12.3% CeO, on MgA1204.Mg0 is reported in Figure 4. During preheating (zone A) the material desorbed 7.6% moisture. This catalyst gains 23.3% weight by the absorption of SO3 which is nearly as high as the CeO,/MgO catalyst. The SOAI of this material is 6.7, indicating that this catalyst is 2.7 times more active than the CeO2/Al20, catalyst. Since SO3 absorption by alumina is negligible and MgO is an efficient SO3 absorbing agent, this high SO3 absorption by this catalyst indicates that in a spinel it is the -Mg-O- structural fragment that is reacting with the SO3. The absorption activity of the -Mg-0- structural fragment in spinel is much higher than that of pure MgO. Linearity of the absorption plot (zone B) indicates that the absorption in the second 15-minute period is as efficient as the first 15-minute period. When the passage of the SQ-containing gas was stopped and replaced by pure N, (zone C) no weight loss was observed. This indicates that in zone B only -MgS04- is formed, although this catalyst is composed of nearly 50% alumina. Unlike the Ce02/Mg0 catalyst, this material regenerates efficiently under H, (zone D).

-Zone A- -Zone 8- -Zone C -Zone D+

118

114 t

98

94

g o L : : . : : : : : ; : : : : : : : : : : : : : I 0 20 40 60 80 100

Time (min.)

Figure 4. TGA test of a CeO2/MgA1,O4-Mg0 catalyst. Reprinted with permission from reference 16. Copyright 1988 Am. Chem. Soc.

A key factor for these solid solution spinels to be an effective SO, reduction catalyst lies in the fact that the reduction of -MgS04- to hydrogen sulfide (reaction 4) and regeneration of the spinel can proceed rapidly to establish a catalytic cycle. By incorporating a transition metal component (M) such as the oxides of iron and vanadium into the solid solution structure (Ce0,/MgA1~~,M,04~MgO), it was possible to carry out reaction 4 much faster even at lower reduction temperatures. This allowed a rapid establishment of a catalytic cycle for effective SO, emission control from FCC regenerator units.

Comparison of these three catalysts discussed above clearly indicates that only Ce02/MgA1,04.Mg0 meets the requirement that an effective SO, reduction catalyst should have three functions: oxidation, chemisorption, and reductive decomposition.

2.3. XRD Analysis of Solid Solution Spinels

Here we describe a systematic X-ray diffraction study [20] to characterize the stoichiometric spinel (MgAl,O,) as well as it's solid solutions with both alumina (MgAl,O,.xAl,O,) and magnesia (MgAl,O,-yMgO). The SO, reduction activity was correlated with different solid solution compositions. X-ray diffraction analyses of the catalysts were also made at different stages of the catalysis cycle to gain some insights into the possible catalytic mechanism.

A. Characterization of Spinel Solid Solutions

A stoichiometric magnesium aluminate spinel, MgA1204, was prepxed by the coprecipitation method [2,3,18,19]. The precipitated double hydroxide of Mg and A1 was dried and calcined at 700 "C to form the spinel structure. The X-ray patterns of a nearly stoichiometric spinel and its precursor are shown in Figure 5. The four intense peaks of the stoichiometric spinel, MgA1204, at 28 values 65.25, 59.5,44.8, and 36.9"; are indexed as (440), (511), (400), and (311) planes, respectively. However, in all our XRD characterization work, we will refer to the (440) peak at 65.25 degrees as observed in Figure 5 as the characteristic stoichiometric spinel peak. These XRD results clearly indicate that well crystallized sharp spinel peaks result from calcining the poorly crystallized precursor at 700 "C.

100 - 80 - 60 - 40 - 20

Dried uncalcined

O L 3 27.25 51.50 75.75

80

60

40

20

10 37.50 65.00 92.50 120 2 e degrees 0 37.5 65

2 8 degrees

Figure 5. X-ray diffraction pattern of Figure 6. X-ray diffraction pattern of MgAI,04-(0.176) MgO (A) and its double alumina excess (A, x = 0.47, hydroxide precursor (B). Reprinted with stoichiometric (€3, x = y = 0), and permission from reference 20. Copyright three magnesium-excess spinels 1991 Am. Chem. SOC. (C, y = 0.54; D, y = 1; E, y = 3).

S = Spinel, Ce = CeO,. Reprinted with permission from reference 20. Copyright 1991 Am. Chem. SOC.

539

8.20

8.15

8.10

8.05

8.00

7.95

7.90

The alumina excess spinel, MgA120,~xA1203, and magnesia excess spinel, MgA1204*yMg0, were also prepared by the coprecipitation method. Cerium oxide was then supported on the calcined spinel matrix by impregnation followed by calcination. The XRD patterns for the stoichiometric spinel (y = x = 0), an alumina-rich spinel (x = 0.47), and three magnesia-rich spinels (y = 0.54, 1, and 3) are shown in Figure 6.

B -

A: MgAI;?O4. yMgO, y=3 8: MgAIz04. yMgO, y=l

-

- C: MgAI204- yMgO, y=0.54 D: MgAI2 0 4 , y=X=O

- E: MgA1204. XA1203, ~=0.33

-

I I I 1 I I I I I

One way of looking at these solid solution spinels is the amount of total mole% alumina content. For example the total mole% alumina content of the stoichiometric spinel, MgA1,04 (x = y = 0), is 50. The total alumina content of several solid solution spinels (x = 0-9, y = 0-3) and their corresponding lattice parameter (a) as determined by XRD are given in Table I and plotted in Figure 7.

Figure 7. Vegard’s plot of spinel solid solutions. Reprinted with permission from reference 20. Copyright 1991 Am. Chem. SOC.

540

Table I Total mole% alumina content of several solid solution spinels (x = 0-9, y = 0-3) and their corresponding lattice constants (a). Reprinted with permission from reference 20. Copyright 1991 Am. Chem. Soc.

Molar Value

X Y

9 1 0.33 0.11 0.053 0 0

- -

0.11 0.54 1 1.7 3

Total Alumina Mole %

90.91 66.67 57.14 52.60 51.3 50.0 47.4 39.4 33.33 27.03 20.0

7.95 8.01 8.05 8.076 8.08 8.09 8.095 8.12 8.147 8.095 8.085

67.00

66.50

66.00

Magnesia 65.50 excess Alumina

( 0 1

4 312 2/3 114 0 114213 312 4 f Y + x -

MgAI2O4. yMgO Mgh204 MgAI204. xA1203

Figure 8. Location of 440 peak in alumina excess and magnesia excess spinels. Reprinted with permission from reference 20. Copyright 1991 Am. Chem. SOC.

Two-theta values of these spinel materials are also plotted as a function of the composition in Figure 8. The following conclusions can be drawn from the results.

1. The lattice constant of the stoichiometric spinel is 8.09 A. As the amount of alumina in the solid solution (x) increases, the two-theta for the spinel peaks shifts to higher values,

54 1

and thus the lattice constant decreases linearly and approaches the lattice constant value of pure gamma alumina (also a spinel) of 7.89 A. This observation is consistent with other works in this area [21-241.

2. A continuous solid solution can be formed between alumina and magnesia starting from pure alumina to y = 1 in MgAl,O,.yMgO.

3. As the amount of MgO in the solid solution increases, the lattice constant increases linearly up to the y value of 1, that is, up to a 1:l molar ratio solid solution of MgO and MgA1,0,. When y > 1, no more solid solution exists, spinel and MgO coexist, and the lattice constant reverts to a stoichiometric spinel.

4. By using both X-ray diffraction and optical technique, Alper et al. [25] found the limits of magnesia excess solid solution, MgAl,O,-yMgO, with y values between 10.5 (82 wt% MgO, 18 wt% A1,0,) and 0.62 (39 wt% MgO, 61 wt% Al,O,). The apparent discrepancy between the previous work and the study by Yo0 et al. [20] may be attributed to the different sample preparation/testing conditions. In the earlier work, the limits of solid solution were determined by induction heating of pressed pellet specimens at 1975 "C for two hours and followed by water quench at room temperature. In the present study, granular samples were subjected to straight air calcination at 760 "C for 16 hours, and allowed to cool down gradually in air.

5. There is no evidence to suggest that CeO, interacts with any spinel matrices studied. In fact, CeOz peaks can be used as the internal standard [26].

Steaming was found to be a tool to simulate catalyst aging in a commercial FCC unit. Hydrothermal stability is related to catalyst structure: increasing magnesia content of the spinel tended to increase steam stability. An optimal magnesia range of around 50 mole% excess was found. Magnesia-rich spinels were more stable than bulk magnesia. X-ray diffraction investigation of the magnesia-rich spinels was extended to both fresh and steamed samples (Figures 9, 10). Reflections from several planes were used to determine cell constants. It is significant that, after steaming, the expanded solid solution spinel lattice returns to normal size because the magnesia is exsolved (Figure 11). However, this magnesia could not be detected by XRD until the value of y > 1. The results shown in Figures 9-1 1 lead to these conclusions:

1. In the CeQ/MgAl,O,*yMgO (y = 1) catalyst, CeO, and spinel solid solution, MgA1204*Mg0, are major phases in the virgin (calcined) catalyst. Upon steaming, the solid solution matrix reverted back to the structure close to the stoichiometric spinel (lattice constant 8.0757-8.092 A). Excess magnesia in the solid solution exsolved from the lattice framework, as reflected in the shrunken lattice constant. Some crystalline MgO began to appear on the shoulder of the spinel peaks, but no CeA10, was detected in the steamed sample.

542

I -

2. In CeOz/MgA1204.yMg0 (y = 3) catalyst, C e 4 , MgO, and MgAlZO4 were found in both virgin-calcined and steam-treated samples. Steaming clearly enriches the sample with free MgO, but the spinel lattice constants did not change.

3. In the Ce02/MgAl,04.yMg0 (y = 0.5) catalyst, the expanded lattice of the solid solution was also reduced to the structure close to the stoichiometric spinel on steaming. Although magnesia must have exsolved from the expanded lattice structure on steaming, the XRD failed to detect the free periclase phase. It is speculated that the magnesia molecule that left the expanded spinel lattice must be associated with the lattice structure by some mechanism.

100

80

60

40

20

- Calcined

--- Steamed

Figure 9. XRD pattern of CeQ/MgAlZO4*MgO. Reprinted with permission from reference 20. Copyright 1991 Am. Chem. Soc.

100

80

60

40

20

1 0

Calcined

---- Steamed

-

0 37.5 65 92.5 120

2 8 degrees

Figure 10. XRD pattern of Ce02/MgAl,0.,.3Mg0. Reprinted with permission from reference 20. Copyright 1991 Am. Chem. Soc.

543

- V = Virgin

" (v) S = Steamed

I I I I

114 213 312 4 8.05

Y - MgAl2O4 e yMgO

Figure 11. Change in spinel lattice constant as a function of excess MgO. Reprinted with permission from reference 20. Copyright 1991 Am. Chem. Soc.

B. Composition-Activity Relationship

Calcined cerium oxide-containing solid solution spinels, as well as their steam-treated analogs, were evaluated as SO, reduction catalysts. For a reference catalyst, cerium-containing magnesia, Ce02/Mg0, was also tested under identical conditions. The test gas consisted of 0.93% O2 and 0.1% SQ in nitrogen. Usually, 1.5 wt% of the SO, removal catalyst blended with a commercial equilibrium FCC catalyst was used for testing. The SO, activity of a catalyst is defined as:

SO, pick-up (DeSOx blend) - SO, pick-up (FCC c a t a l y s t alone) SO, Act iv i ty =

concentration of De.SOx i n the blend

The results of SO, activity are plotted against the composition of these catalysts in Figure 12. The analyses of the SO, removal activities of alumina-excess, magnesia-excess, and stoichiometric spinel catalysts (all catalysts contained 12.3% CeOJ reveal the following:

1. Alumina excess spinel catalysts have considerably lower SO, activity than the stoichiometric or magnesia-excess catalysts. The activity decreases in the order: MgA1204*Mg0 > MgA1204 > MgAl2O4.xAl2O3

2. As the value of x increases in the alumina-excess spinel series, the activity decreases.

544

3. As the value of y increases in the magnesia-rich spinel catalysts, the activity increases up to y = 1. This is the most active SO, reduction catalyst. Activity and steam stability decreases where y > 1.

4. The SO, activity of these Ce0,-containing spinel catalysts are directly proportional to the amount of MgO in the catalysts as long as the MgO is maintained in the solid solution state with stoichiometric spinel.

8.20 B 70 -

- 50

7.95 -

7.900 I0 20 i o i o 50 60 ;o 80 90 100

c .- > 0 m .- w

A: 0: C: D: E:

Mole% Alumina

Figure 12. Spinel solid solution lattice constant versus SO, activity. Reprinted with permission from reference 20. Copyright 1991 Am. Chem. SOC.

It is possible that this solid solution spinel provides an optimum structure so that the SO3 molecule can be picked up as sulfate on the -0-Mg-0-A1-0- moiety in the spinel surface. The SO, activity of the virgin CeO,/MgO catalyst was higher than that of the solid solution catalyst, Ce0,/Mg2A1,05, (1:l molar ratio of MgO in MgAl,O,). Contrary to the solid solution system, the pure MgO catalyst was drastically deactivated by the same steam treatment.

For the SO, control catalyst to successfully work in a commercial FCC operation, the SO, picked up in the regenerator must be stripped by the reducing atmosphere in the cracking reactor section. If the sulfur cannot be removed, the sulfur buildup would saturate the SO, control catalyst and render it useless. Commercial tests have shown that the sulfur can easily be removed from the spinel catalysts.

At a temperature of 732 "C the reduction is extremely fast, reaching approximately 90% sulfur removal in the laboratory in less than two minutes. This is important since the regenerated catalyst at temperatures in excess of 700 "C will see the reducing atmosphere

545

of the reactor section for a short time before the temperature is reduced to reactor temperatures (approximately 540 "C) due to the endothermic cracking reaction and heat transfer.

The sulfated spinels resulting from the laboratory SO, activity test were then reduced with H, or propane to regenerate the catalyst. The SO, pick-up test is the oxidation and absorption half cycle, and the H,/propane reduction reaction is the regeneration half cycle [16,27]. The effect of reduction temperature on sulfated spinels was studied and illustrated in Figure 13. Reduction of sulfate becomes easier and requires lower temperature as the alumina content in these spinel materials increases. The reduction efficiencies of these sulfated catalysts decreases in the following order:

A1,03 > MgAl2O,.xAl2O3 > MgA120, > MgA1204*yMg0 > > MgO

)r

> c .- .- c

2

50

40

30

20

10

I l i i I I I

I I I I I 05!?3 593 621 649 677 704 7&C

Reduction temperature

Figure 13. Reduction of sulfated catalysts.

We have studied the XRD pattern of sulfated spinel catalysts. When the degree of sulfation approaches the stoichiometric point, XRD clearly shows the presence of MgSO, peak at two theta = 24.6 degrees. Figure 14 shows the XRD pattern of a highly sulfated spinel having the composition Ce02/MgA1,04.yMg0 where y = 0.54. The MgSO, peak at 24.6 degrees is clearly evident. This peak disappeared when the sulfated catalyst was reduced by hydrogen (Figure 14). More importantly, no MgO peak appeared in the reduced sample. This strongly suggests that -MgO- regenerated by the sulfate reduction may still be associated in some manner with cubic close pack lattice structure of the solid solution.

546

10 37.5 65 92.5 120 2 degrees

Figure 14. XRD patterns of steamed 12.3% CeO,/MgAl,O,. (0.54)MgO Catalyst. A: Sulfated Catalyst; B: H2 Reduced Sulfated Catalyst. Reprinted with permission from reference 20. Copyright 1991 Am. Chem. Soc.

This study shows that the Ce0,-containing magnesia excess spinels, Ce0,/MgA1,04.yMg0, are the best SO, removal catalysts for refinery flue gas. The SO, removal activity increased as the MgO content increased in the solid solution matrix, and became a maximum at y = 1. This Ce0,-containing solid solution material, MgAl,04.Mg0, was found to be hydrothermally very stable and was commercialized by the Katalistiks Unit of UOP as DESOX.

2.4. Preparation Procedures for Spinels

One of the factors that is important when producing a catalyst for commercial use is the ability of the solid particle to withstand the forces exerted on it in a modem FCC unit. Besides thermal shock, the particles undergo great mechanical stress. A measurement of the particles’ ability to withstand these shocks is known as attrition. Typically, a particle is fluidized in a chamber, and the fines generated (due to particle collisions) are measured with respect to time. Plotting the amount of fines generated versus time, one can then calculate the rate of fines production (the slope of the curve). This value is known as the materials attrition index. A hard particle will have low attrition index and thus will stay in the unit longer. This critical physical property is very much dependent on the preparative route used. How the SO, abatement activities and some of the important physical properties of the spinels discussed in the previous section are related to the preparative route used, is discussed in this section.

