DOE/PC 95050-

Exploratory Research on Novel CoalLiquefaction Concept

Task 4 and Final Report

November 1998

S. D. Brandes, R. A. Winschel CONSOL Inc.Research & Development4000 Brownsville RoadLibrary, PA 15129

F. J. Derbyshire, G. M. Kimber, R. K. Anderson, D. N.Jacques, T. D. RantellUniversity of Kentucky Center for Applied Energy Research3572 Iron Works PikeLexington, KY 40511

M. PelusoLDP Associates32 Albert E. Bonacci DriveHamilton Square, NJ 08690

Under Contract to:

United States Department of EnergyContract No. DE-AC22-94PC95050

U.S. DOE PATENT CLEARANCE IS REQUIRED PRIORTO THE PUBLICATION OF THIS DOCUMENT

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DISCLAIMER

This report was prepared as an account of work sponsored by an agency of the United StatesGovernment. Neither the United States Government nor any agency thereof, nor any of theiremployees, makes any warranty, express or implied, or assumes any legal liability o rresponsibility for the accuracy, completeness, or usefulness of any information, apparatus ,product, or process disclosed, or represents that its use would not infringe privately owne drights. Reference herein to any specific commercial product, process, or service by trade name,trademark, manufacturer, or otherwise does not nec essarily constitute or imply its endorsement,recommendation, or favoring by the United States Government or any agency thereof. Th eviews and opinions o f authors expressed herein do not necessarily state or reflect those of theUnited States Government or any agency thereof.

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ABSTRACT

The report presents a summary the work performed under DOE Contract No. DE-AC22 -

95PC95050. Investigations performed under Task 4 - Integrated Flow Sheet Testing ar e

detailed. In this program, a novel direct coal liquefaction technology was investigated b y

CONSOL Inc. with the University of Kentucky Center for Applied Energy Research and LD P

Associates. The process concept explored consists of a first-stage coal dissolution step i n

which the coal is solubilized by hydride ion donation. In the second stage, the products ar e

catalytically upgraded to refi nery feedstocks. Integrated first-stage and solids-separation steps

were used to prepare feedstocks for second-stage catalytic upgrading. An engineering an d

economic evaluation was conducted concurrently with experimental work throughout th e

program.

Approaches to reduce costs for a conceptual commercial plant were recommended at th e

conclusion of Task 3. These approaches were investigated in Task 4. The economic analysis

of the process as it was defined at the conclusion of Task 4, indicates that the production o f

refined product (gasoline) via this novel direct liquefaction technology is higher than the cos t

associated with conventional two-stage liquefaction technologies.

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TABLE OF CONTENTS

Page

CONTRACT SUMMARY 1Program Description 1

Technical Concept 1Program Organization 4

Technical Accomplishments 5Task 2 - Evaluation of Process Steps 5Task 3 - Flow Sheet Development 7Task 4 - Integrated Flow Sheet Testing 8Task 5 - Engineering and Economic Study 10

CONTRACT CONCLUSIONS AND RECOMMENDATIONS 12Conclusions 12Recommendations 13

TASK 4 15Summary 15Background 16

First-Stage Solubilization 16Filtration 17Catalytic Upgrading 18Engineering and Economic Evaluation 19

Experimental 20Reagents 20Equipment 21Procedures 22

Results and Discussion 25First-Stage Tests 25Filtration 31Second-Stage Catalytic Upgrading 34Conclusions and Recommendations 40

REFERENCES 42

LIST OF TABLES

Table Page

1 Analyses of Black Thunder Mine Subbituminous Coal 432 Analyses of Reilly Industries Anthracene Oil and Recycle Distillate

from Run 23-LA Series 443 Analyses of Molyvan L, Molyvan A, and Dimethyldisulfate 444 Boiling Range Distribution of Molyvan L 455 Task 4 - First-Stage One-Liter Autoclave Tests Summary 46

iv

LIST OF TABLES (Continued)

Table Page

6 Task 4 - First-Stage System-Condition Goals 477 First-Stage Microautoclave Tests - Effect of KOH Catalyst Loading on

Coal Conversion 478 First-Stage Microautoclave Tests at Reduced Solvent/Dry Coal Ratios 489 First-Stage Microautoclave Tests - Effect of Recycle Solvent on Coal

Conversion; Reilly Industries Anthracene Oil vs. Recycled Distillatefrom Run 23-LA Series 48

10 Task 4 - Off-Line Filtration Tests 49 11 Mass Balance Data 50 12 Filter Cake Vacuum Drying 50 13 Microautoclave Hydrotreating Tests on Distillation Residues from

Run 23h-LA Filtrate 51 14 Microautoclave Catalyst Tests on the Distillation Residue of Run 23h-LA

Filtrate 52 15 Microautoclave Tests of Residue Recycle 53 16 Mass Balance Data for Feedstock Tests, 300 mL CSTR Runs 54 17 Conversion Data for Feedstock Tests, 300 mL CSTR Runs 55 18 Overall Elemental Balance for Feedstock Tests, 300 mL CSTR Runs 55 19 Elemental Analyses for Feedstock Tests, 300 mL CSTR Runs 56 20 Boiling Range Distribution for Feedstock Tests, 300 mL CSTR Runs,

wt %, as Determined 57 21 Feed and Product Distribution for Feedstock Tests, 300 mL CSTR Runs 58 22 Mass Balance Details for Catalyst Tests, 300 mL CSTR Runs 59 23 Hydrotreating Results for Catalyst Tests, 300 mL CSTR Runs 60 24 Overall Elemental Balance for Catalyst Tests, 300 mL CSTR Runs 60 25 Elemental Analyses for Catalyst Tests, 300 mL CSTR Runs 61 26 Boiling Range Distribution for Catalyst Tests, 300 mL CSTR Runs,

wt %, as Determined 62 27 Feed and Product Distribution for Catalyst Tests, 300 mL CSTR Runs 63 28 Mass Balance Data for Residence Time Tests, 300 mL CSTR Runs 64 29 Conversion Data for Residence Time Tests, 300 mL CSTR Runs 65 30 Overall Elemental Balance for Residence Time Tests, 300 mL CSTR

Runs 65 31 Elemental Analyses for Residence Time Tests, 300 mL CSTR Runs 66 32 Boiling Range Distribution for Residence Time Tests, 300 mL CSTR

Runs, wt % as Determined 67 33 Feed and Product Distribution for Residence Time Tests, 300 mL

CSTR Runs 68

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LIST OF FIGURES

Figure Page

1 Simplified Block Flow Diagram, Liquefaction System, Original NovelConcept Case 69

2 Simplified Block Flow Diagram, Conceptual Commercial Plant,Original Novel Concept Case 70

3 Simplified Block Flow Diagram, Liquefaction System, Improved NovelConcept Case 71

4 Effect of HI "E" on Coal Conversion 72 5 Effect of Boiling Point Distribution on Resid Conversion 73

6 Effect of Catalyst Concentration on Conversion 747 Feedstock Conversion and Hydrogen Uptake 758 The Influence of Catalyst Type and Concentration on the Conversion of

Feedstock 41LA 769 The Influence of Residence Time on the Conversion of Feedstock 41LA 77

10 Feedstock Hydrocracking: Effect on Boiling Point Distribution 78 11 Catalyst Testing: Effect on Boiling Point Distribution 79 12 Residence Time Testing: Effect on Boiling Point Distribution 80

LIST OF APPENDICESAppendix Page

1 Hydride Ion Source Production and Hydride Ion Coal LiquefactionStudies - Annotated Bibliography A-1

LIST OF CONFIDENTIAL APPENDICES

Appendix Page

CA-1 Hydride Ion Sources (NOT IN GENERAL DISTRIBUTION) CA1-1 CA-2 Engineering and Economic Evaluation (NOT IN GENERAL DISTRI-

BUTION) CA2-1 CA-3 Economic Implications of Filtration (NOT IN GENERAL DISTRIBUTION) CA3-1

CH3OOCH H2O HCOO CH3OH H

HCOO COAL COAL H CO2

COAL H H2O COAL H2 OH

CH3OOCH COAL H2O COAL H2 CH3OH CO2

CH3OH CO CH3OOCH

COAL H2O CO CO2 COAL H2

1

(1)

(2)

(3)

(4)

(5)

(6)

CONTRACT SUMMARY

PROGRAM DESCRIPTION

TECHNICAL CONCEPT

A novel direct coal liquefaction technology is being investigated in a program being conducted

by CONSOL Inc. under DOE Contrac t DE-AC22-95PC95050 in collaboration with the University

of Kentucky Center for Applied Energy Research and LDP Associates. The novel concep t

consists of a new approach to coal liquefaction which avoids some of the inherent limitations of

current high-temperature thermal liquefaction processes. The chemistry employed is based on

hydride ion donation to solubilize coal at low temperatures. These temperatures (350-400 C)

are significantly lower than typically used in conven tional coal liquefaction. Reaction at these low

temperatures results in high conversion of the coal to a solubilized form, with little hydrocarbon

gas make, and avoids the thermally induced retrograde reactions that are unavoidable i n

conventional thermal processes. In addition, for low-rank coals, a substantial portion of th e

oxygen in the coal is removed as CO and CO during the dissolution. The higher selectivity to2

liquid products and rejection of oxygen as the carbon oxides were expected consequences o f

coal solubilization via this chemistry and were found to result in improved hydrogen utilization.

The basis of the novel concept is the discovery made by CONSOL R&D that certain reagents are

capable of hydrolyzing to form hydride transfer agents that are very active for coal dissolution at

temperatures in the range of 350 C. Although a detailed reaction mechanism has not bee n

defined, the overall chemistry of the reaction is believed to involve the steps outlined below. This

example is based on methyl formate as the hydride transfer reagent.

Liquefaction:

Regeneration of Hydride Ion Source:

Overall Reaction:

2

A key element of the proposed research was to identify the optimum hydride transfer chemistry;

thus, a suite of hydride transfer reagents were investigated. Because of the proprietary nature

of the prior work, hydride ion reagents tested in the program are held confidential. They ar e

referred to throughout this report by a code letter. Confidential Appendix 1 (not in genera l

distribution) is the key to the code.

Several features of this reaction sequence are particularly important. First, as illustrated in this

example, it is possible for the hydrogen in the liquefaction reaction to be supplied to the coal by

water, through hydrolysis of, for example, methyl formate. It is possible to supply the water for

hydrolysis as the moisture inherent in the coal. Therefore, there is no need or less need fo r

hydrogen (or any reducing gas) in the dissolution rea ctor. In fact, gaseous hydrogen is liberated.

In Reaction 2, the formate species produced by hydr olysis reduces the coal in, what is presumed

to be, a hydride transfer mechanism. Because both of these reactions take place sequentially

in the liquefaction reactor, it is necessary to operate in a temperature range where the rates of

hydrolysis and hydride transfer are approximately the same. In th is (and earlier) work, convincing

evidence was developed that such a temperature range exists.

In this example, for the regeneration reaction, the methanol produced by methyl format e

hydrolysis is reacted with carb on monoxide to make methyl formate. For low-rank coals, a good

portion of the CO needed for this reaction may be produced in the dissol ution reactor directly from

the coal. In a sense, methanol acts as the i ntermediate in a two-step water gas shift reaction, as

suggested by the overall reaction.

A process design was formulated at the onset of this program (Figure 1). It was modified a s

experimental data became available and economic evaluations were completed. The initia l

design concept was that of a two-stage system: the first-stage product is fed to a separatio n

system where ash and the s mall amount of unreacted coal are removed; the solubilized product

then is fed to a second-stage reactor where it is catalytically hydrogenated and converted t o

distillate products. The process economics were evaluated in the program with the solid s

removal step located between reactor stages and after the second stage. In the fina l

configuration, an economic advantage is gained if the solids removal system is located after the

second stage. This is necessitated, in part, because of the high capital costs required for th e

first-stage processing at the level of severity necessary to achieve high coal conversions .

Additionally, interstage depressurization and repressurization of process streams in and out of

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the interstage filter added significantly to the capital cost. As part of the experimental program and

the resulting process design , a number of hydride transfer agents were explored. Each required

a different regeneration process configuration. The effect on the process economics wa s

determined for each reagent.

In the conceptual configuration, coal is slurried in a recycle solvent and fed to the low-temperature

reaction stage. The purpose of the solvent is to transport the coal and the solubilized products.

The solvent must be a good physical solvent for the low-temperature reaction products, but i t

does not need to be a hydrogen donor because the necessary hydrogen is provided by th e

hydride reagent. Solvents with relatively high aromatic hydrogen contents promoted coa l

conversion in the solubilization step. The mechanism for this effect was not investigated ;

however, it was speculated that the aromatic solvent could solubilize reaction products as soon

as they formed and prevent retrogressive reactions. Therefore, solvents tested were those that

have the most desirable physical properties for coal conversion, the subsequent transport, and

separation step. The principal product gases from the low-temperat ure solubilization step are CO,

CO , H , H O and hydrocarbon gases. Although H S and NH likely were formed, none wa s2 2 2 2 3

detected in the product gases.

In the early conceptual process design, the separation section was located between the coa l

solubilization and upgrading steps. The ash and IOM were removed, and gases are recovered

for recycle or further processing. Solids separation was effected through filtration. Othe r

techniques, similar to solvent deasphalting, were considered. However, filtration was found to

be so effective that no other method was tested. The decision to locate the separation section

ahead of the second stage was made contingent on the economic evaluation and th e

determination of the benefits to be had in resulting improvement in second-stage performance.

One predicted advantage of interstage deashing in the conceptual process would be if an all -

dispersed catalyst second stage is used. This would allow the catalyst to be recycled withou t

significant rejection with the ash and unconverted coal.

In the early conceptual process design, the solids-free separator product was vacuum distilled

and fed to a catalytic hydrogenation reactor for conversion to distillable products. Slurry-phase

catalyst could be used in the second-stage because the solids-free feed allows for high catalyst

concentrations while avoiding th e loss of catalyst with the ash and IOM. As stated above, in the

absence of significant solids formation, almost the entire second-stage residual product can be

4

recycled around the hydrocracker. This could allow operation at lower temperature to improve

selectivity and hydrogen utilization, and to decrease the potential for retrograde reactions. Some

level of second-stage recycle was found to be necessary to achieve high overall conversions.

At the first-stage conditions used to carry out the final tests in the program (375 C, 60 min,

solvent/coal of 1.4), the economic evaluation indicates that interstage separation is too costly .

Post second-stage solids filtrat ion was investigated. Solids-containing first-stage products were

upgraded in the second stage and then filtered. The reactivity of first-stage solids-containin g

products was the same in the second stage as was the reactivity of the solids-free first-stag e

products. However, filtration of the upgraded produc ts was not as facile as interstage separation

of solids. Post-second-stage filtration could be improved with the use of filter aids. An economic

evaluation of the filtration process is presented in Confidential Appendix 3.

In the example presented above, the methanol that results from methyl formate hydrolysis i s

recovered and recycled to a formate synthesis reactor. A suite of hydride transfer agents were

tested in the program and shown to be active for liquefaction. The recovery and regeneratio n

scheme depends on the hydride agent.

PROGRAM ORGANIZATION

The work in this contract followed a Task structure. The first Task of the project was th e

production of a management plan . An Engineering and Econ omic Study (Task 5) was conducted1

throughout the project to establish economic feasibility and to provide guidance at key decision

points at the completion of project Tasks. Task 2 was an evaluation of reaction steps for th e

envisioned process. Task 2 included a parametric study of low-temperature hydride io n

liquefaction, an evaluation of filtration and solvent separation to deash the liquefaction product,

evaluation and selection of catalysts, and a literature and engineering study of hydride io n

regeneration. The results of Task 2 were used to define the reaction steps and establish a

conceptual flow sheet and design for partially integrated testing. In Task 3, experimental work

was done to partially integrate the process steps of coal solubilizat ion and solids separation. Also

under Task 3, a second-stage catalytic reactor design was developed and the reactor system was

constructed and used to evaluate catalytic conversion options pertinent to a second-stag e

design. Concurrent with Tasks 2 and 3, a conceptual, f ully integrated design for the process was

devised under Task 5. The goal of Task 4 was to test experimentally the design devised under

5

Task 5 in a manner as fully integrated as possible. Even though this work was exploratory, the

concurrent engineering and economic analysis was proved to be an excellent way to asses s

results and provide criteria for deciding on experime ntal alternatives. Task 2 and Task 3 findings

were reported previously. The conceptual process design was revised based on the findings2,3

of Task 4 and recommendations for further work were made. The findings of Task 4 are included

in this report. Recommendations made specific to work performed under Task 4 also ar e

included. Conclusions and recommendations for the program are provided below.

TECHNICAL ACCOMPLISHMENTS

TASK 2 - EVALUATION OF PROCESS STEPS

The goal of Task 2 was to test the concept that hydride ion promoted coal liquefaction could be

accomplished under low-severity conditions and that it would provide a solubilized coal product

at high coal conversions. In Task 2, a systematic study was condu cted in 45 mL microautoclaves

to determine the range of low-severity first-stage reaction conditions that produce high coa l

conversions. Variables tested were coal rank, residence time, reaction t emperature, solvent type,

solvent-to-coal ratios, hydride ion sources, hydride ion reagent-to-coal ratios, and total reactor

charge size. Low-severity reaction conditions ( 400 C, 60 min) were discovered for which high

coal conversions (>90 wt %) to soluble products were obtained with five different coals of three

different coal ranks. Some solvent is necessary to achieve high coal conversion, although it does

not have to be a hydrogen donating solvent. The aromaticity of the solvent appears to b e

important for obtaining high coal conversions at the lowest severity reaction conditions. The coal-

derived soluble products are enriched in hydrogen and depleted in oxygen relative to the starting

coal. Scale-up of the first-stage reaction to a one-liter stirred autoclave was successfull y

accomplished.

Filtration was tested for the removal of unreacted coal and mineral matter from the liquefaction

process. The preliminary conceptual design, which formed the basis for Task 2 work, places a

filter directly following the fi rst-stage solubilization reactor and ahead of a second-stage catalytic

upgrading step. Hot pressure filtration effectively removed solids from first-stage products .

Experimentally determined filtration rates of first-stage products were very fast (in some cases,

>200 kg/m •h); cake resistivities were on the order of 1 x 10 m/kg, and filter cake solids contents2 10

were as high as 90%. The rates of filtration achieved at around 300 °C and 0.3 MPa wit h

feedstocks made from several coals (bituminous, s ubbituminous and lignite) indicate the filtration

in a commercial plant would have only a minor impact (40¢/bbl of oil) on processing cost. Th e

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successful demonstration of filtration of the first-stage product confirmed the viability of filtration

as the primary solids-separation technique.

Catalytic hydroprocessing was used to upgrade the first-stage filtrate. Preliminary upgradin g

studies were performed usin g a conventional two-stage liquefaction distillation resid to establish

conditions and select an appropriate catalyst. Subsequent work was done with 566 C filtrate+

resids of the first-stage products. At 440 C, filtered first-stage products gave significantly higher

resid conversions than a conventional two-stage liquefaction deashed resid (DAR) fro m

Wilsonville Run 258. Hydrogen consumptio n per unit resid conversion was lower for filtrate than

for the DAR. Lignite-derived filtrates gave the highest resid conversion to distillate products.