547

Magnesium aluminate spinels can be prepared by various methods [2,19,26,28] such as (1) thermal cocondensation, (2) coprecipitation, and (3) cogel formation. The results obtained from each of these procedures are discussed below. Here, we will describe the Ce02-containing stoichiometric spinel, MgA1204, and the solid solution spinel, MgA1204.Mg0.

Co-condensation of Oxides-The stoichiometric spinel MgA1204 can be prepared by reacting boehmite-type reactive alumina with high surface area MgO at a temperature higher than 1200 "C (equation 9).

MgO + Al,03 -----> MgA1204 (9)

This procedure is very useful when preparing ceramic-type spinels. The MgO can be replaced by MgC03, if desired, which eventually, at 1200 "C, yields MgO. Repetitive pulverization and calcination steps are required to quantitatively generate high quality spinel. At a temperature below 1200 "C, the spinel formation is very poor. High temperature calcination, which is essential for the preparation of high quality spinel, causes sintering. Because of sintering the spinel material becomes harder and denser (sharper XRD peaks) and loses surface area and pore volume considerably (Table 11).

Table I1 Physical and chemical properties of spinels prepared by different methods. Reprinted with permission from reference 26. Copyright 1989 Am. Chem. Soc.

Preparative Route Spinel Surface Type Area m'lg -

Co-condensation MgA1204 2

Co-precipitation MgA1204 180 MgA1204-Mg0 150

Co-gellation MgA1204 165 MgAl,04*Mg0 169

Pore Vol % SO, Absorbed cclg

0.15 3.0

0.40 10.9 0.41 13.6

0.32 8.8 0.36 11.4

When impregnated with Ce(NO,), solution and calcined, this material does not exhibit good SOx reduction activity, as determined by the thermogravimetric study (Table 11). This low activity is possibly due to the low surface area and pore volume of this material. The amount of SO3 picked up by this catalyst in 15 minutes is only 3 % of its initial weight. The reduction of the sulfate was also very unsatisfactory. The rate of reduction was slow and not all the absorbed sulfur was released even after 10 minutes of hydrogen reduction.

548

The solid solution spinel, MgA1204.Mg0, cannot be prepared by this thermal method because such a high temperature (> 1200 "C) would promote the dissociation of MgO from the solid solution spinel framework (equation 10).

Co-DreciDitation of Hvdroxides-Reaction 9 would be considerably favored if instead of isolated oxides, double hydroxides of magnesium and aluminum are used as spinel precursors. The appropriate double hydroxide can be coprecipitated (equation 11) from the basic solutions of magnesium and aluminum salts [2,19,26,28].

Mg(N03), + NaAl02 ---> MgAl(OH)~.,(NO3),

700 "C MgAl(OH)~.,(NO3), --------> MgAlZO4.MgO + NO2

All spinel preparations described in the XRD analysis section were prepared by the co-precipitation method. This method, unlike the co-condensation method, does not require a very high temperature calcination to produce the spinel structure from the co-precipitated double hydroxide. Usually a temperature of 700-800 "C and 2 to 3 hours of calcination are sufficient to convert the double hydroxide to spinel. The characteristic XRD patterns are shown in Figure 6. The physical properties of the spinels prepared by this method (Table 11) are very different from those prepared by the co-condensation method. The spinels prepared by this method have a very high (- 150 m2/g) surface area but are very soft. When impregnated with Ce(N03h solution and calcined, the resulting spinel, or solid solution spinel catalysts, are very active toward SO, abatement as seen in a previous section. Some of these results are shown in Table 11. We see that the stoichiometric spinel catalyst prepared by the co-precipitation method is nearly four times more active than the same catalyst prepared by the co-condensation method. The solid solution catalyst, MgAl,04.Mg0, is more active than the stoichiometric spinel catalyst.

Co-pel Formation-In addition to the co-precipitation of the two hydroxides it has been found that a very homogeneous mixture of Mg2+ and A13+ species can be obtained by co-gellation [26,29]. This co-gel, usually prepared by combining aqueous slurries of pseudoboehmite alumina, high surface area MgO and a monoprotic acid, is dried and calcined at 700 to 800 "C to produce both stoichiometric and high magnesium spinels (equation 12).

Mg2+ + Acidic AlO(0H) -----> MgAl(OH),

700 "C MgAI(OH), --------> MgAlzO,.MgO

The XRD pattern of this product is identical with the spinels prepared by the co-precipitation method. The surface area and pore volume were also very high (Table 11). Unlike the co-precipitation method this method yields much harder spinels.

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When these spinels are impregnated with Ce(N03)3 and calcined, the resulting catalysts are nearly as active as the co-precipitated catalyst (Table 11). The amount of SO3 absorbed in the oxidation and pick-up half-cycle is 11.4% of its initial weight in 15 minutes which is about four times higher than the catalyst prepared by the co-condensation method. Once again, the high magnesium spinel is about 25% more active than the stoichiometric spinel. Similar to the catalyst prepared by the co-precipitation method, this catalyst releases 70% of the absorbed sulfur within two minutes of hydrogen reduction in a TGA test. Superior hardness of the product, low raw material cost, and high surface area and pore volume are the major advantages of this preparative procedure.

2.5. Commercial Trial Results

Research at Arc0 Petroleum Products Company, as described in Section 2.3, led to the development of SO, reduction catalysts HRD-276 and HRD-277. HRD-276 was a cerium dioxide-containing stoichiometric spinel, MgAI,O,, commercialized in 1983. HRD-277 was a cerium dioxide-containing solid solution spinel, MgAl,O,.MgO, marketed in 1984. Commercial tests showed that these materials achieved SO, emission reduction of over 70% for a 0.4% sulfur feed with a catalyst concentration of only 0.7 to 3% [30]. HRD-277 showed much higher activity than HRD-276, verifying laboratory results. For example, in a single commercial FCC unit, HRD-277 reduced SO, emissions to the same level as HRD-276, but, at nearly one-half the daily addition rate.

The chemical composition and physical properties of these materials were further improved and a new catalyst, DESOX KD-310, was introduced in 1986 by Katalistiks International Inc. (now a unit of UOP). This catalyst is considerably more attrition resistant than the HRD series and also contains a minor amount of a transition metal species which promotes the reduction of the sulfated spinel. This DESOX KD-310, along with two other commercially available, alumina based SO, catalysts, were evaluated commercially [3 11.

A trial of Katalistiks DESOX KD-310 with a competitor’s alumina based catalyst was conducted in a UOP high efficiency FCC unit operating at a 18,000 B/D feed rate and under complete CO combustion. The results of this trial are summarized in Table 111. Initially, addition of 120 pounds of the competitor’s alumina-based catalyst was required to reduce the SO, emission from 400 ppm to 220 ppm. After about four months the refiner switched from this competitor’s catalyst to DESOX catalyst at an addition rate of only 50 pounds per day. The SO, emission remained at the 220 ppm level. This showed that the DESOX catalyst is about 2-3 times more active than the alumina-based competitor’s catalyst. In addition, the concentration of DESOX in the inventory was 0.8% compared to an estimated 3.5% for the alumina-based catalyst.

Less effective SO, additives have been found to decrease FCC yields by acting as diluents to the cracking reactions. Use of a spinel based DESOX catalyst minimized these diluent yield effects. Because of the much superior activity of the DESOX catalyst, less amount is needed to remove equivalent amounts of SO, [31].

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Table I11 Summary of Trial Results: DESOX versus Competitor

Parameters Competitive Catalyst DESOX

Unit Design UOP UOP

Addition Rate, lbs/Day 120 50 % SO, Reduction 45 45

Unit Feed Rate, B/D 18,000 18,000 Regenerator Temp., OF 1,350 1,350

% Agent Concentration 3.5 0.8

2.6. Competitive Materials for FCC SO, Reduction

Regulations limiting the emissions of SO, from FCC unit regenerators have prompted the development and commercialization of the SO, control technology. Historically, the major companies involved in the catalyst technology development are Ammo, Chevron, Arco, Texaco, Union Oil, Engelhard, Davison, and Katalistiks.

Initially, there were two approaches to commercialize SO, control agents. In one, the SO, capture agent was incorporated within the FCC catalyst particle (single particle). In the other, the SO, agent is a separate particle (dual particle), preferably having physical properties very similar to the FCC catalysts. Examples of single particle technology are Davison’s DAS 250, DAS 300, and Engelhard’s Ultrasox-560 introduced about ten years ago. Davison claimed [8,32] that the SO, emission reductions ranging from 30 to 80% were achieved with the reduction catalyst comprising from 60 to 91% of the catalyst inventory. Engelhard claimed up to 80% SO, reduction in a commercial trial using Ultrasox-560 [33]. This single particle approach was soon replaced by the dual particle approach. The main advantage of using a separate SO, catalyst is that it permits the independent control of SO, reduction and cracking activity and selectivity. The amount of SO, additive used can be varied depending on the degree of SO, reduction required. This is very important when using feedstocks which vary in sulfur content. Most catalyst developments for the past several years were based on the dual particle technology.

By the early 1980s, work on SOx removing catalysts had accelerated with the use of alumina [34,35] and rare earth metal oxides [36]. Chevron [32] provided results of commercial scale tests of their Transox catalyst technology. Commercial tests conducted at Chevron’s Hawaii and El Paso refineries, showed that Transox achieved SO, emission reduction of 50 to 79% for a 0.98 to 1.06 wt% sulfur feed. Davison introduced a rare earth on alumina-based material Additive R in the early eighties. Davison provided test data showing SO, reductions up to 63% with Additive R displacing 6% of the cracking catalyst inventory [32]. Subsequently, the physical properties of this material were improved and the catalyst enjoyed commercial success. However, as described in

551

Section 2.5, DESOX KD-310 emerged as the most effective and technologically advanced SO, reduction catalyst available today.

Recently, in a U.S. patent, E. H. van Broekhoven of Akzo N.V. (Netherlands) described the use of anionic clay based materials as SO, reduction catalysts [37l. These anionic clays may have a hydrotalcite, an ettringite, or a hydrocalumite-type layered structure and have a general formula M:+M~+ (OH) (n+m)~:;a , where I@+ and M3+ are divalent and trivalent metals, m and n are their relative proportions, A is an anion such as OH-, CO?-, SO:-, etc, and a is the charge of the anion A. An example of one sucb material is the regular hydrotalcite, Mg&(OH),,CO,. It may be noted here that these clay materials are thermally unstable and may produce MgO and MgA1,0, under FCC regenerator conditions [38,39].

2.7. Refinery NO, Reduction

During the combustion of nitrogen bearing fuels, nitrogen oxides labeled as NO, are produced [40] by two major routes: (1) oxidation of molecular nitrogen from the combustion air and (2) oxidation of nitrogen which is chemically bound in the fuel and is released during combustion. These routes lead to the formation of so-called "thermal NO," and "fuel NO,," respectively. Considering the huge amount of NO, being released every year (about 25 million tons/year) and the toxic effects of NO, on humans, animals, and plants, it is important that effective and economical means of reducing NO, emissions from various sources be developed [41,42]. Scientists and Engineers have been challenged by the difficulty of finding either a way to prevent the formation of NO, (combustion modification) or a way to convert NO, chemically to a nonpollutant. Here we will focus our attention on the latter method in connection with FCCU flue gas treatment. The fraction of NO, emitted from FCC units is still very small compared to the total emission from other sources. Currently, the emission of FCCU NO, is not very tightly regulated. However, in California, there are three recent South Coast Air Quality Management District (SCAQMD) rules that directly impact NO, emissions from refinery boilers and process heaters [43]. These types of restrictions are expected to increase in the future. Researchers are already paying a lot of attention to find ways to reduce the emission of NO, from various sources including FCC units.

A. Chemistry of Nitrogen Oxides

Equations (13-15) indicate that at a reasonably high temperature (600 "C) the composition of NO, is primarily NO because reaction 15 starts at 160 "C and is complete at about

552

600 "C [44]. Since all FCCU regenerators operate at a temperature higher than 600 "C, NO, in the regenerator flue gas consists primarily of NO.

25 "C 3N0 <=====> N20 + NO,

This NO reacts with different reducing agents in the gas phase to form products that are basically nonpollutants (reactions 16-21).

2N0 + 4NH3 + 202 = 3N2 + 6H20 (18)

2N0 + 2H2 = N2 + 2H20 (19)

4N0 + CH4 = C02 + 2N2 + 2H20 (20)

2N0 + 2CO = 2C02 + Nz (21)

These reactions are used in various ways to find methods for the reduction of NO,.

B. Catalytic Reduction of NO,

An attempt to reduce NO, by NH, (reactions 16-18) using a catalyst is called Selective Catalytic Reduction (SCR) [45-491. SCR uses a metal oxide or copper exchanged zeolite catalyst at 200-400 "C to attain over 80% NO, reduction using 0.85-0.95 mole of NH3 per mole of NO,. Various catalysts have been developed that have an optimum reaction temperature of 300 to 400 "C, so that SCR is applied to a boiler economizer outlet at 300 to 400 "C. Japan is the leader in commercializing this technology and in the mid-eighties there were over 140 commercial units [46] operating. The main disadvantage of using this method for FCCU NO, abatement is the poisoning of the catalyst by the presence of SO,. The formation and deposition of (NH4),S04 and (NH4)HS04 adversely affects the catalytic reduction of NO.

553

Much less work has been done on Non-selective Catalytic Reduction (NSCR) where reducing agents such as CO, H2, CH, or their mixtures are used (reactions 19-21) for the reduction of NO using the same type of catalysts as used in SCR [50-531. The reduction temperature is usually about 300-600 "C and is dependent upon the type of the active catalyst and the nature of the reducing agent used. Several other methods such as the liquid membrane technique, electron beam process, and photochemical disproportionation are also reported in literature [54-551. The NSCR method may be more suitable than the SCR method for FCC processes because of the possible availability of CO in the regenerator.

Fluid catalytic cracking units operating under partial combustion mode may contain a sufficient amount of CO to promote reaction 2 1 in the presence of a suitable catalyst. One reaction that may compete with reaction 21 is the exothermic oxidation of CO to C02 by oxygen (reaction 22).

CO + %02 = COz AH = -67.6 Kcal/Mole (22)

In partial combustion mode, in the absence of any excess amount of oxygen, reaction 22 may not be able to compete with reaction 21, particularly when a suitable catalyst is used to promote reaction 2 1 .

Bhattacharyya et al. [29] and Yo0 et al. [56] have examined several catalysts to convert NO to N2 including the spinel-based SO, reduction catalyst, 12.3 % CeOz/MgA120,.Mg0. A 2: 1 molar ratio of CO and NO was used for the laboratory reactions. Nitrogen was used as the carrier gas. The amount of NO in the gas blend was kept at 272 ppm. A quartz tube (l-inch diameter) equipped with a fitted disc was used as the reactor. A Beckman 951A NO/NO, analyzer was used for the measurement of both NO and NO, downstream from the reactor. The NO and CO gas mixture was passed through the reactor at a temperature of 500 to 1450 O F (260-788 "C) and the NO emission from the reactor was continuously measured by the NO/NO, analyzer. A space velocity of 520/h was maintained throughout the reaction.

It was observed that an equilibrium FCC catalyst does not have any NO, reduction activity (Figure 15). Figure 15 also shows that a 5% blend of solid solution spinel, MgA120,.Mg0, in an FCC catalyst is inactive towards NO, reduction by CO. However, under the same reaction conditions a 5% blend of 12.3% Ce02/MgA120,.Mg0, is very active (Figure 15). A complete reduction of NO, was evident at 1100 "F (593 "C) and above. Since both the equilibrium FCC catalyst and the spinel base, MgA1204.Mg0, were found to be inactive towards NO, reduction, it is likely that the redox properties of Ce02 are responsible for this activity. This was confirmed by the fact that a 5 % blend of 12.3% Ce02/Mg0 in an FCC catalyst was found to be as active as a similar blend of 12.3% Ce0,/MgA120,~Mg0 (Figure 15). However, in an FCC regenerator, in addition to NO,, there will be a sufficient amount of SO, present and CeO2/MgAl20,.Mg0 would be a much better choice because of its excellent hydrothermal stability and higher SO, reduction activity. In fact, this catalyst could be an effective [29,56] simultaneous SOx-NO, control catalyst because of the following sequence of reactions.