Concurrent with Task 2, a literature and patent survey was completed on hydride ion reagen t

synthesis under Task 5. The survey was conducted to identify potential transfer agents and their

respective production methods, assess the cost of production of potential transfer agents, and

identify areas requiring further investigation. This survey resulted in the compilation of a n

annotated bibliography (Appendix 1).

The merit of the tested novel concept was assessed by comparing the resul ts of this program with

the results achieved by current, state-of-the-art processing technologies. A major finding of the

economic evaluation based on the experimental findings of Task 2 was that the hydride io n

reagent “A” consumption was the most costly component of the process economics.

TASK 3 - FLOW SHEET DEVELOPMENT

There were three main objectives under Task 3. The fi rst objective was to use the findings of the

experimental program conducted under Task 2, and the engineering and economic evaluation

(Task 5), to refine and revise the original conceptual proc ess design. Thus, a main focus of Task

3 was to address the unfavorable costs determined at the conclusion of Task 2 for the use o f

hydride ion reagent "A". Consequently, Task 3 included the exploration for a low-cost hydride ion

reagent-promoted, high coal conversion, low-severity first-stage reaction system. The criterion

used to screen parametric changes in the first stage was the maintenance of 90 wt % coal

conversion. Economic criteria were then applied to the conditions to determine the feasibility of

further employing the chosen conditions. The reaction system conditions specified were Black

Thunder Mine subbituminous c oal, Reilly Industries anthracene oil, hydride ion source "E", KOH

catalyst, 375 C, and 60 min residence time.

7

A second objective of Task 3 was to integrate the lo w-temperature solubilization stage and solids

separation and investigate the combined system on a larger scale than w as used in the screening

studies under Task 2. Under Task 3, an integrated first -stage and filtration step was successfully

tested. The solids content of the product filtrates were extremely low and the filter cake s

contained almost all of the coal ash. Filtratio n rates were high and filter cakes contained, without

washing, about 50 wt % soluble material.

A third objective of Task 3 was t o scale the design of the catalytic upgrading stage to match that

of the integrated first-stage and solids separation units. The construction and operation of a

larger scale second-stage catalytic upgrading reactor system was successfully completed. The

second-stage reactor system allowed for continuous hydrogen gas addition and product ga s

recovery. As in Task 2, resid conversions for first-stage products were consistently greater than

for conventional two-stage liquefaction resids.

In Task 5 work done concurrently with Task 3, elementally balanced material balance data were

derived from experimental results from first-stage and second-stage tests. An integrate d

liquefaction system (ILS) balance was completed. The ILS included support systems (such as off-

line product up-grading stages). The economic analysis indicates that the cost of production of

refined product (gasoline) via this novel direct liquefaction technology is higher than the cos t

associated with conventional two-stage liquefaction technologies. However, it also wa s

determined that reconfiguring of the conceptual commercial plant co uld result in a cost advantage

over the conventional two-stage liquefactio n process. Figure 2 is a simplified block flow diagram

of the conceptual process configuration.

Based on the economic evaluation of the HI "E" system, it was appare nt that the costs associated

with the first-stage reaction system needed to be significantly reduced. It was recognized that in

order to reduce costs, first-stage operating severity had to be reduced. This could be effected

by reducing reactor residence time, minimizing the amount of e xcess HI "E" reacted, reducing the

steam to coal ratio, reducing the catalyst rate, lowering reactor pressure, lowering solvent-to-coal

ratio, and close-coupling the first- and second-stage reactor systems. Recommendations were

made at the conclusion of Task 3 to experimentally determine the effect of altering thes e

operating parameters. At the conclusion of Task 3, the effect of reducing first-stage operating

severity on coal deoxygenation via CO removal was unknown. The experimental work was to2

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be used to assess the impact of lower first-stage severity on coal deoxygenation via CO 2

removal.

TASK 4 - INTEGRATED FLOW SHEET TESTING

The goals of Task 4 were t o implement the recommendations made at the conclusion of Task 3,

to test the combined first-stage and filtration reaction steps experimentally, and to refine th e

engineering and economic analysis. Upon completion of Task 3, the integrated conceptual flow

sheet was updated. The integrated flow sheet incorporat ed the best alternatives identified among

the first-stage reactants and conditions, deashing methodology, and second-stage catalyst type

and reaction conditions. The integrated flow sheet was tested in a blocked-out or partiall y

integrated manner to validate the conceptual design and identify any areas that need additional

development. Sufficient perfo rmance data were developed so that recommendation concerning

further development (i.e., continuous testing) of th e process concept could be made (see below).

In Task 4, the recommendat ions made at the conclusion of Task 3 were followed to reduce first-

stage severity. These included reducing the catalyst loading. Microautoclave tests with th e

reduced catalyst loading had the same coal conversion (ca. 90 wt %) with a wide range o f

catalyst loadings (45 to 500 mmoles KOH/kg coal). 111 mmole KOH/kg coal was chosen fo r

subsequent tests in the one-liter autoclave. The HI "E" loading also was reduced. This resulted

in reduced coal conversions (between 70 wt % and 80 wt % vs. 90 wt % or higher established

as a requirement in Tasks 2 and 3). The solvent/coal ratio also was reduced; there was a

detrimental effect on coal conversion when solvent/coal ratios were less than or equal to 1.0 .

Additionally, the product mobility was greatly reduced at low solvent/coal ratios. The total system

pressure was greatly reduced by reducing the HI "E" loading. Other work done in the first stage

included the investigation of solvent quality on recycle. Coal conversion was found to b e

unaffected when recycled solvent was used.

Second-stage catalytic upgrading in Task 4 also was devoted to reducing processing cost and

improving resid convers ion. The high hydrocarbon gas and IOM yields generated in the 300 mL

CSTR, which were reported in Task 3, were reduced by system modifications. It wa s

demonstrated that first-stage products, which were filtered and distilled to remove the ligh t

fraction or unfiltered with the -250 C fraction removed, can be upgraded at high levels o fo

conversion. A reduction in catalyst loading was examined. This resulted in a decrease in resid

conversion, but high IOM conversions were maintained. A less-expensive dispersed molybdenum

9

catalyst (Molyvan A) was substituted for Molyvan L. Conversions with Molyvan A wer e

comparable to those obtained with Molyvan L. Resid recycle around the second stage wa s

investigated. It was found that convers ion of resid recycled (once) to the second stage was low.

The solids separation (filtration) step was relocated pos t second-stage. Flow rates were reduced

from >200 kg/m h found at the conclusion of Task 3 for first-stage products to ~10 kg/m h. The2 2

rate was improved to 34 kg/m h by increasing the residence time of the second-stage upgrading2

from 60 to 150 min. The addition of a carbonaceous filter aid increased the rate to ~100 kg/m h.2

At this rate, if further testing shows that removal of a large portion of the distillate solvent could

be recovered without increasing the viscosity substantially, fi ltration costs would be under $1/bbl.

A simplified block flow diagram of the improved liquefaction system is provided in Figure 3.

TASK 5 - ENGINEERING AND ECONOMIC STUDY

The goal of Task 5 was to evaluate the technical and economic merit of the novel concept. To

accomplish this objective, a baseline case was established against which the results of thi s

program were assessed. Initially, the engineering and economic study was based on unpublished

CONSOL experimental results and on the results of a review of the patent and technica l

literature. A baseline was formulated for subbituminous coal, which has cost and availabilit y

advantages over lignite. The baseline integrated two-stage liquefaction (ITSL) case for sub -

bituminous coals was based on the Wilsonville pilot plant operation with subbituminous Blac k

Thunder Mine coal in Run 263 Period J.

As sufficient experimental data became available during the program, elemental and materia l

balances were formulated for e ach of the process steps. Where information was initially lacking

(e.g., second-stage upgrading) , speculative estimates of unit performance were made based on

the performance of similar units in other coal liquefaction processes. In this way, a speculative

flowsheet of the process could be prepared early in the program. This conceptual flowshee t

provided feedback to the Task 2 and 3 experimental work by defining information needs an d

providing guidance for process variable studies.

Upon the completion of Task 2, sufficient information was available to prepare a preliminar y

conceptual commercial plant design. When the preliminary design was completed a preliminary

economic evaluation of the process was made and the results were compared to the baselin e

10

case. Capital and operating costs were estimated and the required p roduct selling price to satisfy

a fixed return on investment was determined using the same approach as in the baseline case.

The design and economic evaluation results were used to guide the experimental work in Task

3 and, eventually, in Task 4. As additional result s of the experimental program became available

during Task 3, the conceptual commercial plant design and the economic evaluation wa s

updated.

Based on the information from Task 4, a final conceptual commercial plant design and economic

evaluation was prepared. Compared to the evaluation prepared at the conclusion of Task 3 ,

improvements in cost of the process were attained via the changes implemented in Task 4 .

However, it was determined at the conclusion of Task 4 that the cost of production of a barrel of

gasoline was $0.86 greater than the Wilsonville base case.

An improved conceptual case developed based on the results of the economic evaluatio n

provides further recommendations for reducing first-stage reaction system cost. This hypothetical

paper study identified the potential improvements i n the economics of the process and can serve

as a guide for future recommended experimental work. The econom ic evaluation of the improved

conceptual case indicates that a significant cos t advantage can be achieved over the Wilsonville

base case provided the performance assumptions are confirmed in future experimental work.

11

CONTRACT CONCLUSIONS AND RECOMMENDATIONS

CONCLUSIONS

The results of the work summarized above confirm the proposed concept. Low-severity coa l

solubilization can be achi eved via a mechanism presumed to be based on hydride ion donation.

High coal conversions can be obtained with coals of three different ranks. The concept is most

economically attractive with subbituminous coal. C oal conversion was found to be dependent on

temperature, residence time, solvent-to-coal ratio, solvent aromaticity, and hydride ion reagent

to coal ratio. React ions could be scaled up by a factor of ten (from 10 g of feed coal to 100 g of

feed coal). First-stage products are depleted in oxygen and enriched in hydrogen as compared

to the feed coal.

Filtration of first-stage products was technically successful for solids removal. Filtration rate s

indicate that if used in a commercial plant, interstage filtration would contribute only $0.40 pe r

barrel of oil to the process co sts. Placement of the filter in the conceptual process is dependent

on the associated capital costs. Consequently, t he final economic evaluation of the process was

made based on two process configurations, one with the filter interstage and one with it pos t

second stage. In Task 4, reduced severity first stage conditions resulted in interstage filtration

being adversely affected. Some improvement in filtration rate was achieved by incorporating a

filter aid. Placement of the filter post second stage was recommended at the conclusion of Task

3. Filtration of second-stage products was diff icult. It is probable that the poor filtration was due,

in part, because of the reduced first-stage s everity. However, increasing the severity of second-

stage conditions by increasing the residence time resulted in improved filtration performance .

The addition of filter aids further improved filtration performance. If further testing shows tha t

removal of a large portion of the distillate solvent could be recovered without increasing th e

viscosity of the product substantially, post second-stage filtration costs would be under $1/bbl.

Catalytic upgrading of the first-stage products was performed on both filtered and subsequently

distilled products and unfiltered, distilled first-stage products. In all cases, resid conversions were

demonstrated to be higher than the base-case Wilsonville materials. Conversions under th e

conditions used reached ~85 wt %. It was shown in preliminary tests that conversion of second-

stage products recycled to the second stage was low, indicating that the conversion reaction is

nearly complete in a single pass through the reactor. Reduction in second-st age reaction severity,

12

such as reduced catalyst loading, resulted in reduced resid conversion, but provided high IOM

and quinoline insoluble (QI) conversions. Increasing residence time increased conversions.

The execution of an economic and engineering evaluation concurrently with the experimenta l

program was an important element in this program. Each step of the conceptual process wa s

evaluated before it was accepted as a potential component of the proposed process. The merit

of the novel concept was assessed by comparing the results of this program to the result s

achieved by current, state-of-the-art liquefaction technologies. It was determined from thi s

comparison that implementation of the novel concept on a commercial scale could produce 3.58

barrels of gasoline per ton of MF coal and would be $0.86/bbl more expensive than th e

Wilsonville base case. The operating costs of the process are sensitive to reagent costs. The

capital costs are dominated by the cost of the CO removal and CO recovery and compression2

systems, and the solids separation (filtration) unit. These operations alone account for 62% of

the increased capital cost as compared to the base case. Another major factor in the increased

cost is the second-stage processing unit, which is required to achieve higher resid conversion

than the base-case second stage. This unit accounts for an additional 24% of the higher capital

costs.

RECOMMENDATIONS

It apparent that the costs associated with the first-stage reaction system need to be furthe r

reduced. The hypothetical paper study developed during Task 3 estimated a required product

selling price $2.50/bbl lower than in the Wil sonville Run 263J base case. Some of the first-stage

parameters estimated to be required to achieve this price included: reactor residence time of 24

min, steam/coal ratio of 0.25/1, KOH catalyst rate of 0.28 wt% on MF coal, and a solvent/coa l

ratio of 1. Experimental work performed under Task 4 can be view ed as the first of several steps

toward these goals to improve the process and to reach the estimated selling price. Cos t

reductions beyond those achieved in Task 4 are anticipated if the first-stage operatin g

parameters more closely approach those suggested in the h ypothetical study without significantly

affecting the extent of coal deoxygenation, product yield and product quality. Therefore, it i s

recommended that a systematic experimental investigation of first-stage parameters be carried

out to assess the extent to which the process can be improved.

The intention of this program was not to study the fundamenta l chemistry of hydride ion-promoted

coal solubilization, but instead to provide a direction for its utilization in the form of a cost -

13

competitive process. In doing so however, a number of interesting observations were made for

which a fundamental study of the chemistry of the system may provide an explanation. One of

these observations is that different coal conversions were obtained with solvents of differen t

aromatic hydrogen contents. What is the mechanism for this effect? The fundamentals of th e

formate ion source chemistry also warrant a study. Questions that could be addressed include:

Is the chemical mechanism the same for different hydride ion sources? Is the mechanism th e

same for a single source at different hydride ion/coal ratios? Are the kinetics of reaction fo r

systems with different hydride ion sources the same once the hydride ion is formed from th e

source chemical?

It is recommended that such fundamental studies be carried out to more fully characterize and

understand the coal dissolution strategy used in this program. The findings of these studies may

help direct the effort to improved process economics.

14

TASK 4

SUMMARY

The findings of Task 4 and the corresponding engineering and economic evaluation carried out

under Task 5 are described here. There was a general objective under Task 4, based on th e

results of work performed under Tasks 2 and 3 and the results of the economic evaluatio n2 3

done concurrently with Task s 2 and 3 under Task 5, to improve process economics by reducing

processing severity. Reaction conditions, process flow design, and subsequent proces s

economics at the conclusion of Task 4 are different than those described at the conclusion o f

Task 3. 3

In Task 4, variables in both first and second stages that could lead to economic improvements

in the process were explored. These variables included the effects of hydride ion source to coal

ratio, catalyst loading, solvent-to-coal ratio, solvent recycle, and first-stage system pressure. The

interstage filter was altered to more closely approximate filtration conditions that would b e

encountered in a commercial pro cess. This included the use of more robust filtration media and

the production of thicker filter c akes. At the conclusion of Task 3, interstage filtration was found

to have a large negative impact on the process economics. Consequently, in Task 4, the first-

and second-stage reactors were (conceptually) close-coupled, removing the interstage filter to

post second stage. In the second stage, process conditions wer e altered to reduce gas yield and

maintain acceptable resid conversion levels. Lowered second-stage catalyst loading also was

tested. The post-second-stage products did not as filter as readily as the first-stage products.

Filtration aids improved filtration; however, filtration for these materials does not show a n

economic advantage over other solids separation techniques (such as critical solvent deashing).

The reduced severity operating con ditions resulted in improved capital costs (from the projected

costs based on work done under Task 3) and a reducti on in the required selling price of a refined

gasoline product. However, as compared to the Wilsonville pilot plant Run 263J base case, the

required selling price of a barrel of gasoline produced by the Novel Concept process is $0.8 6

higher.

15

BACKGROUND

FIRST-STAGE SOLUBILIZATION

The findings of Task 2 supported the concept of first-stage coal dissolution by hydride-ion -

promoted reaction and prepared the way for Task 3 and 4. Low-severity reaction condition s

( 400 C, 60 min) were determined for which high coal conversions (>90 wt %) to solubl e

products were obtained with five differe nt coals of three different coal ranks. The effects on coal

conversion of temperature, residence t ime, coal rank, hydride ion-to-coal ratio, solvent type, and

total reactor charge size were determined. I t was determined that some solvent in the first-stage

reaction is necessary to achieve high coal conversion, although it does not have to be a

hydrogen donating solvent. The aromaticity of the solvent was found to be important for obtaining

high coal conversions at the lowest severity reaction conditions.

Under Task 2, experiments were successfully scaled up from 45 mL microautoclaves to a one-

liter stirred autoclave. A hot transfer of first-stage products from the one-liter autoclave to a

contiguous receiver vessel was successfully completed. This allowed for the subsequen t

integration of a filtering unit with the first-stage reactor under Task 3 and for production of large

quantities of product for subs equent second-stage catalytic upgrading in the 300 mL continuous

stirred tank reactor (CSTR).

At the conclusion of Task 3, several recommendations were made to improve the proces s

economics. These included reducing the costs associated with the first-stage reaction system.

It was determined that first-stage operating severity and cost could be reduced by:

• Reducing reactor residence time

• Minimizing the amount of excess CO shifted

• Reducing the steam to coal ratio

• Reducing the catalyst rate

• Lowering reactor pressure

• Lowering solvent-to-coal ratio

These recommendations formed the basis for Task 4 first-stage work.

FILTRATION

16

In Task 2, filtration of the first stage reactor products derived from a range of coals (bituminous,

subbituminous, and lignite) and solvents (aliphatic and aromatic) and different hydride io n

reagents, prepared under a range of conditions (temperature and residence time), were carried

out at temperatures around 300 C and at pressures of up to 350 kP a. Two principal criteria usedo

for assessing the filter performance were met: 1) filtrates were produced which contained very

low concentrations of mineral matter and insoluble organic matter (IOM) - the undissolve d

products from the dissolution stage, and 2) high filtration rates were achieved regardless of the

conditions used in the first stage. The high flow rates achieved under these conditions indicated

that filtration of these materials would contribute only 40¢/bbl of crude oil equivalent (COE) i n

processing cost (i.e., capital and operating costs) when translated to a commercial-scal e

liquefaction plant. This cost is similar to solid/liquid separation costs in conventional direc t

liquefaction (e.g., ROSE-SR or filtration in an integrated two-stage liquefaction facility ). It was3

concluded at the conclusion of Task 2 that filtration is potentially a viable and economic means

for removing undissolved solids from the first stage reaction products.

The majority of Task 2 work was carried out in small-scale equipment because first-stag e

products were limited in size having been produced in microautoclaves. Although microscal e

filtration gave adequate qualitative data, it was difficult to obtain quantitative data. At the end of

Task 2 and in Task 3, reactor products from the first-stage one-liter autoclave were assessed and

larger solids separation equipment was employed.