554

1. CeO, may be reduced to CqO, by both SO2 and CO (reactions 23,24).

2Ce0, + SO2 = CqO, + SO3 (23)

2Ce0, + CO = CqO, + CO, (24)

2. Sulfur trioxide, SO3, may be chemisorbed by the spinel as discussed before. Sulfation of the spinel will not affect the NO, control process since spinel does not participate in the NO, reaction (Figure 15).

3. The Ce(II1) oxide, Ce203, may now be oxidized to CeO, by NO or 0, (reactions 25, 26).

2Cq03 + 2N0 = 4Ce0, + N2 (25)

2Cq03 + O2 = 4Ce0, (26)

In partial combustion mode, that is, in the absence of any excess oxygen, reaction 25 is favored over reaction 26.

100

80

4 60 E)

-0 a, 0

x

40

0 300 500 700 900 1100 1300 1500

Temperature ( O F )

0 = 5% MgO/CeO2 in FCC A = 5% MgA1204 MgO/CeO in FCC

Figure 15. NO, reduction activities of MgAl,O,.MgO, 10% Ce/MgA1,04 *MgO and 10% Ce/MgO: NO, level = 272 ppm.

Under the conditions discussed, simultaneous reduction of SO, and NO, is possible as observed in the laboratory. A proposed mechanistic cycle showing simultaneous reduction of SO, and NO, from an FCC unit operating under partial combustion mode is shown in Figure 16.

555

= Two active sites of 12.3% CeO~lMgA1~04 MgO

Figure 16. A proposed mechanistic cycle showing simultaneous reduction of SO, and NO, from an FCC unit operating under partial combustion mode.

2.8. Refinery CO Control

An important effect of CO oxidation in an FCC unit is the drastic reduction in the CO content of the regenerator flue gas. This eliminates the need for the CO boiler and improves the environment by reducing one of the most poisonous components of the FCCU flue gas. A typical modem U.S. refinery with a 50,000 BPD FCC unit processing a low sulfur, low metals content vacuum gas oil feed and operating on complete combustion mode typically emits about 200 ppm of CO in its flue gas outlet (1.5 to 2 TPD).

During the cracking of gas oil in the FCC reactor, coke is deposited on the FCC catalyst. One of the many factors which determines the amount of coke deposited on the catalyst is the type and composition of the FCC catalyst. This coke deactivates the catalyst and has to be burned off to regenerate the catalyst. The main function of the FCC regenerator is the combustion of the coke deposited on the FCC catalyst. The coke can be oxidized to CO and/or C02.

c + 0 2 = co2 AH = -94 KcaVMole (27)

c + so2 = co AH = -26.4 Kcal/Mole (28)

CO + SOz = C02 AH = -67.6 K d M o l e (22)

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A substantial portion of the heat thus generated is transferred to the regenerated catalyst which in turn supplies the heat requirements of the FCC reactor. In older FCC units, the coke was oxidized to an equimolar mixture of CO and COz [57,58]. If the remaining CO is further oxidized to COz by using a CO oxidation promoter, the heat release can be increased by as much as 50%, because of the higher exothermicity of the total combustion. Beyond the dense catalyst bed, where catalyst is not available to act as a heat sink, the high heat release can raise the gas temperature above 760 "C causing equipment damage. Ammo's non-promoted Ultracat Regeneration Process circulates sufficient catalyst through the dilute phase and cyclones to absorb this heat and return it to the dense bed [59].

Because the FCC process operates in heat balance, changes in the regenerator combustion will affect the reactor operation. If all the CO is burned to C02, less amounts of coke will be required. Normally, a lower coke yield would be associated with a lower gas-oil conversion. However, the lower carbon on regenerated catalyst (CRC) associated with full CO bum will offset this with a better coke-conversion selectivity.

Catalytic Combustion Promoters-Researchers at Mobil discovered [60,61] that some of the group VIII metals, platinum in particular, could be incorporated in an FCC catalyst formulation at very low levels (1-3 ppmw) to effectively catalyze the combustion of CO to COz in the dense bed. They also found that at these low levels these group VIII metals did not catalyze undesirable dehydrogenation reactions during cracking. The cracking reactions were thus unaltered when such a CO combustion promoter was added to the catalyst inventory.

Catalyst manufacturers utilized this technology many ways based on how these CO promoter species were introduced. One way was to incorporate 1-3 ppmw amounts of platinum directly into the FCC catalyst matrix. This way, the addition of CO promoter is achieved by adding fresh FCC catalyst to the regenerator. A liquid promoter containing a solution of chloroplatinic acid was also developed. The mode of addition was to spray the liquid onto the catalyst as it circulated through the unit. Because of corrosion difficulties, this method has been largely supplanted by use of dry powder additives. Another approach was to develop a solid additive. These additives typically contain 500-1000 ppmw of platinum. It is added to the FCC unit via a small metering system, or in small batches, which is independent of the fresh cracking catalyst addition system. In the United States the separate additive approach is usually used while in Europe it is very common to incorporate the additive component in the fresh cracking catalyst. The biggest advantage of the additive approach is that the amount of promoter addition may be varied depending on the degree of CO oxidation desired.

The preparation of the popular solid additive containing 500 to lo00 ppmw of platinum on a gamma alumina or on a gamma alumina precursor can be achieved by impregnation or pore filling [60,61]. The support alumina usually has a surface area of 150 to 200 m2/g and an average particle size of about 80 microns. The pores of this alumina may be partially or completely filled by using appropriate concentrations of [Pt(NH3),]C1, in water. The product may be sold as is or may be dried.

557

Platinum catalyzed CO combustion occurs readily in the dense phase at temperatures well below 700 "C. Guegin [62] has reported promoted combustion occurring satisfactorily at 650 "C in a commercial FCC unit. Through the use of a combustion promoter, the after bum problem is greatly reduced. By oxidizing CO in the dense phase, the potential heat release from burning in the dilute phase is also greatly reduced. For example, a commercial unit which switched from non-promoted to promoted operation reduced the amount of after bum by 55 "C [58,63] (Table IV). A number of other metals have been proposed, but only Pt and Pd [64] are currently used.

Table IV Reduction of After-burn Temperature with Promoted CO Combustion. Reprinted with permission from reference 58.

Component Temperature "C

Without Promoter With Promoter

Typical flue gas 700 660

Dense phase 660 675

After-burn, AT +40 -15

The quantity of CO oxidation promoter used can also be a variable in promoted CO combustion systems, particularly when a separate additive is used. The systems can be classified as fully promoted or partially promoted. The degree of promotion is best defined by the response of the unit to additional promoter. A partially promoted system is one where an increase in promoter addition results in a decrease in the dilute-dense phase AT. A fully promoted system, on the other hand, sees no effect upon after-burn AT when the promoter concentration is increased. The effect of promoter concentration upon the regenerator after-burn was studied [63] in commercial units. The changes in the dilute-dense phase AT was mainly due to changes in dilute phase temperature. Variations in promoter concentration only have a small effect on the dense phase temperature but greatly affect dilute phase temperature, depending on the catalyst entrainment there.

The CO oxidation reaction (reaction 22) competes with the coke burning reaction (reactions 28, 27) for the available oxygen in the regenerator. If the oxygen level is low a high level of CO oxidation promoter will enhance the combustion of CO and retard burning of the coke. In one refinery where large variations in promoter content have been used, Upson, et al. [63] have found that a fully promoted catalyst with essentially no excess oxygen in the regenerator has equilibrated at a coke content on the regenerated catalyst (CRC) of about 0.1 wt% higher than a partially promoted system at the same degree of CO combustion.

558

The oxygen concentration in the regenerator affects the coke burning which also affects the CRC values. Upson et al. have observed that an increase in regenerator temperature of 20 "C at a constant flue gas O2 content results in a CRC decrease from 0.22 to 0.15 wt% [63]. The CRC is an important parameter which affects the activity of the FCC catalyst in the reactor. The high activity of the zeolite cracking catalysts can be severely reduced by the presence of high concentrations of CRC. In a study [58] , the CRC was deliberately varied in commercial units by adjusting regenerator air rate and allowing the unit to equilibrate to a new set of operating conditions. The results show that the catalyst with the highest activity was the most severely deactivated by the presence of coke. Reduced CRC also improved the selectivity to gasoline.

The regenerator temperature is usually controlled at some maximum temperature allowable by regenerator metallurgy. Variations in FCC feed quality may result in severe fluctuations in the regenerator temperature if the extent of CO combustion is not controlled. The degree of CO oxidation may be controlled to hold the regenerator temperature constant at the desired maximum even though significant feed quality changes have occurred. This flexibility is achieved through regenerator air rate control and CO combustion promoter addition.

Any changes in the extent of CO combustion, either from operational changes or by use of CO combustion promoter, affect the unit operation as discussed before. The exact response will result from a complex interrelation of carbon on regenerated catalyst, regenerated catalyst temperature, and catalyst circulation rate, and must be evaluated on a case by case basis.

3. CONCLUSION

One of the major developments in the areas of FCCU additives is the discovery of the magnesium aluminate spinel-based SO, reduction catalyst, DESOX. This catalyst, now available commercially, is composed primarily of CeO, and MgAI,O,.MgO. The latter is a solid solution of MgO and stoichiometric spinel, MgAl,O,. The SO, removal activity of this well-characterized catalyst not only depends on the composition but also on the method of preparation. This same material was also found to be effective for the reduction of NO to N2 in the FCC regenerator, especially under partial combustion mode. A problem with another major FCCU flue gas pollutant, CO, was also solved by designing a solid additive. This additive promotes the oxidation of CO to C02 without any adverse effect on the cracking reactions and is composed of minute amounts of platinum or other group VIII metals on alumina.

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4. ACKNOWLEDGEMENTS

The authors thank Douglas Rundell, Emmett Burk, Joseph Powell, John Magee, John Karch, Cecelia Radlowski, Gerald Woltermann, and William Cormier for valuable suggestions. The original work on DESOX was done at ARC0 Petroleum Products Company, Harvey Research Center, Harvey, Illinois, and commercial development was done by Katalistiks, Baltimore, Maryland.

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60 Pennick, J. E. U.S. Patent 4,064,039, 1978.

61 Schwartz, A. B. U.S. Patents 4,072,600, 1978, and 4,093,535, 1978.

62 Guegin, Y. Preprints from Katalistiks 1st FCCU Symposium, October, 1980, Bordeaux.

63 Upson, L. L.; Van der Zwan, H. Paper Presented at Katalistiks 7th Annual FCC Symposium, Venice, Italy, May 1986.

64 Ambur Chemical Company, Sales Literature, received January 1992.


563

CHAPTER 15 ENVIRONMENTAL CONSIDERATIONS AFFECTING FCC

RICHARD E. EVANS, GEORGE P. QUINN

Refining and Engineering Department Amoco Oil Company

200 E. Randolph; Chicago, Illinois 60601

1. INTRODUCTION

The Fluid Catalytic Cracking Process has recently celebrated its 50th anniversary, in which time it has become the primary refining conversion process for upgrading higher boiling components of crude to gasoline blending stocks. The prominence and longevity of the FCC process is clearly the result of the unique flexibility of the fluid process to respond to the changing refining demands. As we enter the "Environmental Decade" of the ~ O ' S , the flexibility of the FCC process to respond to these issues will play an important role in the overall refining industry's solutions. There is little doubt that environmental issues have always been important in the FCC process, and the record of improvements is there for all to see. Improvements that were made without their being required by regulation.

The challenge of the 90's for the FCC process is to work with the various regulatory agencies to meet the goals of legislative initiatives, such as the Clean Air Act of 1990. Only in this way will it be possible to prevent environmental regulation from being viewed as burdensome restrictions. Rather, the proper view should be meeting the challenge of environmental regulation with good science and technology, is good business. The environmental impact of the FCC process is large. All FCC products taken together represent a total of 17% of the US energy demand of about 80 Quads, as shown in Figure 1. FCC naphtha in Figure 2 makes up about 37% of the motor gasoline pool [l]. FCC coke burned in the regenerators of US FCC's totals over 15.5 million tons per year of a sulfur containing solid fuel. For comparison, this would be equivalent to about 2% of the coal burned in the US for all purposes [2]. This large potential impact is causing various agencies to look at the FCC process with regulation in mind. Technical innovation in the past has shown process improvements that were good for business were often times good for the environment.

To put the current environmental issues in proper perspective, it is instructive to try to see how the FCC process has evolved over the last 50 years, and in particular, how the FCC process has impacted the environment. To do this, we will review the history of a typical FCC originally built in 1950, and revamped to state-of-the-art technology in 1970 and 1990. In doing this some of the highlights in the development and improvement of the FCC process will be apparent. The important point of this story is that economically driven improvements in the operation, design and catalyst technology did impact on the

564

U.S. ENERGY DEMAND 1980

8 1 QUADRILLION BTU’S

- FCC PRODUCTS 14 QUADRILLION

FIGURE 1

TOTAL U.S. GASOLINE POOL

7.3 MM B/D

\ FCC NAPHTHA 2.7 MM B/D

FIGURE 2

565

quality of the fuels and emissions from the FCC process. What will also be apparent is that over time the regulatory impact on the FCC process has grown.

Having set the FCC process development and environmental issues in proper perspective over the last 50 years will be the basis for the primary purpose of this chapter, a look ahead to the "Environmental Decade" of the 1990's. The intent here is to get an idea of what may happen, and why. The review will cover process emissions, waste disposal, and the effect of product quality on environment. In addition to looking at the possible course of regulation on the FCC process and its products, there will be an attempt to speculate at what process changes and catalyst changes will be needed to respond to the environmental regulations.

2. HISTORY OF FCC UNIT BUILT IN 1950

The FCC process really came into prominence in the period immediately following World War 11. The first fluid unit went on line in 1942 as a result of a consortium set up as part of the war effort to increase the production of aviation gasoline [3]. The period of growth and prosperity that followed the war saw rapid development of the FCC process for the manufacturing of motor gasoline. As a result by far the majority of the FCC units running today were built during the "Cat Cracker Boom" after World War 11, 1946 to 1960. They have been revamped numerous times to keep the design near state-of-the-art, employing the latest catalyst and process technology. To understand and tie together the evolution of the various environmental issues relating to the FCC process, it is essential to see how the process itself evolved over the last 40 years.

Figure 3 shows the case study unit that shall be considered as it was originally built in 1950. The regenerator was designed to run at about 1050"F, and had 3 stages of internal cyclones. On the flue gas line from the regenerator was an Electrostatic Precipitator (ESP) to catch fines, and return them to the regenerator. On the reactor side of the process, 20 mb/d of feed was preheated in a furnace to about 700°F. The feed was then charged through a sloping riser into a large bed of catalyst in the reactor. Conversion, even with substantial recycle, was only 60%. The reactor vessel also had 3 stages of internal cyclones. On the fractionator bottoms a slurry settler was used to decant the bottoms product to remove some catalyst fines, which were returned to the process via slurry recycle. The vapor recovery unit (VRU) for the process produced a fuel gas that contained almost all of the H,S and propane/propylene. The VRU also produced a butane/butylene stream and a total naphtha stream.

By 1970 this same unit had undergone several revamps, and looked like the unit in Figure 4. The major vessels were still in place, but much of the equipment had been modified. In the regenerator, only 2 stages of regenerator cyclones were used. On the flue gas system the ESP disappeared, and a boiler to consume CO from the regenerator appeared. Also disappearing was the preheat furnace, and the large bed of catalyst in the reactor vessel. By 1970 the feed rate of the unit had increased to 40 mb/d, and conversion was about 67%. In the VRU a depropanizer had been added to collect propane and propylene, and a fuel gas scrubber was in place to remove H,S from the refinery gas.

After several more revamps the case study unit by 1990 was as shown in Figure 5. The regenerator was operating at 1300" F in full CO burn. As a result the CO boiler was

566

FCC COMPLEX AS BUILT 1950

OIL r'

c

REACTOR AND RECOVERY SYSTEMS

FIGURE 3

567

FCC COMPLEX REVAMPED IN 1970

IEGENER ATOR SYSTFMS CO BOILER

OIL r'

i c

FIGURE 4

REACTOR AND RECOVERY SYSTEMS

OEETHANIZER DEPROPANIZER DEBUTANIZER FUEL GAS SCRUBBER SLURRY SEl lLER

568

gone, but was replaced by a sensible heat recovery unit. While a few units had external 3rd stage cyclone systems, most units were as shown in Figure 5 with an ESP on the flue gas system. The reactor side also had a number of differences. The sloped riser was replaced with a vertical riser that terminated in some sort of separation device, usually a rough cut cyclone. The unit was now cracking 60 mb/d of feed at a riser outlet temperature of 985"F, and conversion was 76%. On the fractionator the slurry settler had been removed. While there were no major changes in the VRU, there were substantial changes in the products to be separated.

environmental impacts are obvious, such as the addition of fuel gas scrubbers. Others raise curious questions, such as why were the original ESP's dropped, and why did they return in recent years? Finally, how did the significant changes in process conditions and catalyst technology which clearly altered the yields of the process impact the environmental quality of these products? These will be addressed in the next 3 sections of this chapter.