In Task 3, hot pressure filtration was demonstrated as an efficient and cost-effective method for

the removal of solids from the first-stage products. From tests with feedstocks produced under

a wide range of conditions in the firs t-stage reactor, it was found that filtration characteristics are

closely related to the conversion of the coal to quinoline and THF solu bles. Rapid filtration occurs

at low conversions, because the integrity of the coal particles is essentially unchanged. A t

moderate conversions, a significant proportion of the coa l is quinoline-soluble/THF-insoluble, and

filtration is slow. At high conversions to both q uinoline- and THF-solubles, filtration is rapid. This

latter desirable condition was achieved using Black Thunder Mine subbituminous coal, Reill y

Industries anthracene oil, hydride ion reagent "E" and KOH catalyst at 375 C for 60 min

residence time in the first stage. Under these conditions the cost of filtration for a commercia l

process was again estimated to be about $0.40/bbl COE.

17

Recommendations were made at the conclusion of Task 3 to change the process operatin g

conditions and the process flow configuration to improve the overall economics of the process.

Close coupling the first and secon d stages could be considered by moving the solids separation

function downstream of the second-stage reactor, avoiding the need to depressurize an d

repressurize process streams between the re actors. The change in system configuration would,

however, negate the advantages of interstage filtration detailed previously. It would no longer2

be possible to recycle in the second-s tage an ash-free resid fraction without rejection of some of

the second-stage catalyst. Movi ng the filtration step downstream of the second stage, the liquid

phase will contain less resid and more distillate, and therefore the loss of resid with the filter cake

will be lower. Consequently, it might be possible to eliminate t he washing stage. These changes

could significantly affect the performance of t he filtration stage and test the viability of filtration as

the best option as a means of removing solids from process streams. These issues formed the

basis for Task 4 solids separation tests.

CATALYTIC UPGRADING

The primary second-stage upgrading objectives in Task 2 were to investigate the upgradin g

capability of the first-stage products, to determine the suitability of the dispersed catalysts, and

to identify a set of reaction conditions that could be used in subsequent larger scale testing. All

tests in Task 2 were carried out in 50 mL microautoclaves. Th e performance of the catalysts was

assessed by determining the conversion of a deashed resid (DAR) from the Wil sonville pilot plant.

The variables investigated included catalyst type and concentration, pretreatment and ru n

temperature, and residence time. It was shown that Molyvan L becomes ac tive within the first few

minutes of initiation of the reaction and that pretreatment to convert the Mo to the active sulfide

form is not required. It also was shown that conversions of the DAR achieved with Molyvan L

were comparable with those obtained with Mo naphthenate, and with a presulfided supporte d

Ni/Mo catalyst with no pretreatment. However, a higher conversion was achieved when th e

supported Ni/Mo catalyst was used following a pretreatment .2

It was shown that the first-stage product prepared from Glenharold Mi ne lignite was more reactive

than the coal-derived deashed resid from the Wilsonville pilot plant. Under severe reactio n

conditions (440 C for 60 min), resid conversion averaged 72%. Under the same conditions, theo

DAR conversion averaged only 31% conversion. Reduced vehicle solvent concentration in the

feed to the hydrotreater was found to give high resid conversion with good efficiency. Simulated

recycle of the dispersed catalyst showed that recycled Mo (from Molyvan L), sulfided in the

18

presence of a coal liquefaction resid, initially exhibit ed activity in recycle comparable to its activity

in the first pass.

In Task 3, large-scale tests wer e carried out in a 300 mL continuous stirred tank reactor (CSTR)

with a continuous flow of hydrogen. This reactor system more closely resembles a commercial

hydroprocessor than did the 50 mL microautoclaves. In Task 3, work continued using bot h

dispersed and supported catalysts. Liquid and solid feedstocks and catalyst precursors wer e

added batch-wise to the CSTR. To sulfide the catalyst, H S addition was effected b y2

decomposition of dimethyl disulfide (DMDS).

At the conclusion of Task 3, it was apparent that second-stage processing costs also required

reduction to make the overall process economically viable. Reducing catalyst cost and reactor

cost (by reducing residence time) were two of the recommendations made with the proviso that

resid conversion not be adversely affected.

ENGINEERING AND ECONOMIC EVALUATION

Task 5 of this contract is the eng ineering and economic evaluation of the test results obtained in

the program. A key element of this assessment is the comparison of this new process concept

with state-of-the-art two-stage liquefaction (TSL) technology. The Wilsonville pilot plant operation

with Black Thunder subbituminous coal was cho sen as the basis for defining state-of-the art TSL

technology. During Wilsonville Run 263, the pilot plant unit was operated in the so-called hybrid

mode with dispersed iron and molybdenum catalysts used in the first reaction stage and a

supported nickel-molybdenum catalyst used in the second stage ebullated bed reactor. Run 263,

material balance period J was chosen as the basis for developing a baseline conceptua l

commercial plant case against which the results of this program could be compared. During each

Task, the experimental results were used to perform an engineering and economic evaluation.

Recommendations based on these evaluations (as detailed above) were used to guide th e

program. The results of this evaluation for Task 4 are provided in Confidential Appendix 2 of this

report.

EXPERIMENTAL

REAGENTS

Coals

19

In Task 4, Black Thunder Mi ne subbituminous coal from Wyoming was used for all experimental

work. Analyses of the coal used is given in Table 1.

Solvents

Anthracene oil was obtained from Reilly Industries; this oil is designated R-AO and was use d

throughout Task 4. Analyses of R-AO are given in Table 2.

Catalysts

First stage. Potassium hydroxide (KOH) was used throughout Task 4.

Second stage. Molyvan L and Molyvan A, commercially available lubricating agents that contain

8% and 29% Mo, respectively, were used as hydrotreating cata lysts throughout Task 4. Molyvan

L is supplied in the form of an oil. Molyvan A is s upplied as a powder. Dimethyldisulfide (DMDS)

was used as an in situ sulfiding agent. Elemental analyses for Molyvan L, Molyvan A, and DMDS

are provided in Table 3. A boiling point distribution for Molyvan L is shown in Table 4.

Feedstocks to Second Stage

Microautoclave Tests. The filtered product from first-stage Run 23h-LA was used as a

feedstock to assess the effect of boiling point distribution upon resid co nversion in microautoclave

tests. Filtrate derived from Run 23h-LA was divided into four parts. Each was vacuum distilled

to provide feedstocks containing 24.2% (23hLAFDR1), 66.1% (23hLAFDR2), 36.3 %

(23hLAFDR3) and 79.0% (23hLAFDR4) material boiling >566 °C (resid). To investigate resid

recycle around the second stage, the products from the first-stage Run 28-LA series (28oLA ,

28aLA and 28bLA) were combined, filtered, and vacuum distilled to give a residue (28oabLAFDR)

containing 82.4% boil ing >566 °C. This material was hydrotreated in the 300 mL CSTR and the

products vacuum distilled to remove the light fraction. The residue containing 70.8% boilin g

>566 °C was used in second stage microautoclave recycle tests.

300 mL CSTR Tests. Feedstocks were prepared from the fi rst-stage Run 28-LA series and Run

41-LA series of tests. The feedstocks from the Run 28-LA series were prepared as described

above. Feedstocks produced from the Run 41-LA series were not filtered prior to use. Th e

<250°C fraction was removed by distillation. Elemental analyses and boiling point distributions

for the feedstocks are given in Tables 19, 20, 25, 26, 31 and 32.

20

EQUIPMENT

First-Stage Reactor System

The first-stage reactor system wa s described in the Task 2 Topical report for this contract. The2

integration of an in-line filter with the reactor system was described in the Task 3 Topical Report. 3

Filtration

The in-line and off-line filters were described in the Task 3 Topical report. Off-line, the 35mm3

diameter pressure filter was used to remove solids from first-stage and second-stage reacto r

products during Task 4. The larger 118 mm diameter unit was used in one test. A filter liner for

the in-line filter was used briefly in Task 4. This liner, constructed of 316 stainless steel, fit s

inside the filter with less than 1 mm wall clearance. The center is machined as a conical funnel

terminating in a 2.54 cm cylinder, thus providing a reduction of filter area from 45.6 cm to 5.12

cm . The liner weighs 2.44 kg; its weight resting on the glass fiber filter paper was considered2

sufficient to prevent filter cake from building under the liner.

Second-Stage Catalytic Upgrading

The 300 mL continuously stirred tank reactor (CSTR) system was de scribed in the Task 3 Topical

Report. Modifications were made to the 300 mL CSTR to ensure that the temperature of th e3

reactants are accurately measured and that the required control temperature of the reactor i s

maintained within established limits. This was achieved by replacing the 6.4 mm thermocouple

well, which was positioned ab ove the level of the agitator, with a 3.2 mm diameter thermocouple

well positioned in line with the base of the agitator. In addition, the stirrer speed was increased

from ~350rpm to ~700rpm to ensure adequate mixing.

PROCEDURES

First-stage Products

Details are provided in the Task 2 and 3 Topical Reports. 2,3

Filtration

Off-Line Filtration. The 35 mm diameter pressure filter containing a GFA glass fiber membrane

was used to remove solids from either first-stage or second-stage reactor products during Task 4.

The larger 118 mm diameter unit wa s used for filtering the combined first-stage reactor products

21

from Runs 28-LA, 28a-LA, and 28b-LA. On-line filtration of these materials was not successful;

the various product streams (filtrate, unfiltered reactor product, and washings from reactor, filter

vessel and receiver) were combined and blended in a stirred vessel at 250 C for 30 minutes.

The total amount recovered for the test represented ~90% of the maximum possible feed to the

on-line filter from the three runs (a small amount of filter cake was excluded since this was not

easily combined with the other components). The mixture was transferred hot to the 118 m m

diameter filter unit for filtration at 250 C and 210 kPa through a GFA membrane.

A cake drying test was carried ou t using filter cake from Run 28-LA. The cake was crushed and

transferred to a glass vessel and vacuum dried at a vapor temperature of 240 C at 2.8 mm Hg

(equivalent atmospheric temperature of 536 C).

Filtration tests were conducted with second-stage products made from unfiltered first-stag e

products was carried out. Filtration temperature was reduced to a value at or below 150 C,

because of the presence of low-boiling components in these materials. In one test, the effect of

a filter aid on filtration rate was examined. A comme rcially available carbonaceous filter aid, BTU

plus, was mixed with first-stage reactor product a t a concentration of ~2.5%. The mixture filtered

through a GFA membrane in the 35 mm unit at 330 C and 210 kPa.

In-line filtration. At the conclusion of the one-liter first-stage autoclave tests, the low-boilin g

reactor contents were flashed into a separ ator vessel. The product remaining in the reactor was

filtered through a stainless steel, six-in ch diameter filter. The filter was located directly under the

one-liter stirred autoclave. Directly under the filter is located a one-gallon receiver. Details o f

these procedures are provided in the Task 3 Topical Report. In order to increase the filter cake3

thickness in this laboratory-scal e apparatus, a filter liner was installed in the stainless steel filter.

The liner provides a reduction of the filter area from 46 cm to 5 cm . Cake drying via N blowing2 22

was accomplished by press urizing the reactor (open to the filter) to 30 psia of N . The pressure2

was allowed to fall and the process repeated several times.

Second-Stage Catalytic Upgrading

Procedures are detailed in the Task 2 and 3 Topical Reports. 2,3

Product Work-Up

First-Stage Products. Details are provided in the Task 2 and 3 Topical Reports. 2,3

22

Second-Stage 300 mL CSTR Tests. The procedure for product work-up was modified. Th e

reaction products in Task 4 were recovered as four separate samples:

(i) Reactor contents - a mixture of the dry scrape from the reactor plus the product recovered

from the reactor walls, stirrer, dip-tube and thermocouple pocket (tetrahydrofuran is used

to wash any remaining material from the reactor and the washings, then distilled to 70°C

at 380mm Hg vacuum to remove the tetrahydrofuran).

(ii) Light distillate - collected in the cooled distillate accumulator.

(iii) Water - collected in the distillate accumulator and separated from the light distillate b y

means of phase separating paper.

(iv) Gases - collected as three samples throughout the run.

The reactor contents and the light distillate were analyzed by proximate and ultimate analyses,

and simulated distillation. Quinoline (QI) and tetrahydrofuran (THFI) insolubles were determined

on the reactor contents. The gas samples were analyzed for h ydrocarbon and heteroatom gases.

23

RESULTS AND DISCUSSION

At the conclusion of Task 3, a set of first-stage conditions (Run 23-LA, Table 5) was chosen to

produce a large quantity of product. This material was subsequently upgraded in the second -

stage catalytic reactor system. Process economics for a conceptual commercial plant wer e

determined based on the experimental data obtained from these tests. It was found that several

key components made the process too costly to be competitive with conventional two-stag e

liquefaction. A set of conditions (Table 6), which if proven viabl e, was formulated that would bring

the costs below those of the Wilsonville Run 263J base case. Task 4 first-stage efforts wer e

devoted to achieving a balance between the goals of reducing costs and maintaining produc t

quality.

FIRST-STAGE TESTS

MICROAUTOCLAVE TESTS

The primary goal of first-stage Task 4 experim ental testing was to reduce the severity of reaction

conditions in order to reduce both the capital and operating costs of the conceptual proces s

derived at the conclusion of Task 3. A suite of microautoclave tests was made in which th e

severity of each cost-contributing factor was reduced independently. Catalyst loading, solvent-to-

coal ratio, and hydride ion to coal ratio were reduced (from the levels used in the Run 23-L A

series of tests that concluded Task 3 experimental work) to determine the effect on coa l

conversion.

Five microautoclave tests were made to examine the effect on coal conversion of lower KO H

catalyst loading (Table 7). Tests were made at 375 C for 60 min with solvent/dry coa l

ratio of 1.8, and catalyst loadings of 111, 83, 65, 55, and 45 mmole/kg dry coal. Coal conver-

sions were 89.9 wt %, 90.6 wt %, 90.3 wt %, 87.9 wt %, and 88.2 wt %, respectively, assuming

the KOH reports to the insolubles. ( There is little difference in these conversions, even if all the

KOH reports to the solubles, because such small amounts of catalyst were used, 0.025-0.010

g.). The results indicate that the first-stage catalyst loading can be reduced from 222 mmole/kg

dry coal used in the one-liter autoclave series (Runs 23-LA) to 111 mmole/kg dry coal withou t

negatively affecting coal conversion. The lower usage rate would result in a cost reduction o f

$1.00 to $1.50 per barrel of gasoline.

Microautoclave tests also were made at re duced solvent/dry coal ratios at 375 C for 60 min with

83 mmole KOH/kg dry coal. (Table 8). The solvent/dry coal ratios used were 1.5, 1.25, and 1 .

24

This is considerably lower than 1.9, which is the ratio used in the one-liter autoclave series 23x-

LA. Coal conversions for these tests are all within 2 wt % (abs.) of 90 wt %. An additional test

(Run 247, Table 9) was made at a solvent/dry coal ratio of 1, but at a reduced KOH loading (50

mmole/kg dry coal). 90 wt % coal conversion w as obtained for this test. This encouraging result

merits further investigation of reduced catalyst loading and solvent to dry coal ratios.

A series of microautoclave tests were made with progressively lower HI "E" loadings. The results

of these tests show a direct correlation of HI "E" loading and coal conversion (Figure 4). A n

economic evaluation of the novel concept process performed concurrent with Task 4 indicates

that first-stage coal conversions can be lower than ca. 90 wt % (the benchmark for first-stag e

conversion throughout this project), if the costs associated with high HI "E" loadings can b e

lowered. The data shown in Figure 4 for microautoclave test s indicate that the HI "E" loading can

be reduced from a value of 1500 to 800-900 (relative scale) while maintaining coal conversions

between 70 wt % and 80 wt %. Confirmation of these data was obtained in one-liter autoclave

tests.

Recycle of first-stage oil was tested in Task 4. The combined distillate (454 C ) of first-stage–

products from one-lite r autoclave Runs 23-, 23a-, 23c-, 23d-, 23f-, and 23g-LA (see below) was

analyzed by GC/MS and for ultimate analyses. The ultimate analyses of the homogenized 23h-LA

recycled oil shows it to be essentially the same as the feedstock Reilly Industries anthracene oil

(Table 2). A precipitate forms in the composite sample at room temperature. Ga s

chromatography/mass spectrometry (GC/MS) of the whole, homogenized sample indicates that

this precipitate is not 'wax', although n-paraffins are evident. A major portion of the precipitate

was identified as phenanthrene. The oil was used in a single microautoclave tests to simulate

recycle around the first stage. The use of the recycled oil did not effect coal conversion under

the conditions used (Table 9).

One-Liter Autoclave Tests

At the conclusion of Tas k 3 , a series of one-liter autoclave tests (Run 23-LA series) was made;

the process economics were based on these tests. In Task 4, one-liter autoclave tests wer e

continued (Table 5). The goals of these Ta sk 4 tests were to reduce processing severity without

reducing coal conversion. Later in the task, this goal was altered to reducing severity, allowing

conversion to fall below the 90 wt % goal that had been established from the beginning of th e

program. The relaxing of the high coal conversion criterion came about from an economically-

25

driven requirement of removing the filter from the interstage position (to remove solids prior t o

catalytic upgrading of the product) to post second-stage. Thus, unreacted coal could be sent to

the second stage to be processed and not rejected in the filter. This approach would b e

acceptable provided additional conversion of the first-stage products could be achieved in th e

second-stage to result in an overall conversion of about 90 wt %. Consequently, Task 4 first -

stage tests were made at significantly lower severity than the Run 23-LA series. Variable s

reduced in severity in these tests included solvent/coal ratios, catalyst loadings, HI "E" loading,

total system pressure, residence time, and water/coal ratio.

Several tests were made prior to relocation of the filter with a different filter configuration than that

used in Task 3. The post first-stage filter diameter was reduced to produce a thicker filter cake

which would more closely simulate commercial operations of the filter.

One-liter autoclave test data are provided on Table 5. Run 24-LA was made at the sam e

conditions as the Run 23-LA series except with half KOH catalyst loading (111 mmole KOH/kg

dry coal vs. 222 mmole KOH/kg dry coal). The mass balance was 100.6%. Coal conversion was

81.7 wt %, considerably lower than the conversion (88.8 wt %) obtained in the correspondin g

microautoclave test (Run 242) made under the same conditions.

A one-liter autoclave test, Run 25-LA, was made at reduced catalyst loading and reduce d

solvent-to-dry-coal ratio. The sample was filtered on-line.

Run 26-LA was designed to test a low solvent-to-dry-coal ratio (1:1). The reactor syste m

developed a leak at the transfer valve to the filter. System pressure approached the valv e

pressure limit of 5000 psi (at 450 F); however, it did not exceed it when the leak occurred .

Measured pressure in the one-liter autoclave was 4900 psi. The run was safely aborted.

Run 27-LA was made with a solvent/co al ratio of 1, and half the catalyst loading used in the Run

23-LA series tests (111 mmole KOH/kg dry coal). Only a small amo unt of product filtered through

the on-line filter; the majority of the product was f ound in the autoclave. Examination of the valve

located between the autoclave and the filter unit sho wed it to be partially blocked. It appears that

any filtration which took place occurred across the partially blocked valve. Gas analyses ,

including direct hydrogen analysis were obtained. All liquid and solid products were recovered

and re-filtered off line.