In looking at the evolution of the hardware in the FCC process, some

3. HISTORY OF FCC REGENERATOR STACK EMISSIONS

By far the most obvious impact from an environmental perspective of the FCC process on the community in which it operates is at the stack. A dusty stack or visible plume can be seen by thousands of people for many miles. While at times the visible nature of the plume is an indication of excess pollution, it is not always the case. Nonetheless, attention has properly been focused on the FCC stack as a pollution source from the combustion of coke. This section will focus on the stack emissions of particulates, carbon monoxide, sulfur oxides, and nitrogen oxides.

3.1 Particulates

not for environmental reasons, but rather it was dictated by process economics. Referring to the case study unit as it was built in 1950, Figure 3, catalyst losses from the regenerator were about 1 ton per hour even with 3 stages of cyclones. With only 2 stages the losses would have been over 5 tons/hr. Such was the state of cyclone technology at the time. Indeed, even with 3 stages of cyclones, most units found it justifiable to have an Electrostatic Precipitator (ESP) down stream of the regenerator to collect about 90% of the dust in the flue gas. This catalyst was collected for economic reason, and returned to the process. Net losses to the stack were thus on the order of 2 1/2 tons per day, or about 12 lbs of particulates per lo00 lbs coke burn. A typical stack opacity would have been 30-50%, and catalyst losses to the stack were minimized to the extent that was economically feasible.

cyclones had improved significantly. Most notably improved refractory systems and more flexible cyclone hanger systems allowed high velocities, which improved efficiency, and a more mechanically reliable system which lasted years rather than months. Moreover, as shown in Figure 6, the basic design of the cyclones changed from 1950 to 1990. Among the key design changes were longer barrels, addition of dust bowls and flapper valves at the bottoms of the diplegs, and reducing the size of the secondary diplegs to match the

In 1950 the typical FCC regenerator had 3 stages of internal cyclones. This was

By 1970 improvements had occurred in cyclone technology, and the efficiency of

569

F C C COMPLEX REVAMPED IN 1990

I -ECTROSTATIC PRECIPITATOR

REACTOR AND RFCOVERY SYSTEMS

DEETHANIZER DEPROPANIZER DEBUTANIZER FUEL GAS SCRUBBER

FIGURE 5

FCC CYCLONE EVOLUTION 1950

CYCLONF

I

FIGURE 6

t

1 LONGER L/D

DUST BOWL

TRICKLE VALVE

571

catalyst flux. As a result, the case study unit in 1970, Figure 4, had increased feed rate and coke bum significantly. Yet, due to the higher efficiency cyclones, catalyst losses from the regenerator had actually dropped to about 1.5 to 2 tons per day. Two stages of cyclones without an ESP were achieving a particulates loss of about 5 lbs per lo00 lbs coke burn, and stack opacity was typically 20 to 30%, both of which were well below the 1950 levels.

In the late 1970’s and early 1980’s improvements in catalyst manufacturing hardened the catalyst, reducing attrition rates. Also during the 1970’s and 1980’s a number of improved schemes have been tried. In some units a 3rd stage of cyclones was installed in the vessel with the flow from the diplegs withdrawn from the vessel to a collection hopper. This scheme only achieved moderate improvements in stack losses. For units that installed Power Recovery Turbines (PRT), an external 3rd stage was used to protect the turbine. While successfully protecting the turbine from high catalyst loadings in upset conditions, these 3rd stage separators only removed about 40% of the remaining catalyst from the flue gas during normal operations. A few units installed flue gas scrubbers for combined S& and particulates control. While costly and requiring on-going maintenance, flue gas scrubbers do remove over 80% of the fine catalyst dust that leaves the regenerator. By far the most common and cost effective means of removing the catalyst from the flue gas is to again put an ESP in the system. With the finer dust and lower loadings the ESPs are still over 80% efficient. The key point to note here is that the ESP’s returned to the FCC flue gas system for environmental reasons, not for economic reasons. With either flue gas scrubbers or ESP’s, the case study unit which is now charging 60 m b/d is emitting 20 to 40 lbs per hour of particulates, or 0.5 to 1 lb per lo00 coke bum, and stack opacity is 10 to 20%. So, the 1990 state-of-the-art FCC releases 90% less particulates than in 1950 even though )feed rate has been increased 3 fold. The majority of this improvement came as a result of technical developments in cyclone design and catalyst manufacturing, and were economically driven. This is because recovery and reuse of the’catalyst, which cost about $500/ton in 1950, was essential to the overall process economics. Only the last increment was environmentally driven that put pollution control devices on the system.

3.2 Carbon Monoxide In the 1950’s carbon monoxide (CO) leaving the FCC stack was about 7 ~ 0 1 % . At

the low regenerator temperatures then practiced, carbon remaining on the catalyst after the regeneration step was often 0.5wt%. While the environmental impact of this high CO was recognized, it was also apparent from an energy efficiency point of view that a lot of useful energy was going up the stack. By 1970 most units had added a CO incineration boiler to the flue gas system. These gas fired boilers consumed almost all of the CO that came from the regeneration step. The energy was converted into high pressure steam, and these boilers became an integral part of the plant utilities. However, fired CO boilers were not without their problems. Because of the integration with the plant utilities, an FCC upset was invariably a plant wide utility upset. Also, having a fired boiler in-line with the process created some safety issues, such as the potential for an FCC to put hydrocarbon vapor in the flue gas going to the CO boiler. Over the time CO boilers were used in the FCC process there were no significant improvements in the design, they were

572

a significant maintenance cost, and presented potential safety problems for the process. Needless to say, most were happy to see CO boilers being removed from the process.

1971, when Amoco Oil demonstrated for the first time the complete combustion of CO in the vessel [4]. This major technological development earned Amoco Oil Company the prestigious Kirkpatrick Award from Chemical Engineering in 1975 [5]. This process assured the maximum utilization of coke in the process in terms of energy efficiency, and burned the carbon remaining on the regenerated catalyst to 0.05 wt% . This reduction of carbon improved the utilization of catalyst zeolite and, as a result, yield selectivity improved. So, full CO combustion in the regenerator vessel had a huge economic incentive, and also reduced CO emissions from the vessel to less than 500 ppm. This meant that the fired CO boilers were no longer needed, and could be replaced by sensible heat recovery units. Thus, the state-of-the-art in 1990 for the case study FCC as shown in Figure 5 was full CO combustion in the regenerator with a waste heat recovery unit. The technical developments driven by economic, environmental, and safety considerations have resulted in a process charging 3 times its initial feed rate, yet CO emissions have dropped 99% from 17,500 lb/hr in 1950 to less than 200 lb/hr in 1990.

By far the most significant change in the regenerator side of the process occurred in

3.3 Sulfur Oxides

the aromatic sulfur in the feed condensing with the coke on to the catalyst. The amount of sulfur in coke on the spent catalyst can vary depending upon feed type, process conditions and catalyst properties. For typical feed sulfur and operating conditions in 1950, the case study FCC would have emitted 200 to 500 lbs/hr of SC&, or 500 to 1500 ppm in the flue gas. This was little changed by 1970, except that as the feed rate and coke burn doubled so did the SO,. The concentration remained the same. The development of full CO burn in the FCC regenerator was the first step toward a reduction in sulfur oxide emissions from an FCC. The more severe regeneration and the fully oxidative conditions in the vessel along with the excess oxygen gave the FCC catalysts, particularly the higher alumina catalysts, a chance to adsorb some of the Sq( as a sulfate. The result was a reduction in SO, emissions by about 20%. This effect was first recognized by Amoco Oil and was an important first step in the development of S& control catalyst. All that remained was development of more active catalyst or additives for the removal of SQ. Also, in the 1970’s and 1980’s more high sulfur crudes were processed as the US supply of light crudes declined. A number of feed hydrotreaters were built at this time to upgrade the poor quality gas oils from the high sulfur crudes. This later move was significant enough to result in a reduction in overall FCC feed sulfur. The net result of both full CO bum and the feed hydrotreaters that were built was a reduction in FCC stack SO, emissions without special control technology to 300 to 900 ppm. When the threefold increase in throughput is considered, the SO, emissions from the case study unit have slightly increased to 400 to lo00 lb/hr in 1990, or about double the 1950 rate.

FCC have not been matched by reductions in SQ, has brought SO, to its present point as the key targeted pollutant. Driven by Acid Rain and other Clean Air legislation, development of a cost effective means of SO, reduction has been major area of research for the last 15 years. It should be noted that much of this effort has been the direct result

Sulfur oxide (SOJ emissions from an FCC regenerator are the result of some of

The fact that the spectacular reductions in particulates and CO emissions from an

573

of pressure placed on California refiners by the California Air Resources Board. By contrast most of the US refining industry have the FCC SO, emissions "Grandfathered" at about 1970 levels. Therefore, the refining industry in California has taken the lead on SO, reduction. Among the options investigated have been severe FCC feed hydrotreating, flue gas scrubbing, and SO, control additives.

Severe feed hydrotreating to low sulfur levels could reduce SO, emissions. However, SO, emissions are not linearly related to feed sulfur, and there is a diminishing return. As a result 90% feed desulfurization typically only reduces stack SO, emissions by about 75%. Excessive feed hydrotreating beyond a certain point results in a feed that makes so little coke that the process heat balance can no longer be satisfied. As a result the feed cannot be processed with the existing FCC process. Finally, FCC feed hydrotreating units are high pressure hydrogen units, and are expensive.

Wet scrubbing is one of the few scrubbing technologies to have been tried on an FCC. The equipment needed to contact and react the sulfur oxides out of a large volume stream like FCC flue gas is significant, and represents a substantial capital investment. The on-going chemical costs, maintenance costs, and the handling costs for the waste stream all contribute to the high cost of this technology. Other scrubbing technologies could be considered for the FCC process, but few have been tried. Perhaps one of the more interesting possibilities is the NATEC dry scrubbing approach [6]. While this would be less capital intensive, the chemical and waste handling costs still are present. The advantage that all the scrubbing technologies should have is the ability to reach 90% SO, reduction. The main concerns are the costs and onstream reliability. Achieving the latter may require redundant equipment, which will further increase the cost.

catalyst technology. This approach to reducing SO, was outlined in the 1970's and numerous early formulations were tried [7]. The reaction chemistry is straightforward: a metal oxide in the regenerator would absorb the SO, as a sulfate on the catalyst. The catalyst would then move the sulfate to the reducing conditions of the reactor, where the sulfur would be released as hydrogen sulfide. A downstream sulfur plant would then collect the sulfur in the elemental form. Like most catalyst developments, the evolution of the technology and the problems to be solved were certainly challenging. It now appears that in situ SO, control additives are available that are able to reduce SO, in many cases by 80% at a cost of less than 1/2 of wet scrubbing. In terms of cost effectiveness, $/lb SO, removed, the additive approach seems to be the preferred route.

For all the work that has been invested in developing control technologies for FCC flue gas SO, emissions, the fact is that only a small, albeit growing, number of FCC units use any Sq( control. There are perhaps 5 units in the US using wet scrubbers, and maybe 15 or so units using additives to reduce SO,. Unlike particulates and CO, there is no economic driving force to reduce SO,. While the EPA has promulgated new source performance standards (NSPS) for FCC SO, emissions, few units have to meet the 300 ppm maximum SO, requirement. Most FCC's in 1990 were "Grandfathered," and without any control technology are emitting 300-600 ppm SO,.

The most promising technology at this time is an in situ control of the SO, using

3.4 Nitrogen Oxides

low, on the order of 50-150 ppm. Low feed nitrogen content and the low severity The emissions of nitrogen oxides (NO,) from the original FCC units in 1950 were

574

regeneration were considered the reasons for these low emissions. With CO boilers in the 1970's some interesting things about the chemistry of NOx became apparent. While the emissions of NO, out of the regenerator remained low, at high firebox temperatures (1600°F) with excess oxygen in the flame, high NO, emissions out of the CO boiler were possible. At times the emissions would reach 400 ppm and the plume would become visible due to the NO,. Since the advent of full CO bum, high NO, emissions have at times occurred as the result of local hot spots in a vessel due to mechanical damage, or when excessive CO Oxidation Promoter has been used. The resultant understanding of NO, formation in the FCC process implies that particle temperatures of 1550°F or higher are required for nitrogen fixation. This means that control of NO, at low levels, ie 50-150 ppm, is achieved by proper control of the regenerator. Even temperature distributions, and avoiding "excessive" oxidizing conditions will be the key. This implies that keeping the level of NO, emissions low will be related to maintenance practices in the regenerator. So, for most FCC units in 1990, NOx emissions remained low at about 50-150 ppm.

4. HISTORY OF ENVIRONMENTAL IMPACT OF FCC PRODUCTS

In comparison to stack emissions, the impact of FCC product quality on the environment is less obvious, and not nearly as well understood. Indeed, this is an area of current research for all refined products. One reason refiners seem less in tune with this area is that these fuels are consumed beyond our facility limits, oftentime weeks after their production. Yet, the quality of our fuel products is one way in which we present our industry to society as a whole. With the increasing emphasis on environmental issues, it is clear that we must modify refining processes to make them cleaner burning. Over the period 1950 to 1990 the yields of products from the FCC process had changed dramatically, as the data in Table I indicate.

Table I FCC Yield History

1950 Fuel Gas, Wt. % 8 cg 'S, Vol % c4 'S, Vol % 8 Naphtha Vol % 40 LCO, Vol % 33 DCO, Vol % 7 Conversion, Vol % 60

- 1970 1990 3.5 3.5

9 12 13 16 50 56 28 20 5 4

67 76

We will now review the environmental impact of the FCC products, and see that those changes are as significant as the yield changes.

575

4.1 Fuel Gas Fuel gas production from an FCC is sometimes not thought of as a product per se,

because it is consumed within the refinery. However, because it is a fuel it does have an environmental impact. Looking back at the case study unit in 1950, we find that the fuel gas consisted of all cracked products lighter that butanes. So, even though the process seventy was low by current standards, the net fuel gas make in 1950 was about 8 wt% of feed. Moreover, its sulfur content was about 5 to 10 wt% because of the hydrogen sulfide. So, when this fuel was burned in the plant, the resulting flue gas had an SO, concentration of 3000 to 6OOO ppm. This high level probably contributed to high furnace maintenance. By 1970 this situation had changed dramatically. In most gas concentration units an additional tower had been added to recover the propane and propylene from the FCC products for use as LPG and chemical plant feedstocks. Perhaps even more important from an environmental standpoint, the fuel gas from the absorber was then directed to a fuel gas scrubbing unit, for example an amine scrubber using monoethanolamine. This unit removed hydrogen sulfide from the fuel gas, and directed it to a sulfur recovery unit. These two changes reduced the amount of fuel gas and greatly reduced environmentally harmful SO, emissions when the gas was burned. By 1990, process seventy, including riser outlet temperatures, had been increased. However, the designs of the units emphasized short contact times to minimize low valued fuel gas make. Also some gas concentration units added refrigeration units to enhance propylene recovery. As a result, the typical fuel gas yield in 1990 was about 3 to 4 wt % of feed, and, as mentioned before, is scrubbed to remove sulfur. So, the fuel gas that an FCC made in 1990 should be considered an environmentally acceptable fuel, while the fuel gas in 1950 was very high in sulfur.

4.2 Fuel Oil/Decanted Oil Decanted oil, as the bottom product of an FCC fractionator, is something to be

minimized. The history of decanted oil (DCO) yield and quality in some ways parallels that of fuel gas. In 1950 a typical yield of DCO was about 7 vol% with a gravity of 17" API. Sulfur content was on the order of 2.5 wt%. Because of the state of cyclone technology at time, even with 3 stages of reactor cyclones, the catalyst losses were high. So, most units had external slurry settlers to decant this oil. Hence the oil's name of Decanted or Clarified Oil. These settlers collected about 80% of the catalyst lost by the reactor cyclones and returned it to the process via Slurry recycle. Even with this system, the ash content of typical decanted oil was about 0.15 lb/gal. The net loss of catalyst on this side of the process for the case study 20 mb/d unit in 1950 was about 3 to 4 tons per day. By 1970 increased process seventy had reduced DCO yield to 5 vol%, and the gravity had dropped to about 7" API. Sulfur content was still about 2.5 wt%, and with 2 stages of reactor cyclones and a slurry settler the ash content was about 0.1 lb/gal. By 1990 further improvements in design and catalyst had reduced DCO yield :to,about 4 vol% with a 2" API gravity. The harder catalysts that are being used allow the ash content to be kept at or under 0.1 lb/gal ahead of the slurry settler. Removal of slurry settlers from the FCC process has been something that refiners have strived toward for many years. That is because they are a maintenance headache, and are a potential safety problem. With the combined improvements in catalyst and cyclone technology, settler removal is now possible. So, the result of 40 years of FCC process development on DCO has been a

576

reduction in almost half of the yield of this low valued and poor quality fuel. While its sulfur content is largely unchanged, it is now more aromatic but contains less catalyst dust. From an environmental perspective, the changes in quality are probably balancing, but the reduction in volume is a big environmental plus.