26

Because of the filtration problems e ncountered in Run 27-LA, Run 28-LA was made at the same

conditions as Run 25-LA (solvent/dry coal = 1.4). This test was made with the conical filter -

diameter reduction-liner in place. The products were found in the filter unit, not in the receiver.

It is not known if the products did not filter beca use of insufficient time or whether too low an over

pressure was used during filtration. All products were collec ted and reserved for off-line work-up.

A duplicate of Run 28-LA (Run 28a-LA) was made. Because it was speculated that the filtration

time had not been long enough in Run 2 8-LA, filtration time for Run 28a-LA was extended to two

hours with an over pressure of N at all times 30 psi. All of the product was found in th e2

receiver; no material was found on the filter. It is not clear at this time how the filter wa s

breached. Run 28b-LA (a duplicate of Runs 28-LA and 28a-LA) was made to test the filter liner

and to provide additional sample for a second-stage hydrotreating test. All products transferred

to the filter unit and approximately 80 g of product filtered though the glass fiber paper and the

Conidur plate. A small dry filter cake and about 50 g of tarry product remained in the filter unit

after one hour of continuous N flow through the filter. It was apparent that the filter liner was not2

providing an appropriate simulation of comm ercial filter application and its use was discontinued.

Run 29-LA was made at the same solvent/dry coal ratio and catalys t loading as Run 28-LA series

tests, however, at a reduced water loading. No added water was fed to the reactor, only th e

inherent ca. 26 wt % moisture of the coal. This resulted in reducing the t otal system pressure from

4325 psi (average of all Run 28-LA tests) to 4170 psi. Coal conversion was found to be only

ca. 80 wt %.

Run 30-LA was made a t the same solvent/dry coal ratio and catalyst loading as Run 29-LA. As

in Run 29-LA, no additiona l water was fed with the 26 wt % moisture coal. In this test, however,

the HI "E" charge was reduced. This reduced the total system pressure at t emperature from 4170

psi obtained in Run 29-LA to 3150 psi. Filtration occurred through the Conidur plate but not the

glass fiber filter paper. This was unintentional and occurred because the filter paper shifted in

assembly of the filter unit. Products found in the receiver were very viscous and, unlike othe r

one-liter autoclave products, did not flow at room temperature. Coal conversion was 75.2 wt %.

This verifies the conversion values obtained in microautoclave tests made with similar HI "E "

loadings (Figure 4). The significant reduction in first-stage system pressure meets the process-

economic goal established at the end of Task 3 (Table 2).

27

Residence time was addressed in the next several tests. Run 31-LA was made under th e

conditions of Run 29-LA. However, res idence time was reduced to 30 min from 60 min. Material

recovery for Run 31-LA was 100.5%. About 19 g of so lid material that contains 19 wt % ash was

recovered from the autoclave. One-liter au toclave test Run 32-LA was made also at 30 min with

a reduced catalyst loading and solvent-to-coal ratio. A solid residue (ca. 8 g) also was found in

the autoclave at the end of the test. This material contains ca. 12 wt % ash. Coal conversion for

both runs was only 50 wt %. This is significantly lower than the approximately 70 wt % specified

to meet the economic goals.

Another test (Run 33-LA) was made at even less se vere conditions than Runs 31-LA and 32-LA.

Residence time was 30 min and the solvent-to-coal ratio was reduced to 1. The products, after

flashing light products, were predominantly solid. They did not transfer to the filter. At th e

reduced solvent-to-coal ratio of 0.8, the residence time in Run 34-LA was extended to 60 min.

Solvent-to-coal and water-to-coal ratios were the same in Run 33-LA as Run 32-LA. However,

the MF coal loading was reduced from 84 g used in the runs described above to 75 g in a n

attempt to reduce the total system pressure. Additionally, the HI "E" to coal ratio also wa s

reduced. Again, the product after flashing was not fluid and did not transfer to the filter. Smal l

aliquots of Run 33-LA and 34-LA products were placed in an oven and reheated to 375 C at

atmospheric pressure. The materials did not melt, become tacky, or fluid. Coal conversion t o

THF solubles was determined to be 32.0 and 36.7 wt % for the two tests, respectively. It wa s

concluded that the reduced solvent-to-coal ratio was responsible for the poor coal conversions

and product quality.

Run 35-LA had a residence time of 30 min. The solvent-to-coal ratio was increased from tha t

used in Run 33-LA to 1.4. The product, which transferred to the receiver, was tarry at roo m

temperature. About 5% of the product was solid and remained in th e autoclave. Coal conversion

to THF-solubles was 52.0 wt %.

In each of the next three tests, the severity of r eaction conditions was increased. Run 36-LA was

made with twice the amount of catalysts as Run 35-LA (100 vs. 50 mmole KOH/kg coal). This

resulted in reducing the amount of material remaining in the one-liter autoclave by half (from ca.

10 to 5 g). Coal convers ion to THF-solubles was 57.1 wt %. Run 37-LA used the same catalyst

loading as Run 36-LA; ho wever, the residence time was increased to 45 min. This resulted in a

reduction of the material remaining in the autoclave to 2.9 g and a coal conversion to THF -

28

solubles of 59.0 wt %. Run 38-LA was made at 60 min resid ence time. However, plugging of the

transfer valve caused the products to be held in the autoclave for an additional 20 min befor e

transfer to the receiver. Appr oximately 25 g of material was found in the autoclave at the end of

the run. Coal conversion was 62.6 wt %, which is higher than that fo r Runs 33-LA through 37-LA,

but the gas yield also appears to be greater.

Run 39-LA was a replicate of Run 38-LA. About 16 g of solid material was left in the autoclave

after transfer of products to the receiver. Coal conversion to THF solubles was 54.6 wt %

In order to produce a fluid product and to determine if the low solvent-to-coal ratio and reduced

residence time were responsible for the prod uction of solid, non-transferred materials, a one-liter

autoclave test, Run 40-LA (a replicate of Run 28-LA), was made. The products of the run were

transferred to the receiver, with the exception of a small amount (<0.7 g) left in the autoclave .

Coal conversion was 84.5 wt %. Several conditions differ between this run and others in which

we were not successful i n producing a totally fluid product. These include water/coal ratio (0.51

in Run 40-LA and 0.35 in all other tests (Runs 31-39)), sys tem pressure (4300 psi), and catalyst

loading (>100 mmole KOH/kg coal). However, Run 30-LA was made at redu ced system pressure

(3150 psi), and an acceptable coal conversion of ca. 75 wt % was obtained. Consequently, a

one-liter autoclave test (Run 41-LA) was m ade to replicate Run 30-LA except that no in-line filter

was used; the products of the run were fluid and transferred to the receiver. Only a small amount

(ca. 1 g) of material was left in the autoclave. Coal conversion was ca. 72 wt %. Examination

of products found in the receiver indicated that there is a 60-65% deoxygenation of the feed coal

under these test conditions. The conditions are intermediate in severity between those used in

the Run 23-LA series made under Task 3 and the goals set for improved economics (Table 2).

It was decided to use the products of Runs 30-LA and 41-LA for second- stage catalytic upgrading

tests. Additional repeat tests (Run 41a-, 41b-, and 41c-LA) were made to provide sufficien t

material for second-stage tests.

Run 42-LA was made at the same conditions as Run 41-LA, but at a slightly higher catalys t

loading. The transfer valve between the autoclave and receiver plugged. The autoclave heat

was immediately reduced and products were recovered from the autoclave. Residence tim e

above 350 C was closer to 70 min than the planned 60 min. The products were reserved, but

not analyzed.

29

FILTRATION

The results of tests made in Task 4, which explored the effects of modifications to the process

operating conditions, were compared to the results achieved in the Run 23-LA series of test s

completed in Task 3 (Tables 10 and 11). The results of filter cake vacuum drying are shown in

Table 12.

In Run 27-LA, the solvent/coal ratio was reduced from 1.8 to 1.0 and the concentration of KOH

catalyst was halved. It was not possible t o transfer all of this material to the on-line filter because

of a partial blockage in the transfer valve. When the recovered material was filtered off-line in the

35 mm unit, virtually no filtrate was obtained. Analysis of the cake showed that its composition

was typical of filter cakes produced previously (47% quinoline insolubles (QI) and 51 %

tetrahydrofuran insolubles (THFI)). These findings support the view that filtration ha d

inadvertently occurred at the partial blockage at the transfer valve, and that the materia l

recovered from the reactor was, in effect, representative of a filter cake.

To overcome the transfer problems associated with the low mobility of the reactor produc t

produced in Run 27-LA, the solvent/coal ratio was increased to 1.4 in the Run 28-LA series of

tests. In these tests, filtration on-line was partially successful (see above). The filter inser t

introduced to improve the quality of the filtration data by reducing the filter area and increasing

filter cake thickness, did not perform as expected and some of the reactor product did not pass

through the filter (see above). When re-filtered off-line in the 11 8 mm filter, the rate achieved with

the recombined Run 28-LA material, ~50 kg/m h with a cake resistivity of ~300 x 10 m/kg,2 10

although acceptable, was significantly lower than the very high rates seen for the reactor products

from the Run 23-LA test series. In those tests, rates of >200 kg/m h with resistivities of <50 x2

10 m/kg were routinely obtained (Table 10). A dry filter cake, which had QI of 42% and THFI10

of 47%, was produced that was about 4 mm thick. This cake cont ained small flecks of filter paper

that had not been removed fro m the feedstocks. Due to the unknown quantity and nature of the

losses in these tests, it is not possible to accurately det ermine the coal conversion achieved from

the mass balance data. However, calculation from the ash balance al lows an estimate of the coal

conversion to be made without recourse to the mass balance data. Coal conversion to quinoline

solubles was 90.1 % and to THF solubles, 88.5 %, on an dry, SO free, ash-free basis (Table3–

11).