4.3 Light Cycle Oil Light Cycle Oil (LCO) is the distillate boiling range product of an FCC. It is

generally lower in gravity and higher in sulfur than most other distillate streams. As such most is blended off, and some will be desulfurized to 0.3 wt% sulfur. From 1950 to 1990, as unit severity and improved designs have increased conversion, the yield of LCO has dropped from about 35 vol% to about 20 vol%, and the gravity has also dropped slightly from 30" API to about 26" API. The sulfur content has dropped slightly from about 0.9 wt% to about 0.7 wt%, primarily due to about 30% of FCC feed now being hydrotreated. Indeed, even if all FCC feed were severely hydrotreated, LCO sulfur would be between 0.3 and 0.4 wt%. So, on the whole, the quality of LCO as a fuel, and hence its environmental impact, has not changed significantly. Only severe hydrotreating of the FCC LCO product will improve its fuel quality. The only thing that changed from 1950 to 1990 was the yield of this non-premium distillate blending component was reduced by about 40 % .

4.4 FCC Naphtha and LPG Olefins Affect the Gasoline Pool

will have a major impact on the process and also the environment. The upgrading of gas oil to gasoline blending components by the FCC not only almost doubled the amount of gasoline that could be produced from a refinery, but it had profound impact on the quality. Without an FCC, motor gasoline would be essentially virgin naphtha and reformate. As the FCC process started to develop in 1950 the yield of FCC naphtha was about 40 ~ 0 1 % . With the technology of the process and the catalyst of the time, this naphtha had about a 0.15 % sulfur level. The octane was fairly good at about 92 RON and 79 MON, primarily because it was high in olefin content. Table 2 contains typical data on the evolution of the composition of FCC naphtha over the last 40 years.

FCC naphtha, or gasoline, is the primary product of an FCC and as such its quality

Table 2 FCC Naphtha Composition

1950 Paraffins, Vol% 43 Naphthenes, Vol% 8 Aromatics, Vol % 25 Olefins, Vol% 34 Sulfur, PPM 1500 RON 92.5 MON 79.5

1970 50 13 30 17

lo00 89 78

1990 50 13 27 25

500 92

80.5

577

By 1970 the advent of zeolite catalyst and increased process severity had increased the yield to about 50 ~01%. The gasoline olefins had dropped significantly due to the zeolite catalyst, and the sulfur level was also lower. Isoparaffins were preserved as a result of the catalyst shifts, and aromatics were somewhat higher, a result of zeolitic hydrogen transfer. On the whole the naphtha produced by the FCC process in 1970 was an improvement from an environmental point of view, ie, lower sulfur, more isoparaffins, etc. On the otherhand, octane had dropped to 89 RON and 78 MON. The point to emphasize, however, is that the changes were made for economic reasons, not environmental reasons.

During the 1980’s the emphasis on environmental concerns began to take shape. One of the first big targets of this growing concern for the US refining industry was the lead that was used to enhance the octane of gasoline. In order to meet the challenge of Lead Phase Down, the octane that it provided had to be replaced. The response by the FCC was to increase the process seventy even further and to use catalyst containing Ultra Stable Y Zeolite to enhance the octane. So, by 1990 the average octane of FCC naphtha was up to 92 RON and 80.5 MON, from the low levels of the 1970’s. The combination of severe operations and USY catalyst coupled with the designs to take advantage of them increased the net production of branched hydrocarbons in the gasoline boiling range, and reduced sulfur level to about 0.05 wt% Another thing that this operating mode did was significantly increased the yield of butylenes. Indeed by 1990 the yield of total C4’s was about 16 vo l l , or just about double the yield of the unit in 1950. This stream, when charged to a downstream alkylation unit, will produce a high octane and an environmentally good quality gasoline stock. When taken as a whole, the changes that have occurred in the FCC naphtha, and the increased yield of alkylate have had a positive effect on the environment. The early changes were not as dramatic, and were made for economic reasons. Only the later changes were made for environmental reasons, primarily in support of Lead Phase Down.

4.5 Waste Issues Waste streams from a refining process should be considered products, albeit that

they are undesirable products. Waste generation, management, and proper disposal are appropriate environmental issues in their own right. Indeed, there is a growing body of legislation in the area of controlling waste and contamination of water resources. The refining industry as a whole has a major challenge and responsibility in the 1990’s. The FCC process will have to respond to those items as well. For the most part the waste issues for an FCC are the same as most refining processes, and will not be covered here. The one unique item for the FCC is the generation of a potential waste in spent cracking catalyst. While spent cracking catalyst is primarily a silica-alumina clay-like material, two issues contribute to concern about it and it’s disposal. First it is a dusty material, and some of it may be less than 10 microns, raising possible respiratory issues. Secondly, heavy metals such as nickel will be on the catalyst. While these are usually at very low levels compared to other refinery catalyst, or wastes from other industries for that matter, the possibility of these metals leaching from the catalyst when contacted with acidic water @H 3 or less) does exist. If the heavy metals do leach off, then it is likely that they will reach the drinking water supply. In general FCC catalyst will pass Federal leaching tests, (TCLP), indicating the metals are fairly tightly bonded to the catalyst.

578

In the 1950’s the relatively poor cyclone technology resulted in high losses from the FCC process to the flue gas stack and/or the decanted oil. As a result there was very little spent catalyst to be disposed. As both cyclone technology and the technology to harden the catalyst were developed, the need to withdraw some spent catalyst started to occur by 1970. This catalyst was typically sent to landfills. By 1990 the catalyst were even harder, and ESP’s have been added to get the fine dust out of the flue gas stack. Also more resid has been charged to FCC’s, which has increased catalyst usage rates. As a result, a substantial portion of FCC catalyst now must be disposed of either as spent equilibrium catalyst with an 80 micron average particle size, or ESP fines which are under 40 microns. The heavy metals on this waste catalyst are low, generally under 1500 ppm, and they are not easily leached off the catalyst. As such, FCC spent catalyst is not usually considered a hazardous waste. Most of the spent catalyst in the 1990’s is being sent to landfills. Alternate disposal options are also under investigation, eg. using spent catalyst as a raw material supplying alumina to cement plants. Demetallation technologies exist, but have not been widely implemented [8].

5. SUMMARY OF 1950-1990 FCC ENVIRONMENTAL IMPACT

In reviewing the history of the FCC process and it’s impact on the environment one comes away with a feeling of substantial accomplishments that have been made on the part of the people in the refining industry. New designs, catalysts, process options and technologies have been developed and put in service. Challenging problems have been solved, and the unique flexibility of the FCC process has allowed these to be implemented. As a result, the FCC process has made steady progress on environmental issues, whether they are driven by economic reasons or legislative action. The FCC process in the 1990’s is a cleaner more environmentally conscious process than ever before. Moreover, the products from the FCC process are cleaner burning and more environmentally acceptable than ever before. To turn a phrase from the 1970’s around a bit, ..the FCC process may have been part of the problem, but it is now clearly part of the solution.

6. FUTURE FCC ENVIRONMENTAL CHANGES 1992-2001

In the preceding section the history of the FCC process has been woven together with the environmental impact of the process. Some improvements were direct, and others indirect. Most early improvements were economically driven, and only in the last ten years have the environmental improvements been driven by legislative action and regulatory initiative. This in and of itself implies that the early history was one of tuning to optimize the process. A good example was full combustion of CO in the regenerator, which had both an economic and environmental benefit. As the process developed, issues that had both economic and environmental impact became fewer. As another example, the refiner is hard pressed to identify the economic benefit to be derived from reducing SO, in the FCC flue gas. Yet, the environmental benefit to society is clearly there. The dilemma for the refiner is to balance the technology available to improve the environment, the extent to which the environment is improved, the cost of that technology, and the cost of alternate technologies. A number of issues come into play, eg. the cost and benefits for any action to improve the environment, the inherent safety of the technology, and the

579

impact versus alternatives on the overall economy, etc. For each environmental question of the 1990's the complex issues and consequences will be different and unique. The US refining industry through it's various organizations, API and NPRA, must be part of the on-going dialogue that our society will go through as we wrestle with these tough issues. We must become part of the process that establishes the targets and specifications of future environmental regulation. A good example of this was the "Regulatory Negotiation Process" used in developing the guidelines for Reformulated Gasoline [9]. It is in this type of activity that environmentally and economically sound solutions to these profound problems will emerge. This objective can not be accomplished with out our deep and committed involvement.

With regard to the FCC unit in particular, it will be essential that the previous history of running cleaner and producing cleaner fuels be maintained. We anticipate that increasing environmental regulation of the process is likely. We anticipate that the proven flexibility of the FCC process will again be tested as the process design, operation, and catalyst technology respond to these environmental demands. In the remainder of this section some of the key environmental issues impacting on the FCC process will be addressed. This discussion is offered as part of the dialogue on these important issues for not only the refining industry, but for environmental organizations and regulatory agencies.

6.1 Flue Gas Stack

now for most units under the NSPS standard of 1 lb/1000 lb of coke bum. Improvements in cyclone, 3rd stage separator or ESP technologies are not likely to reduce this number very much. Even wet scrubbing technologies do not do much better than NSPS. Stack testing on units with wet scrubbers generally have stack losses of particulates of 0.3 to 0.9 lb/IOOO lb of coke bum [lo]. So, at the present there is no demonstrated best available control technology (BACT) to justify particulate targets below current NSPS. On the other hand, much of the dust that does go to the stack is of the 10 micron and less particle size. This fine particle size may be of concern even if it is dispersed over a large area from a high flue gas stack. If so, the only way to make any reduction in this is to employ some sort of direct filtration of the flue gas, eg. a bag house type of system. Such a system would be expensive and complex. It would first have to have some sort of surge capacity like a Shell 3rd Stage Separator to handle upsets. It would also involve a large bank of multiple filter elements cycling between filtering large volumes of hot gas, and blowing the dust back to collection. Disposal of this very fine dust is also a problem. Another concern from a process point of view is that such a system will back-pressure the regenerator, creating safety issues. Indeed, a concern that needs to be addressed in all of these environmental questions is whether we are pushing ourselves against the limit of technologically safe operation. At the very minimum this will add instrumentation and control to an already complex process, and will reduce the reliability of the process. With serious technical development efforts a system could probably be developed that could make a 50 to 70% reduction in these fine particulate emissions and be on line on FCC toward the end of the decade. As such we feel that in the 1996 timeframe the current NSPS particulate standard for FCC may be reduced to about 1 lb/3000 lb of coke bum, and units will start implementing such a technology around 2000.

As previously discussed, particulate emissions of FCC catalyst dust in the stack is

580

Of the gaseous flue gas emissions; CO, NO,, and SO,, it seems fairly certain that only Sq( is likely to come under any Federal Regulatory pressure. Carbon monoxide, as previously discussed, has been reduced over the years, and is now well under 500 ppm. Nitrogen oxides are a factor in smog formation. However, the NO, levels in FCC flue gas are generally less than 200 ppm, which is substantially below levels for other processes, eg. coal combustion. For this reason we feel that few FCC units will be limited by NO,. We would expect analyzers will be required for all of these components in the FCC flue gas on virtually every unit by 1995. As previously pointed out, NSPS for SO, on FCC units is about 300 ppm. However, only a small number of units now fall under NSPS, and most units are emitting 300 to 600 ppm, with some higher. It is likely, and clearly it is the intent of the various agencies, that most FCC units will be brought under SO, NSPS during the next decade. Since very few units would regularly be under 300 ppm without control technology, most units will have to start planning on capital and expense spending simply to reduce FCC flue gas SO,. Some FCC’s, primarily units without ESP’s, will select wet scrubbing to achieve both particulate and SO, reduction. This will be a costly and maintenance intensive route. One risk that it has is that it might not meet substantially tighter particulate emissions that are likely late in the decade, and the waste disposal problems with wet scrubbing might become prohibitive. These concerns will foster the development of improved dry scrubbing technologies just for S4, that may be commercially available around 1997. Considering all these issues, it is our belief that most FCC operating companies will select the SO, control technology carefully. The optimum solution may well be the technology that generates the highest pollution reduction per unit of cost to the industry, and hence society. Indeed, the effectiveness of alternate technologies can be ranked in this manner. Based on currently available information, the SO, reduction catalyst technology would seem to be the most effective means of reducing SO, by 50 to 80% for most FCC’s. If this is the case, a substantial increase in SO, control catalysts will be needed over the next decade.

6.2 Regulatory Effects on FCC Products This is perhaps the area that will have the most impact on the FCC process. Yet,

as we have seen, it is only is the last few years that this subject has been addressed in legislation or regulatory action. As such, the impact and how these issues will shake out and eventually be translated into how the processes are run is not at all clear. What is clear is that the Clean Air Act of 1990 calls for the use of Reformulated Fuels [ll]. Refining industry groups are now working with the EPA in developing the details of how this will be done. The other interesting development is the AUTO/OIL AIR QUALITY IMPROVEMENT RESEARCH PROGRAM which is evaluating the effects of gasoline composition for various engine types on air quality. The data from this program will be essential in guiding the refining industry and the regulatory agencies in developing the criteria for Reformulated Gasoline.

some level of oxygenates will be required in all Reformulated Gasoline. The leading contender to provide most of the oxygen now appears to be MTBE. To this end, the production of isobutylene will be a key refining objective. Also, for both gasoline and distillate products, there are likely to be certain compositional objectives set by forthcoming regulations. For example, gasoline is likely to have a maximum benzene

Turning attention to issues which may impact on FCC operations, it now seems that

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content. Also, broad compound types, for example total sulfur in highway diesel, will be set. For certain items new definitions of compounds are being developed, eg. POM for Poly Organic Material. What will be needed are analytical techniques to measure this growing volume of compositional information, an understanding of how process operations will change the composition, and technology to optimize the production of components that achieve the objectives of the Clean Air Act of 1990.

6.3 FCC Naphtha 1992 to 2001 What will become of increasing importance for FCC naphtha in the next decade

will be its composition. This will of necessity generate a whole new set of terminology for the process. So, "Octane-bbls" will be replaced in importance by "Aromatic-bbls", "Benzene-bbls", "Olefin-bbls", etc. Indeed, one of the first questions that will need to be answered is how to measure some of these components. The objective then will be to optimize the composition within the context of Reformulated Gasoline. To do this will require a fundamental re-examination of the process chemistry. The direction in which the fuel composition is likely to move will be guided by data on the impact of that composition on tailpipe emissions, such as the data now coming from the Auto/Oil program. Potential compositional variables that are affected by FCC operations are aromatics and sulfur. Catalytic reduction of aromatics and sulfur should be possible in the reaction process. The catalyst that should directionally reduce both of these in FCC naphtha would have an active alumina matrix and a zeolite with a lower hydrogen transfer rate. The process conditions that would be favorable would include higher riser temperatures and shorter residence times. The intent of these moves would be to reduce ring formation and dehydrogenation to aromatic compounds, including aromatic sulfur compounds. If the olefins in gasoline are regulated, the conditions indicated here would probably result in olefin increases. Since most of the olefins are in the front end of FCC Naphtha, they could either be selectively hydrotreated, or converted to ethers. In any case, it appears the likely direction for FCC naphtha will be to reduce aromatics and sulfur, and deal with the olefins downstream, if necessary. The emphasis, however, will not be so much on the concentration of any species, but rather the net barrel yield. For example, it is anticipated that the yield of FCC naphtha will actually drop in a Reformulated Gasoline refinery. Even if the concentration of aromatics remains constant, the net barrels of aromatics to the pool will drop. Regardless of whatever might be done catalytically to improve the composition of FCC naphtha in terms of environmental impact, it is possible that late in the period there will be increased regulatory pressure on the end point of the pool and hence on FCC naphtha. This may require either novel process solutions to reduce the yield of the last 40°F boiling range, or hardware modifications to separate it out. In the later case, the challenge would be to figure out what to do with this material.

6.4 Isobutylene

example as alkylation unit feed. But the impact of the Clean Air Act on requiring oxygenated fuel will make isobutylene far and away the most valuable component. Isobutylene is a principal product of the FCC cracking reactions. However, because it is a reactive structure, there seems little doubt that some of it is destroyed in the FCC via

The C4 stream from an FCC vapor recovery unit has always been valuable, for

582

hydrogen transfer, polymerization, or thermal reactions. Recently, published data by Matos as shown in Figure 7 clearly indicates this effect [12]. So the FCC process will need to be tuned even more toward the selective production of isobutylene, and catalytic and process considerations will need to emphasize the preservation of isobutylene once formed.

6.5 Light Cycle Oil The pressure from sulfur specification on Highway Diesel will indirectly affect

FCC LCO. Additionally, the desire to minimize aromatics to improve burning characteristics will have an impact. The net result is likely to be a depression of LCO value as a distillate blending stock based upon it’s current quality. Historically, the quality of LCO has never driven the FCC process, and with the impending pressure on Reformulated Gasoline, and the drive to make isobutylene, this is not likely to change. So, more than likely refiners will accept what quality of LCO comes with process and catalyst changes to optimize the unit for Reformulated Gasoline. Based on the proceeding discussion, these changes will be potentially significant. However, their impact on FCC LCO quality is likely to be of only a secondary importance. FCC feed hydrotreating will mean a slightly less aromatic and lower sulfur LCO from the FCC process, emphasizing the point that feed pretreatment is not the solution. All these items will not offset the depression in value of FCC LCO because of pressure on the distillate pool aromatics and sulfur. As such, the spread between FCC cracked products (isobutylene and FCC naphtha) and LCO will increase by a few cents a gallon. Moreover the spread between LCO and premium distillate products will also increase. So, the likely outcome is that a much higher percentage of FCC LCO will be desulfurized by the mid 1990’s, and by the end of the decade perhaps most will be desulfurized.

6.6 Decanted Oil

reach the point where this high sulfur, high ash fuel will be not be an acceptable fuel on its own. Since any substantial reductions in sulfur or catalyst fines in DCO are not likely, the only outlet for this bottoms product will be as fuel oil to boilers with Flue Gas Scrubbers or as a Bunker C type fuel. This will depress DCO value to that of coal, and mean that minimizing DCO yield will remain a primary objective of FCC operating strategy.

The environmental impact on Decanted Oil (DCO) is likely to continue and will

6.7 Waste Disposal Issues

refining processes. While we will not discuss these in any detail here, we need to keep in focus the impact of them on our daily business. Specifically, while trying to achieve all of the other refining and environmental objectives, it will be necessary to reduce all waste streams by 50 to 90% during the Environmental Decade of the 90’s. The waste stream that is unique to the FCC is spent cracking catalyst. While there will be pressure to minimize the use of FCC catalyst so as to reduce waste, the huge economic leverage of FCC catalyst to effect improvements in yields, including significant enhancement of the environmental quality, will preclude any large reduction in FCC catalyst usage. However, it is certainly likely that disposal of FCC catalyst will be regulated sometime in the next

As previously discussed, most FCC waste disposal issues are similar to other

583

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584

decade. That regulation will probably take the form of some maximum level of metals on the spent catalyst which can be put to landfills. Ideally that level would be set based upon standardized leaching tests of the various metals, including antimony, if present. In any case, whatever the initial regulation in terms of maximum metals, it will probably tightened over time toward the objective of eliminating sending FCC spent catalyst to landfills. Given this, and the previously discussed economic leverage of FCC catalyst, it seems that either a recycling approach, eg. DEMET, or an alternate use of spent FCC catalyst such as in cement will be required.

7. SUMMARY

In this overall review of the environmental considerations involved in the FCC process, it is clear that much has changed over the first 50 years of the history of this important process. Substantial, and in some cases phenomenal, reductions have been made in the emissions of air pollutants. Most of these changes, particularly earlier on, were made primarily for economic reasons. Furthermore, improvements in the environmental quality of the fuels produced by the FCC have also been made. Only recently have regulatory actions affected the FCC process in a significant way. Perhaps the most significant of these was the impact of Lead Phasedown in the early 1980's. Unfortunately, it appears that the opportunities for technical improvements that are economically driven that also have a positive environmental impact are becoming fewer. At the same time our society, both domestic and global, is calling for a cleaner environment. So, this environmental call will become apparent in the various legislative initiatives at both Federal and State levels. As a partner in this process, the refining industry needs to responsibly input on the options and potential impact. Then it will be necessary to work with the various agencies toward meaningful implementation. We have outlined in the final section of this paper what we see this possibly meaning for the FCC process in terms of stack emissions, fuel quality, and spent catalyst waste. We have tried to present a balanced view, and not necessarily the one that would represent the lowest cost or least burden to the operator of an FCC unit. Yet, meeting what has been outlined here will be extremely challenging. We do this based upon the fact that the FCC process has demonstrated in the past a remarkable flexibility to meet challenges with process, design, and catalyst improvements. It is likely that a number of ways of making equivalent improvements to the environment through the FCC process will be found. The key to success will be identifying the least costly and getting them incorporated into proposed regulations. We look forward to the challenge of having the FCC process operate in an even more environmentally responsive manner, making even cleaner burning products, generating less waste, and doing all of this in a cost effective way.

REFERENCES

1

2

"The Fuel Revolution," G . H. Unzelman, Fuel Reformulation, Volume I, No. 1, p. 32, September 1991. "Consumption of Energy 1988, " Department of Energy, Energy Information Administration, Report 0512(89), Category UC-88, May 1991.

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7

8

9 10

11

12

"First Fluid Catalyst Cracking Unit Starts Operating," Oil & Gas Journal, 41, No. 5 , June 11. 1942. Horecky, C. J., Fahrig, R. J., Shields, R. J., and McKinney, C. 0, U.S. Patent 3, 909, 392. "Complete Combustion of CO in Cracking Process, " Chemical Engineering, November 24, 1975, p. 46. "Injection of Dry Sodium Bicarbonate to Trim Sulfur Dioxide Emissions," T. Coughlin, P. Schumacher, D. Andrew, R. Hoopep, EPRUEPA SOz Control Symposium, May 8, 1990. "Amoco's New Ultracat Process for SO, Control," I. A. Vasacos, E. R. Strong, C. K. Hsieh, C. J. D'Souza; API 42nd Mid Year Meeting, Paper No. 20-77, May 10, 1977. Elvin, F. J., and Pave], S. K., "Fluid Cracking Catalyst Demetallization - Commercial Results," Paper AM-91-40, 1991 NPRA Annual Meeting, March 17, 1991. "The Clean Air Act and the Refining Industry," UOP Report, September 1, 1991. "Fluid Catalytic Cracking Unit Flue Gas Scrubbing," Exxon Research and Engineering, Exxon Technology Report 250-55-JDC, May 1985. "Clean Air Act History Marked by Battle, Compromise," F. L. Potter, Fuel Reformulation, Vol. I, No. 1, p. 22, September 1991. "Increased Butylene Yield in the FCC," J. A. Matos, Davison Catalagram, No. 81, 1990.

587

SUBJECT INDEX

A

Acid sites 42,43,45-50,53-6 1,64,65,68,70-72,149,l53,155,167,173,174,175,201-

Active sites in zeolites 106,156 216,299,508,511,524

accessibility 5132,166,169,173 Bronsted acid site formation 45-47,153 characterization with probe molecules 50,53,56-61,201-205,213-215 direct measurement 53-58 formation and transformation 43,45,50 historical 42-44 indirect measurement 53,59,60 Lewis acid site formation 47,48,153 next nearest neighbors 49,61,70,153,154,155,207-209 relationships between Bronsted and Lewis acid sites 47,48 site geometry 49,64,65 ZSM-5 49,56-58,61,64,71,91 ,185-1 87,196,202,204,210,213-214,501

Additives 146,147,149,170,171,175,258,310,531-56 octane boosting 171 oxidation of CO to CO, in regenerators 303,532,553,555-558

catalysts for 29,149,172,403,405,556-557 thermodynamics of reactions 400,403,552,555

removal of NO, from regenerators 532,551-555 catalysts for 552-555 chemistry of 551-552,554-555

catalyst analysis 538-546 catalyst evaluation 534-537,543446,547 catalysts 534-536,543-544,547350-551372,580 chemistry and mechanisms of 29-30,215332-533,555 commercial trials of 549-550 spinels 171,536,538,543,546-549

removal of SO, from regenerators 147,149,531-551,555

resid cracking 310 simultaneous removal of NO, and SO, from regenerators 553-555

Adsorption 56-58,191-195,212,215-216 After burning 403,405 Alkali metals 42,175,483 Alkyl thiols 372 A1 k y la tion 7-9,68,176,178,26 1,499,504,5 153 16,5 19,527,577 Alumina 116,117,119,121,124,146,147,167,371,374,535,539-541,547,556 Aluminophosphates (ALPO's) 86,524 Aluminum 153,l 54,155,156,157,167,371,374,535,539-541,547,556 Aluminum chlorhydrol 118,119,123

structure of 119

588

Aluminum sulfate 107,108,109 Ammonium fluosilicate 110,113,114,115,116 Ammonium ion 42,4547,128,129 Amylene 178 Aniline point 161,464,465,471,493 Antimony 146,147,171,265,339,347-358,584

colloidal pentoxide 347 commercial tests of 350-357,483 lay down efficiency of 347,355,362 MP-25 348 MI’-85 349 Ni-Sb alloy 349 Phil-Ad CA 347,350 trithallate 347

nickel interaction 349 tin interaction 369-370

trisdiproplydithiophosphate 347 water-based 348,370 with CO promoter 357

Arc0 FCCU pilot plant 223,270,503 Aromatics 43,59,66,70,158,159,160,163,166,177,1 78,244,441,443,445- 447’45 1,455,456,457,458,461,462,464,467,468,471,488,493,501,502,508,509,510,517~2 3 Arsenic 371 Asphaltenes 472,473,474,475,476 Attapulgite 376 Attrition resistance 153,167,175

B

Barium 266,371,377 Benzene in gasoline 177,178 Beta-scission 68-71 beta zeolite 41,87,89,97,165 Binders 116,117,166 Bismuth 146,149,171,265,347,359-363

commercial experience 359-363,483

laydown efficiency 362 toxicity 359

CPM-112m 359,361-362

Boron 371 Bottoms cracking 157,159,168,175,176 Bronsted acids 150,152,153,167,201-205,213-215 BSS (Breck Structure Six) 85,86,96 Butene 178,179,261,422,423,429,489,490,5W,517

589

C

Cadmium 371 Calcination 50,109,111,112,118,120,123,124,127,129,547-549 Carbenium ion 42, 50,67-71,505,507,510 Carbocations 153 Carbon dioxide 172,532,553,555-558 Carbon monoxide 372,532,553,555-558

effect of regenerator temperature on 404 oxidation of 1 72,303,482,532,553,555-558,572 promoter 147,172,357,403-405,482,556-557

effect on conversion 428,430 effect on cracking selectivity 431

Carbon-on-regenerated-catalyst 229,389,428,487,556-558,571,572

Carbonium ion 42,50,71 Catalyst circulation

factors effecting 425 effect on conversion 426

Catalyst coolers 418-420 Catalyst entrainment 418 Catalyst/oil ratio 225,237,247,259,420,426,478,487,488 Cation sites 42,45,127 Characterization of cracking catalysts 100,223

bulk measurements 133,134 EXAFS, extended X-ray absorption fine structure spectroscopy 137,191 helium absorption and skeletal density 194

mercury intrusion 136,193 neutron diffraction 190 nitrogen adsorption 192-193

ISO-9000 140

micropore / mesopore determination 192- 193 pore size distributions by 136,192-193 surface area by 136,192-193 T-plot method and 136,192-193

organic absorption and kinetic diameter 194-195 pore size distributions by low pressure adsorption 194 powder X-ray diffraction 100

high resolution 189 identification of crystalline materials by 186-188,538-546 particle size analysis by 188 quantitative analysis by 188 small angle scattering 190 unit cell size measurement by 188-190

quality assurance 137,138 Statistical Process Control (SPC) 138,139

590

surface analysis 136,137,196-201 thermal analysis

DSC 215 DTA and phase changes 215 micro calorimetry 215 TGA and So, cycles, and zeolite dehydroxylation 215,534-537,547 TPD and acid site determination 213-214 TPR, TPO and nickel/vanadium contaminants 215

water absorption and pore structure 136,194 XANES, X-ray absorption near edge spectroscopy 191

Characterization of zeolites (see also zeolite) 100,186-190,192-195,197,201-215 Chinese crude 317,356 Chrysolite 376 Clarified slurry oil 575,582 Clays 43,83,109,110,116,120,123,124,125,146,147,148,149~166,167,171,187,527,551 Cloverite 83 Coke 145,151,155,159,161,162,168,171,174,176,177,~5,~,~7,~,~9,457,463,

464,466,467,468,471,474,487,488>02,516,519,522,523,524,527,532,555-558 burning 397-399,401-403,414 catalyst deactivation by 173,174 delta coke 240,400 effect of pressure on formation of 428 nature of 173,174 origins of 173 shot 377 yield, reduction by metals passivation 347-363,366-370,374,375,377

Combustor style regenerator 395,410-415 coke burning 414 combined with catalyst cooler 419 commercial experience 414 hydraulics 411 precombustion temperature 409 solids mixing in 412

Conradson carbon residue 37-38,473,474,487,488 Contact time 37,199-200,215,246,259,385,386,390,402,403 Contaminant metals 171,172,259,339-346,393,421,431,478,479,~0,~2,483/499,

Controlled combustion catalysts 172 Conversion of FCCU feed

522,523,584

effect of catalyst circulation on 426 effect of temperature on 422 effect on octane 425 effect on product olefinicity 423 effect on product selectivities 423

Copper 47J 72,483 Cracking catalysts 145

591

age, effect of 183,184,232,259,510,511 composition 3-5,116,147-149,175-177,183-184,446 deactivation 172- 175 demetallation 578 dual function (DFCC) 373 equilibrium 1 72,225,258,264,267,421,505,506,508-51 0,525,526 fines 355 formulation for metals tolerance 149,171 gasoline enhancing 148,149,150,151 improvements 4244,387,389 incorporation of zeolite 44,146,387,446 layered 378 matrix 116,166,167,169,175,176 metals passivation 339-380,389 metal resistant 265,266,267,522,523 octane enhancing 148-150,173 particle size and coke burning in moving bed units 15-17 poisoning 174,175 resid 149,150,151,165,172,173 stability 145

Cracking catalyst activity 50,226,238,435,436,454,485 Cracking catalyst design 175-177 Cracking catalyst evaluation and catalyst selection studies 225,238,258,259,260-

262,263,269,283-290,317,511,513,514,524 comparisons between MAT and circulating pilot units,

metals effects on 230,265,266,276,319 pilot plant operation for 269,280,322 pilot plant reproducibility in 276,279 pilot plants for 269,275

245,248,261,262,285,317

ARC0 circulating unit 229,270,503 Davison circulating unit 229,270,290 MAT 224,225,237,245,248,259

steam deactivation before 229,236,238,258,260,261,263,264,324,502,504,507 strategies 258,317,502,526

alumina in 116,117,119,122,124,167 binders in

Cracking catalyst preparation 5,105ff

aluminum chlorhydrol as 118,119,167 insitu binding as 120 peptized alumina as 119 silica as 117

calcination 1 19,120,123,124,127,128,129 clay and 116,120,123,124,125 commercial aspects of 105 drying 132

592

ion exchange 127 mixing and 125,126 pillared clays in 123,124 silica-alumina in 123 spray drying 126,127

additive effects on 483 dealumination and 50,161,162,263 environmental regulations and 177,178,581 gasoline octane enhancement and 70,163,165,499ff,576 matrix roles in 166-170,297 metals effect on 171 non-framework alumina and 50,162 poison effects on 62-64,174,175,265 silica/alumina ratio of zeolite and 48,57,61,71,155-157,162,509 sodium content effects on 42,159,483 structural changes and 173 zeolite acidity effect on 44,48,57,157 zeolite crystallite size effects on 90,163-165 zeolite mixtures and 165,499ff

Cracking chemistry and mechanisms 5,154,441-443 alkylation and dealkylation 4339,154,442,446 carbocations 153 H-transfer 153,155-157,390 metals effects on 171,340-346,522 olefin isomerization 43,53,59,60,443,444,501 paraffins 43,53,59-61,154,442,443,444,501 residence time 385,386,390 residuum feeds 165,415 structural effects 165 thermal cracking 42,43,6648,446 thermodynamics 442,444

Cracking catalyst selectivity 150,161,162,260

Crude oil 455,457,460-462,472,474,478-481,518 Cycle oil 145,451,485,504

Cyclones 27-28,406,568 recycle 386,387,389,395

D

Davison circulating riser (DCR) 183-184,18&189,198,207-211,223,270-278 Dealumination 50,53,!57,64,86,110,152,161,162,263,508,511 Dehydrogenation 48,64,68,265,339,443-445,483,522,523 Decanted oil (DO) 150,157,162,164,168,175 Dehydroxylation 47,48,57,263 Delta coke 240,400,431 DEMET 33,314,584

593

Demetallization 314 Dense bed cracking 24,224,386,526,565 Desalting 341 Diffusion 166 Dimetallic 9P2 359,363

Dry gas 145,150,151,169,170 Dynamic activity 225,238

DM-1152 363

E

ECR-4 85 ECR-30 85 ECR-32 86 Electron microscopy 95,96,101 Electrostatic precipitator 565,568,571,579 Environmental changes to 2001 A.D. 579-582

aromatic compounds in gasoline 177,178,244,581 decanted oil properties 582 isobutylene requirements 178,179 flue gas stack emissions 579 light cycle oil needs 582 product sulfur levels 177,437,581,582 reformulated fuels 39,177,178,437,527,580 waste disposal of spent FCC catalyst 582

catalyst losses as 568 decanted oil and 575 diesel fuel composition 177,178 fuel gas sulfur and 575 gasoline composition 177,178,576 impact on FCC product quality 574 light catalytic cycle oil and 576 regenerator stack emissions and 531-532,551,555,568,571,572,573 waste streams and 577,582

Environmental issues 177-179,568,571-577

Equilibrium catalysts 225,351,355,360-361,366,368 Erionite 87,96,174 Ethers 178 ETBE 178 Extra-framework alumina (see also non-framework alumina) 50,51,58,62-64,188-

189,209-21 2

F

Faujasi te (see also Zeolite Y) 41,45,83,85,86,88,146,186- 190,l 92- 1 94,196- 197,20 1 - 216

594

Feed distribution 431 method of 393,431 effect of 431

Feed preheat 271,272,273,422 Feedstocks

aromaticity 163 Conradson carbon in 295,473,474,478 contaminant metals in 172,478-483

effects on yields 482-484 passivation of 339-380,393,482-484

effect on yields 443,447,470,49931 1,515,518 N.D.M. method -ASTh4 D-3238-74 467,468,489 residual stocks 472,474

hydrogen content 455-459,463,464,473,474,488,489,492 effect on yields 455,456,463,489,492

mid-boiling point 451,454,464 molecular weight 454,455,460,467 nitrogen in 467,471,474,478,485-488,493 product quality as effected by 175,489-493 specific gravity 448,455,464,465,467,472 sulfur in 464,465,467,471,478,487 UOP K factor 158,163,448,450,453,454,455,464,468,470,472,489,491,493

hydrocarbon type

Ferrierite 87,97,100,174,525 Fixed bed cracking 11-14,270,318,526 Fluid catalytic cracking (see also Cracking categories) 385,456 FCCU design

Ashland 24,295,372,390,393,417 effect of design on yields 389,390,392

historical 17-28,386,565-568 1FP (Total) 394,416 Kellogg 25 MAT, comparison to commercial FCCU 245 Mobil 390 modern design 390 Shell 23,387,391 Texaco 394 UOP 24,386,393,394,410,417,43 1

FCCU reactor 22 carbon-on-regenerated catalyst in 428 circulation rate as a variable in 422 contact time in 385,386,390 design of 386-395,565-568 feed distribution in 393,431 pressure as a variable in 426,431

EXXON 17-21,28,386,390

595

temperature as a variable in 421 -423,442,444,446,450,451 FCCU regenerator 22

design for high coke making feeds 415-420 design of 565-568

historical 395 modern 41 0-420 solids mixing 409 two stage regenerators 416

effect of process variables on 401-404 effect of regenerator temperature on product yields 401-404 fluidization in 404-409 heat balance in 395-405 high efficiency types 24,395,410-415

bubbling bed regime 405 fast fluidized regime 408 turbulent bed regime 406

Fluidization

Formic acid 119 Fractionation 277,278 Future issues 436,527,578-584

G

Gadolinium 377 Gallium 371 Gallium ZSM-5 378 Gas chromatographic simulated distillation (GCSD) 224,229 Gas oil 446,450,465-467,474,487,499,500,502,505-507,509,516,518,519~26 Gasoline 145,446-449,451,453,454,459,461,463,466-469,471,485,486,488,489

composition ~,66,70,151,158,160,163,243,244,261,426,459,489,501,576 octane 146,157,158,161,162,166,171,176 reformulation 166,177,178,437,527,581

site acidity and 49,64,65 Geometric effects in cracking

Gmelinite 87,97

H

Hammett acidity function 54-56 Heat balance 240,275,385, 395-405,435,454,478

effect of coke burning variables on 397-399 effect of process variables on 350,379,401

Heat of combustion 400,403,405 Heat of reaction 399,442,444,445 Heat transfer 166,419 Heavy cycle oil 145,151

596

Heavy oil 52 Heavy Oil Cracker (HOC) 25-27,294,350 Hectorite 376 History

fixed bed 7,lO-14

FCC unit design 21-28,386-390,395-396,410-412,415-420,565-568 microactivity test 224 moving bed 12-17 place of FCC in refining operation 7-9 process development 386-395,409 thermal cracking 10,11,42,43,66-68 zeolite revolution 41,42,145

FCC 17-21

Houdry process 2,7,11-13,43 Hydride transfer 43,60,69-72 Hydrocarbon reactions 43,67-72,441-445 Hydrocracking 1,2,457 Hydrodesulfurization catalyst 340 Hydrogen factor 260 Hydrogen sulfide scrubbing Hydrogen transfer 43,60,69-72,153,155,156,157,178,264,390,4-43-

Hydrogen yield, reduction by metals passivation 347-364,366-368,483 Hydrotalcite 376,551 Hydrotreating 340,437,516 Hydroxyl groups 47,153,201-205,208,212-215 Hydroxyl nests 112

445,457,502,505,51 0,511,581

I

Indium 371 Infrared spectroscopy 55-59,64,101,202-204 Insertion reactions of silica species 167 Intergrowth structures in zeolites 96 Ion exchange 42/45 Iron 341,450,474,482,483 Isomerization 43,60,68,69,443,444,499,504,505,507,508,510,511,514 Isoolefins 43,244,51021 1,581 Isoparaffins 43,67,244,422,504,505,510

K

Kaolin 167 Kinetic diameter 194-195 Kinetics 4 Kinetics of coke burning 397-399

597

L

Lanthanum 42,47,370 Lewis acids 54,150,153,167 Lift gas 372,393,431 Lithium 371 Light cycle oil 145,151,157,159,161,162,164,166,168-

170,175,260,261,471,486,493,504574,576,582 Lowenstein's rule 49 LPG 151,154,157,162,168,169,175-1 77,228,428,448,454,457,459,471,489278,574276

M

MCM-22 84/87 Magnesium (see metal traps) Magnetic separation 315 Make-up rate 421 Manganese 359 Mass transfer 224 Matrix 146,166-169,297 Mazzite 95 Mechanism

of antimony-nickel passivation 349,350 of octane improvement with Z5M-5 501,509 of tin-vanadium passivation 365 of zeolite destruction by vanadium 265,341,343-346,522

Metal deposition 265,267,342443,345,349,373,374,522 Metal scavenging 377 Metal spinel 171,376,536,538-543,546-549 Metals deactivation 241,265,340,522,523 Metals flushing 340 Metals passivation 149,171,174,266,339

additive 146,147,265,340,346,483 agent 340,346-350,357-359,363-365,369-371 by antimony containing fines 355 benefits 339,358,483 chemistry 347-350,365 future direction of 380 mechanism of 349,357,365 modeling 359 of nickel (see nickel) operating conditions, effect on 361,372,393,431,483 protocol 350,359,363,366,380 tests, commercial 350-358,359-364,366-370274,375,483 tests, MAT 241,357,361,366-367,369,370 of vanadium (see vanadium)

598

water-based agents 348,370

DVT additive 376 magnesium oxide 266,376 for nickel 374-375 sepiolite 171,375 shot coke 377 sponge coke 377 for vanadium 375-377

Metal resistant cracking catalysts (see Resid cracking catalysts) Metal poisons 241,258,265 Microactivity test 269,285-290

Metal traps 149,167,171,175,266,372-378

applications 225,258,259-269 ASTM reference 225,257,259,269 cat/oil ratio in 225,247,259,264 comparative catalyst evaluations 236,238,260,261,262,526 contact time 246,259 feed pre-heat and vaporization 247 historical 224 metals testing 241 operation 225,228 product analysis 225,229 products from 161,169,170,228,229,259,260,261,262 reactor design 245 schematic of unit 226 steam deactivation of catalysts 229,236,259-261,263-267 temperature and pressure drop effects in 248 time averaging effects 247

Microspheres 120,126~ 67 Mitchell method (see also metals deactivation) 242 Mix temperature control

in combustor 413 in riser 394

Molar expansion 455 Molecular sieves (see also Zeolite) 205,499,524 Mordenite 41,56-58,87,100,205 Morphology 90,196 Moving bed 12-17 MTBE 178 Mullite 109,185,344

N

NMR (see also surface analysis and imaging) 55,57-59,64,65,%,205-211

Napthenates 343,441,443,445,523 NO, 532,551-555,573,580

599

Natural zeolites 41 Next nearest neighbor concept 153,154 Nickel 146,164,171-173,241,265,340-343,346-357,359-364,371-375,377-378

aluminate 342 Betz passivator for 359 dehydrogenation activity of 340,341,483,484322 deposition 198-200,342,349,374,478-483,488,522 passivation of 171,346-357,359-364,372,374-375,377-378,483 support, effect on 342

Nitrogen 165,166,174,175,192,193,450,467,471,474,478,485-488,493 Non-framework A1 64,71,111,1 12,114,115,161,162,198,209-21 1 Non-homogeneity in zeolites 93

0

Octane 145,146,157,158,163,171,176,177,~5,454,489,491,492 effect of conversion on 425 effect of gasoline boiling range on 424 effect of pressure on 428 effect of reactor temperature on 424 measurement of 261,278,279

Octane barrel 147,176 Offretite 87,96,165,524,525 Olefin yields

effect of conversion on 423 effect of pressure on 428 effect of temperature on 422

459,489,493,576,581 Olefins 43,59,66,70,84,154,155,158,160,163,168,176-178,244,441,443,445,455,457-

Operating variables 511,515 Orifice chamber Oxidation Oxygen in fuels 177,178 Oxygen partial pressure

effect on coke burning 402 effect on carbon-on-regenerated catalyst 404

600

P

Palladium 172 Paraffins 43,59,66,70,158,244,44 1,44346,457,455-457,459,467,468,471,473,501-

Passivation (see metals passivation) Phosphoric acid 345 Phosphorus 206,211,347,371 PIANO 244 Pillared clays (PILCs) 32,83,123,124 Platinum 147,172,175 PIONA 244 Plug flow riser cracking 390 Poisoning of catalysts 183-1 84,198-200,263,499,522 Polymerization 68,444 Pore size 44,45,87,162,166,167,169,171,190,192-197,300,474,477,526 Pore structure 41,44,190,192-197,211-212,474,525 Pore volume 167,169,193-1 95,300 Porphyrins 339,478 Pressure 426,431,505 Predicting commercial performance 248,358-359,378-379 Process constraints

Product analysis 277-279,513,527 Product yields

510,514-517,519,520,522,526,527,576,581

interaction with process variables 435

effect of process variables on 421436 historical changes 389

Products and environmental constraints 580-582 Promoter 83,403-405574 Propane 176-1 78,490,517 Propylene 1 76-1 78,429,436,454,490,502,508,516,517,5 19,525 Pseudoboehmite 119,122,167,574-548

R

Radical 42,59 Rare earth 128,131,146,147,149,152,153,155,156,162,165,176,370~05

cerium 42,47,347,364 lanthanum 42,47370 neodymium 47 praseodymium 47,370,377

Reaction networks 4-6 Reaction variables 445

carbon on regenerated catalyst 428-431 catalyst activity 435-436 catalyst circulation 425426

60 1

pressure 426,431 temperature

reactor 421-423 regenerator 434

Reactor 245,508 Reactor design 245,270,386-395 Recycle 386,387,389,395,451 Reduced Crude Conversion (RCC) 295,372 Reformulated fuels 177,178,437,527,580,581 Regenerator 51 1,520

CO oxidation in 172,272,303,403405,532,553,555-558,571,580 design of 270,272,395420,565-568 NO, control in 532,551-555,573,580 particulate control in 568,579 SO, control in 171,531-551355,572,580

Reid vapor pressure 177,178 Resid 146,149,164,169,339,415,461,488,516,523 Resid cracking catalysts 147,172,176,293,522

additives for 310 composition and design of 83,149,165,169,175,176,296 metals management of 171,313,339-380,522 modeling 304 testing and selection of 241,242

Resid cracking processes 294 Riser 269,386,390,393,507,508

S

SAP0 85,211,524,526,527 SEM 88,89,101,196-197 SO, 147,149,166~ 71,177,215,531-551355,572,580 Scavenger (see metal traps) Scrubbing 573,579 Sepeolite 171,375 Shape selective catalysts 87

commercial trials 499, 500,571-523 laboratory evaluation 499,500,502,513-51 6,524,525 octane enhancement by 499-501,504,505,507-510 pore size of 45,87,194-195,499,524 zeolites used in 45,87,499,524 ZSM-5 metals resistance in 522

effect of Si0,/Al,03 ratio on 93,508 effect of zeolite concentration on 508,516 general aspects of 45,71,499,500 mechanisms 500

Shape selectivity

602

olefin isomerization 504,505,507,510,511,514 paraffin conversion 500,502,505-51 1,516 product 500 reactant 500 spatioselectivity 500

Silica 42,109~ 17,118,120,125,146,147,166,34-4 Silicalite 165,346 Silica-alumina 42-44,116,122,123,146,155-157,162,166,167,203-204 Silica/ alumina ratio 108,109,110,111,114,115,128 Silicon insertion 112,113,114,115 Simulated distillation 229 Sites (see also acid sites) 42,43,45-61,63-65,67,68,70-72,201-204,501,508,511,524 Slide valve 270 Slurry 463 Smectites 123 Sodium 107,108,112,127,128,129,130,159,173,259,339,343-346,366,371,474,482,

Sodium aluminate 107,108 Solid state NMR 96,101,112,205-211 Solvent deasphalting 293,472

Spent catalyst 577,583 Spinel 147,167,185,376,536,538-543,546-549 Steam deactivation 183-1&1,188-189,197,229,236,259,263,264,07,541-

Steric hindrance in zeolites 52,59,500 Stripping 272,392,400 Sulfur 347,451,474,487,493,532-533 Superacid 50,57,63,64,153,205 Surface analysis and imaging 195-201

483,485,523

SP~C~~OSCOPY 55-59,64,136,201-211

543

applications 195-201 electron microscopy 97,98,99,196-197 elemental and surface analysis

AES, Auger electron spectroscopy 199 electron microprobe 197-198,245 FARMS, fast atom bombardment mass spectroscopy 200 SEM, scanning electron microscope 196-197 SIMS, secondary ion mass spectroscopy 199-200,342 STEM, scanning transmission electron microscope 196-197,376 STM, scanning tunneling microscopy 200 TEM, transmission electron microscope 196-197 UPS, Ultraviolet photoelectron spectroscopy 198

spectroscopic surface analytical techniques 201-21 1 JX and protonic acidity, Lewis acidity 55-59,64,201-204 MASNMR and zeolite framework structures 5527-59,64,65,206-210 Raman and zeolite framework structures 188,205

603

UV and acid site strength 55,56,59,204-205 Surface area 110,119,120,122,123,124,125,136,192-193,232,341,367,523,524547 Syncrude 226,228

T

TAME tertiary amyl methy ether 178 Tantalum 371 TCC (Thermofor Catalytic Cracking) 7,12,14-17,353,499,500,511-514,519 Tellurium 371 Temperature 17,442,444,446,450,457,502,505,507,508,516

reactor control of 421 effect on performance 422-423

control of 401-405,418,420 effect on reactor yields 434

regenerator

Tetraethyl lead 70 Thermal analysis 212-216,534-537,547 Thermal cracking 1,lO-11,42,43,66-68,446,455,485,500 Thermodynamics 400,405,442,444,532,553,555 Thioethers 372 Time-on-stream 175 Tin 365

additive 146,171,365 commercial tests 366-369

Toxic characteristic leaching procedure (TCLP) 380,577 Transport disengaging height 406 Tungsten 371

U

USY (see Zeolite, USY) Ultra-short contact time 246,380 Unit cell 50,70,111,127,128,131,155,156-160,161,163,175,188-190,263,505,506

V

VPI-5 32/83 Vanadium 147s 65,166,171,163,265,340,343,522,523

deposition 191,198-200,343,345,374,478-483 eutectic 343 passivation of 122,146,171,191,364372,375-378,483 pentoxide 343,534 testing impact 241

604

vanadates 191,344 vanadic acid 122,265,266,345,522 zeolite destruction by 122,265,340-341,343-346,365,522

W

Water pore volume 136,194 Withdrawal of catalyst 421

X

XPS 51,55,198,342,365 X-ray procedures 65,785-190 X type zeolite (see Zeolite,X) Xenon adsorption, NMR 211-212

Y

Y type zeolite (see Zeolite,Y)

Z

Zeolite 245 A 41 ALP0 86,165,524 beta 41,87,165 Bronsted acidity 45-50,54-59,61,63-65,67,71,150,153,159,167,201-205,213-215 BSS 86 cloverite 83 CREY (calcined rare earth Y) 168,344 crystal shape of 88,90,196 crystal size of 88,90,164,165,192,196 csz-1 84 csz-3 84 dealumination of 152,155,161,162,173 dispersion of 92

ECR-4 84

erionite 87,96,174 ferrierite 87,100,174,525 gmelinite 87 high silica Y 146,151,152,153,156,157,161,165,178 homogeneity of 93 intergrowth structure in 95 Lewis acidity 47,48,50,54-57,59,61,63,64,67,71,152,167,243-215

ECR-4 96

ECR-32 85

605

LZ-210 83 MCM-22 84 mixtures of 165 mordenite 41,56-58,87,100,165,205 morphology of 88,89,90,196 non-homogeneity of 93/94/95 offretite 87,%,165,524525 pore structure of 41,44,45,194-195,211-212524,525 primary promoters 85 SAP0 85/21 1,524,526,527 secondary promoters 85 silicalite 165,346 steric hindrance in 52,58,100,500 US-Y 50,52,63,83,111,152,157,16 1 ,I 62,165,166,183,184,187-1 90,197-1 98,207-

VPI-5 83,194

Y 5,41,42,44,45,47-53,59-59,61-65,69-

ZSM-2 85 ZSM-3 85 ZSM-5 6,29,31,33,41,44,45,49,56-59,61,62,64,65,71,85,91,92,146,149,163,

ZSM-20 85

21 1,264,346,499,502,507,508,516,523

X 44,84,105,106,146

71,84,88,105,107,127,146,149,152,153,155,162,174,178,183-184

165,171 ,I 74,176,178,186,187,202,2O4,213,378,499-523

Zeolite/matrix ratio 150,169,170 Zeolite synthesis 41,106

STUDIES IN SURFACE SCIENCE AND CATALYSIS

Advisory Editors: B. Delmon, Universite Catholique de Louvain, Louvain-la-Neuve, Belgium J.T. Yates, University of Pittsburgh, Pittsburgh, PA, U S A .

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Preparation of Catalysts IScientific Bases for the Preparation of Heterogeneous Catalysts. Proceedings of the First International Symposium, Brussels, October 14-17,1975 edited by B. Delmon, P.A. Jacobs and G. Poncelet The Control of the Reactivity of Solids. A Critical Survey ofthe Factors that Influence the Reactivity of Solids, with Special Emphasis on the Control of the Chemical Processes in Relation to Practical Applications by V.V. Boldyrev, M. Bulens and B. Delrnon Preparation of Catalysts II. Scientific Bases for the Preparation of Heterogeneous Catalysts. Proceedings of the Second International Symposium, Louvain-la-Neuve, September 4-7,1978 edited by B. Delmon, f? Grange, P. Jacobs and G. Poncelet Growth and Properties of Metal Clusters. Applications to Catalysis and the Photographic Process. Proceedings of the 32nd International Meeting ofthe Societe de Chimie Physique, Villeurbanne, September 24-28,1979 edited by J. Bourdon Catalysis by Zeolites. Proceedings of an International Symposium, Ecully (Lyon), September9-11,1980 edited by B. Imelik, C. Naccache, Y. Ben Taarit, J.C. Vedrine, G. Coudurier and H. Praliaud Catalyst Deactivation. Proceedings of an International Symposium, Antwerp, October 13- 15,1980 edited by B. Delmon and G.F. Frornent New Horizons in Catalysis. Proceedings of the 7th International Congress on Catalysis,Tokyo, June3O-Julyl, 1980. Parts Aand B edited by 1. Seiyama and K. Tanabe Catalysis by Supported Complexes by Yu.1. Yerrnakov, B.N. Kuznetsov and V.A. Zakharov Physics of Solid Surfaces. Proceedings of a Symposium, Bechyiie, September 29-October 3,1980 . edited by M. LazniEka Adsorption atthe Gas-Solid and Liquid-Solid Interface. Proceedings of an International Symposium, Aix-en-Provence, September 21-23,1981 edited by J. Rouquerol and K.S.W. Sing Metal-Support and Metal-Additive Effects in Catalysis. Proceedings of an International Symposium, Ecully (Lyon), September 14-16,1982 edited by B. Imelik, C. Naccache, G. Coudurier, H. Praliaud, P. Meriaudeau, P. Gallezot, G.A. Martin and J.C. Vedrine Metal Microstructures in Zeolites. Preparation -Properties -Applications. Proceedings of a Workshop, Bremen, September 22-24,1982 edited by P.A. Jacobs, N.I. Jaeger, P. Jik and G. Schulz-Ekloff Adsorption on Metal Surfaces. An Integrated Approach edited by J. Benard Vibrations at Surfaces. Proceedings of the Third International Conference, Asilomar, CA, September 1-4,1982 edited by C.R. Brundleand H. Morawitz

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Heterogeneous Catalytic Reactions Involving Molecular Oxygen by G.I. Golodets Preparation of Catalysts 111. Scientific Bases for the Preparation of Heterogeneous Catalysts. Proceedings of the Third International Symposium, Louvain-la-Neuve, September6-9.1982 edited by G. Poncelet, P. Grange and P.A. Jacobs Spillover of Adsorbed Species. Proceedings of an International Symposium, Lyon-Villeurbanne, September 12-16,1983 edited by G.M. Pajonk, S.J. Teichner and J.E. Germain Structure and Reactivity of ModifiedZeolites. Proceedings of an International Conference, Prague, July9-13,1984 edited by P.A. Jacobs, N.I. Jaeger, P. JiN, V.B. Kazansky and G. Schulz-Ekloff Catalysison theEnergy Scene. Proceedings of the 9th Canadian Symposium on Catalysis, Quebec, P.Q., September 30-October 3,1984 edited by S. Kaliaguine and A. Mahay Catalysis by Acidsand Bases. Proceedings of an International Symposium, Villeurbanne (Lyon), September 25-27,1984 edited by B. Irnelik, C. Naccache, G. Coudurier, Y. Ben Taarit and J.C. Vedrine Adsorption and Catalysis on Oxide Surfaces. Proceedings of a Symposium, Uxbridge, June 28-29,1984 edited by M. Che and G.C. Bond Unsteady Processes in Catalytic Reactors by YuSh. Matros Physics of Solid Surfaces 1984 edited by J. Koukal Zeolites: Synthesis, Structure, Technology and Application. Proceedings of an International Symposium, Portoroi-Portorose, September 3-8,1984 edited by B. Driaj, S. HoEevar and S. Pejovnik Catalytic Polymerization of Olefins. Proceedings of the International Symposium on Future Aspects of Olefin Polymerization, Tokyo, July 4-6,1985 edited by T. Keii and K. Soga Vibrations at Surfaces 1985. Proceedings of the Fourth International Conference, Bowness-on-windermere, September 15-19,1985 edited by D.A. King, N.V. Richardson and S. Holloway Catalytic Hydrogenation edited by L. Cerven+ New Developments in Zeolite Science and Technology. Proceedings of the 7th International Zeolite Conference,Tokyo, August 17-22,1986 edited by Y. Murakarni, A. lijima and J.W. Ward Metal Clusters in Catalysis edited by B.C. Gates, L. Guczi and H. Knozinger Catalysis and Automotive Pollution Control. Proceedings of the First International Symposium, Brussels, September 8-1 1,1986 edited by A. Crucq and A. Frennet Preparation of Catalysts IV. Scientific Bases for the Preparation of Heterogeneous Catalysts. Proceedings of the Fourth International Symposium, Louvain-la-Neuve, September 1-4,1986 edited by B. Delmon, P. Grange, P.A. Jacobs and G. Poncelet Thin Metal Films and Gas Chemisorption edited by P. Wissrnann Synthesis of High-silica Aluminosilicate Zeolites edited by P.A. Jacobs and J.A. Martens Catalyst Deactivation 1987. Proceedings of the 4th International Symposium, Antwerp, September 29-October 1,1987 edited by B. Delmon and G.F. Froment

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Keynotes in Energy-Related Catalysis edited by S. Kaliaguine Methaneconversion. Proceedings of a Symposium on the Production of Fuels and Chemicalsfrom Natural Gas, Auckland, April 27-30, 1987 edited by D.M. Bibby, C.D. Chang, R.F. Howe and S. Yurchak Innovation in Zeolite Materials Science. Proceedings of an International Symposium, Nieuwpoort, September 13-17,1987 edited by P.J. Grobet, W.J. Mortier, E.F. Vansant and G. Schulz-Ekloff Catalysis 1987. Proceedings ofthe 10th North American Meeting ofthe Catalysis Society, San Diego, CA, May 17-22,1987 edited by J.W. Ward Characterization of PorousSolids. Proceedings of the IUPAC Symposium (COPS I), Bad Soden a. Ts., April 26-29,1987 edited by K.K. Unger, J. Rouquerol, K.S.W. Sing and H. Kral Physics of Solid Surfaces 1987. Proceedings of the Fourth Symposium on Surface Physics, Bechyne Castle, September 7-1 1,1987 edited by J. Koukal Heterogeneous Catalysis and Fine Chemicals. Proceedings of an International Symposium, Poitiers, March 15-17,1988 edited by M. Guisnet, J. Barrault, C. Bouchoule, D. Duprez, C. Montassier and G. Perot Laboratory Studies of Heterogeneous Catalytic Processes by E.G. Christoffel, revised and edited by 2. Paal Catalytic Processes under Unsteady-State Conditions by Yu. Sh. Matros Successful Design of Catalysts. Future Requirements and Development. Proceedings of the Worldwide Catalysis Seminars, July, 1988, on the Occasion of the 30th Anniversary of the Catalysis Society of Japan edited by T. lnui Transition Metal Oxides. Surface Chemistry and Catalysis by H.H. Kung Zeolites as Catalysts, Sorbents and Detergent Builders. Applications and Innovations. Proceedings of an International Symposium, Wurzburg, September 4-8.1988 edited by H.G. Karge and J. Weitkamp Photochemistry on Solid Surfaces edited by M. Anpo and T. Matsuura Structure and Reactivity of Surfaces. Proceedings of a European Conference, Trieste, September 13-16, 1988 edited by C. Morterra, A. Zecchina and 0. Costa Zeolites: Facts, Figures, Future. Proceedings of the 8th international Zeolite Conference,Amsterdam, July 10-14,1989. Parts Aand B edited by P.A. Jacobs and R.A. van Santen Hydrotreating Catalysts. Preparation, Characterization and Performance. Proceedings of the Annual International AlChE Meeting, Washington, DC, November 27-December 2,1988 edited by M.L. Occelli and R.G. Anthony New Solid Acids and Bases. Their Catalytic Properties by K. Tanabe,M. Misono, Y. Ono and H. Hattori Recent Advances in Zeolite Science. Proceedings of the 1989 Meeting of the British Zeolite Association, Cambridge, April 17-19, 1989 edited by J. Klinowsky and P.J. Barrie Catalyst in Petroleum Refining 1989. Proceedings of the First International Conference on Catalysts in Petroleum Refining, Kuwait, March 5-8, 1989 edited by D.L. Trimm, S. Akashah, M. Absi-Halabi and A. Bishara

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Future Opportunities in Catalytic and Separation Technology edited by M. Misono, Y. Moro-oka and S. Kimura New Developments in Selective Oxidation. Proceedings of an International Symposium, Rimini, Italy, September 18-22,1989 edited by G. Centi and F. Trifiro Olefin Polymerization Catalysts. Proceedings of the International Symposium on Recent Developments in Olefin Polymerization Catalysts,Tokyo, October 23-25,1989 edited by T. Keii and K. Soga

Surface Analysis edited by J.L.G. Fierro

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Volume 57A Spectroscopic Analysis of Heterogeneous Catalysts. Part A: Methods of

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Volume 69

Probe Molecules edited by J.L.G. Fierro Introduction t o Zeolite Science and Practice edited by H. van Bekkum, E.M. Flanigen and J.C. Jansen HeterogeneousCatalysis and Fine Chemicals II. Proceedings of the 2nd International Symposium, Poitiers, October 2-6, 1990 edited by M. Guisnet, J. Barrault, C. Bouchoule, D. Duprez, G. Perot, R. Maurel and C. Montassier Chemistry of Microporous Crystals. Proceedings of the International Symposium on Chemistry of Microporous Crystals, Tokyo, June 26-29,1990 edited by T. Inui, S. Namba and T. Tatsumi Natural Gas Conversion. Proceedings of the Symposium on Natural Gas Conversion, Oslo, August 12-17,1990 edited by A. Holmen, K.-J. Jens and S. Kolboe Characterization of Porous Solids II. Proceedings of the IUPAC Symposium (COPS Il),Alicante, May6-9,1990 edited by F. Rodriguez-Reinoso, J. Rouquerol, K.S.W. Sing and K.K. Unger Preparation of Catalysts V. Proceedings of the Fifth International Symposium on the Scientific Bases for the Preparation of Heterogeneous Catalysts, Louvain-la-Neuve, September 3-6,1990 edited by G. Poncelet, P.A. Jacobs, P. Grange and B. Delmon New Trends in CO Activation edited by L. Guczi Catalysis and Adsorption by Zeolites. Proceedings of ZEOCAT 90, Leipzig,

edited by G. Ohlmann, H. Pfeifer and R. Fricke Dioxygen Activation and Homogeneous Catalytic Oxidation. Proceedings of the Fourth International Symposium on Dioxygen Activation and Homogeneous Catalytic Oxidation, Balatonfured, September 10-14,1990 edited by L.I. Simandi Structure-Activity and Selectivity Relationships in Heterogeneous Catalysis. Proceedings of the ACS Symposium on Structure-Activity Relationships in Heterogeneous Catalysis, Boston, MA, April 22-27,1990 edited by R.K. Grasselli and A.W. Sleight Catalyst Deactivation 1991. Proceedings of the Fifth International Symposium, Evanston, IL, June24-26,1991 edited by C.H. Bartholomew and J.B. Butt Zeolite Chemistry and Catalysis. Proceedings of an International Symposium, Prague, Czechoslovakia, September 8-13,1991 edited by P.A. Jacobs, N.I. Jaeger, L. Kubelkova and B. Wichterlova

August 20-23,1990

Volume 70

Volume 71

Volume 72

Volume 73

Volume 74

Volume 75

Volume 76

Volume 77

Poisoning and Promotion in Catalysis based on Surface Science Concepts and Experiments by M. Kiskinova Catalysis and Automotive Pollution Control II. Proceedings of the 2nd International Symposium (CAPoC 2). Brussels, Belgium, September 10-13,1990 edited by A. Crucq New Developments in Selective Oxidation by Heterogeneous Catalysis. Proceedings of the 3rd European Workshop Meeting on New Developments in Selective Oxidation by Heterogeneous Catalysis, Louvain-la-Neuve, Belgium, April 8-10,1991 edited by P. Ruiz and 6. Delmon Progress in Catalysis. Proceedings of the 12th Canadian Symposium on Catalysis, Banff, Alberta, Canada, May 25-28,1992 edited by K.J. Smith and E.C. Sanford Angle-Resolved Photoemission. Theory and Current Applications edited by S.D. Kevan New Frontiers in Catalysis, Parts A-C. Proceedings of the 10th International Congress on Catalysis, Budapest, Hungary, 19-24 July, 1992 edited by L. Guczi, F. Solymosi and P. Tetenyi Fluid Catalytic Cracking: Science and Technology edited by J.S. Magee and M.M. Mitchell, Jr. New Aspects of Spillover Effect in Catalysis. For Development of Highly Active Catalysts. Proceedings of the Third International Conference on Spillover, Kyoto, Japan,August 17-20,1993 edited by T. Inui, K. Fujirnoto, T. Uchijima and M. Masai

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