30

Characterization of the filtrate showed additional difference s between the products from Runs 28-

LA and 23-LA. The amount of THF insolubles in the filtrate was significantly higher, 6.8% cf ~2%.

Preparation of the feedstock for the second-stage hydrotreating by vacuum distillatio n

demonstrated other differences. Aiming at a 70:30 split (distillate:residue), the Run 28-LA filtrate

was cut at an equivalent atmospheric temperature of 422 C to yield a 36% residue with a

softening point of 230 C. In comparison, a filtrate from Run 23-LA was cut at only 384 C

yielding 33% residue with a softening point of 115 C. These results show that the reacto r

product generated in Run 28-LA contains a higher proportion of high molecular weigh t

components compared to products of Run 23-LA. This is a direct consequence of the les s

severe process conditions used in the first-stage reactor in the Run 28-LA series.

Vacuum drying of the filter cake to an equivalent atmospheric temperature of 536 C yielded

44.6% distillate and 55.4% dry cake. This total recovery exceeded the prediction of ca. 80 %

recovery made during Task 2 of the program. The QI content of the cake rose from 42.4% t o

89.5% and the THFI content increased from 47.1% to 94.2%. Calculation from the mass balance

data predicts that the QI content would be 76.4% and THFI content would be 85.0%. Thi s

suggests that a small amount of the soluble material in the cake had polymerized during th e

heated vacuum drying (Table 12).

The products from Run 30-LA were made under the same conditions as those used for th e

Run 28-LA series, but with the amount of H O and hydride ion "E" reduced from 0.5 to 0.35 and2

from 1.0 to 0.7, respectively, relative to the amount of dry coal. Run 30-LA was filtered on-line,

but it is believed that the product by-passes t he GFA filter. It was re-filtered off-line in the 35 mm

unit. The filtration characteristics were similar to those of Run 28-LA products. The flow rate was

62 kg/m h and the cake resistivity was ~730 x 10 m/kg.2 10

A filtration test on first-stage products of the Run 41-LA series was made with 2.5 wt % filter aid

added to one subsample prior to filtration through a GFA membrane . This was done to determine

whether this commercially accepted practice could improve filtration rates. The filter aid ca n

produce a more porous filter cake with a corresponding enhancement of the filtration rate .

However, its addition also increases the amount of solids in the feed and, hence, the cak e

thickness. There is, therefore, an optimum amount of filter aid that can be added to spee d

filtration. This is dependent upon the charact er and properties of the solids in the reactor product

and the way in which they pack together to form the filter cake. The optimum value was no t

31

determined. However, it was determined from the single test that was carried out, that th e

filtration rate doubled to just over 100 kg/m h, and the cake resistivity was reduced to a value of2

100 x 10 m/kg. The addition of a filter aid, although it produced the desired increase in rate, may10

not be cost effective. The cost of the r aw materials may be prohibitive and the introduction of an

extra process step adds to both the time and the cost of the filtra tion cycle. The addition of a filter

aid increases the amount of solids in the filter cake, which may then retain more of the liqui d

phase. A cost-effective met hod of recovering the valuable liquid product, by solvent washing or

vacuum drying, then becomes more important in that case.

The milder conditions used in the test with filter aid were reflected in t he properties of the resultant

filtrates. The amount of THF insolubles in the filtrate was 15 .2% for Run 30-LA and 14.3% for the

combined sample from Runs 41b-LA and 41c-LA. This is indicative of an increase in th e

concentration of larger molecular species.

The remaining material from Run 41-LA series tests was used as feedstock for catalyti c

upgrading tests made without the preliminary removal of the solids. Relegation of the solid s

separation stage downstream from the second-stage hydroprocessor could have a significan t

affect upon filtration performance. To determine this effect, the filtration characteristics of a

number of samples from the second-stage upgrading reactor were assessed. In tests wit h

differing catalyst concentration and/or type, very low filtration rates were encountere d

(~10 kg/m h). The filtration tests could not be completed, and no mass balance data ar e2

available. However, the reactor product from one test in which more severe hydrocrackin g

conditions were used (residence time of 150 min) was more amenable to solid-liquid separation.

The filtrate rate was 34 kg/m h. This is well below the target value of 100 kg/m h, but represented2 2

a significant improvement. The implications of these data are that solids separation by filtration

of second-stage products can be accomplished, albeit at a cost significantly greater than th e

value predicted at the conclusion of Task 3 for filtration of first-stage products.

SECOND-STAGE CATALYTIC UPGRADING

Microautoclave Tests

Operating conditions and product yields, gas analysis, resid conversion, and hydrogen uptake

for the microautoclave tests that were carried out in Task 4 to investigate a range of proces s

variables are given in Tables 13-15 and Figures 5 and 6. The results from these tests were used

to guide the experimental work in the 300 mL CSTR.

32

Feedstock Residue Concentration. In previous work under Task 3, three first-stage one-liter

autoclave runs were required for each 300 mL CSTR test to provide enough filtered and distilled

product to produce sufficient residue containing typically 70% material boiling >566 °C .

However, the concentration of resid in the feedstock to the sec ond stage had not been optimized.

It was, therefore, considered i mportant to determine the effect of varying the resid concentration

on resid conversion to lighter material (bpt <566 °C). If the conce ntration of resid in the feedstock

could be reduced by lowering the cut point and leaving more light material in the feedstock, a

larger inventory would be produced from few er one-liter first-stage runs. Also, the lower severity

distillation could reduce the potential neg ative impact on the reactivity of the first-stage products.

A series of microautoclave tests was made to determine the optimum resid concentration. The

results (Table 13 and Figu re 7) show that even though the conversion is high (~57%) in the test

when only 36.3% of the feedstock boils >566 °C, the actual amount of residue converted i s

significantly lower than in the two tests made with higher residue concentrations. Conversion is

expressed as the percentage of the feedstock residue that is converted and not the amount of

residue converted. Thus, the use of feedstocks containing low concentrations of resid make s

interpretation of the results more difficult. Therefo re, it was deemed prudent to limit hydrotreating

work to experiments on feedstocks in which the resid concentration was ~70%.

Catalyst Type and Concentration. From an economic stand-point it is important to minimize

the catalyst loading while maintaining a high level of resid conversion. Using Molyvan L, three

microautoclave tests at different catalyst concentrations were carried out with feedstock fro m

first-stage Run 23h-LA. The feedstock, containing 79% resid, was hydrocracked at 440 C foro

60 min. Molyvan L was added to give Mo concentrations of 1% and 0.01%; no catalyst wa s

added in the third test. In addition, a fourth test was conducted in which Molyvan A, a powder

catalyst containing ~28% Mo, was added to give a Mo concentration of 1%. The result s

(Table 14 and Figure 8), show that adding catalyst at the 0.01% level has no measurable affect

on resid conversion. Th e resid conversions were 41% and 38% for the two tests at 0% Mo and

0.01% Mo, respectively. However, when 1% Mo was added there was a significant increase in

resid conversion to 56%. Hydrogen uptakes for the three tests were 21, 29, and 37 mg/g MAF

feed as the catalyst concentration was increased. With Molyvan A added at the 1% Mo level, a

similar conversion and hydrogen uptake were achieved as for the test using Molyvan L at th e

33

same Mo concentration, although with its higher Mo content, less catalyst needs to be added to

give the same Mo concentration.

Second-Stage Residue Recycle. In a commercial scale plant, the resid conversion may b e

improved by recycl ing resid around the second stage, and it may then be feasible to reduce the

catalyst loading with an overall cost adva ntage. However, the success of recycle depends upon

the activity of the unconver ted resid exiting from the second stage, which may prove to be more

refractory and, hence, not responsive to further hydrocrac king reactions. To assess this concern,

microautoclave tests were carried out to compare the conversions achieva ble with first-pass resid

and second-pass recycle resid, and with mixtures of the two, as feedstock. In some tests, n o

supplementary catalyst was added to the reactor, relying purely upon the catalyst recycled with

the resid. In other tests, fresh catalyst was added. The operating conditi ons, product yields, resid

conversion, and hydrogen uptakes are shown in Table 15. The results show the high reactivity

of the first pass resid when 1% Mo (as Molyva n L) was added, giving a resid conversion of 55%.

However, for the second pass resid, with or without fresh catalyst, conversions were very lo w

(<10%), showing the low reactivity of this material. Conversions of the blends of first- an d

second-pass resid materials with the addition of fresh catalyst were measured at 36 to 37% ,

consistent with the conversions achieved with the indivi dual components. However, conversions

of first- and second-pass resid blends without the addition of any fresh catalyst apparentl y

increased to 46.5%. The reason why the addition of fresh catalyst to the bl ended feedstock would

give a lower resid conversion may be an anomaly of the small quantities used and th e

inaccuracies in the microautoclave scale tests.

300 mL CSTR Tests

Under Task 4, 300 mL CSTR catalytic upgrading tests were continued to provide data for th e

economic and engineering evaluation. Several changes to the process conditions wer e

recommended in Task 3, mostly to improve the economic viability of the process. These included

a reduction in catalyst loading, the testing of an alternative dispersed molybdenum catalyst ,

recycling of the unconverted resid, and the processing of ashy feed. After making preliminar y

tests in the microautoclave, 300 mL CSTR runs were made . Tables 16-33 and Figures 8-12 give

the results for all the 300 mL CSTR runs completed in Task 4. The resul ts include details of mass

balances, resid and IOM conversions, elemental balances and analyses, boiling rang e

distributions and feed and product distributions.

34

300 mL System Modifications. Following the implementation of modifications to temperature

measurement and stirrer speed in the 300 mL reactor (see Experimental section), a CSTR run

(263J-2,300CSTR) was conducted using Wilsonvill e Run 263J interstage resid as the feedstock.

The test was conducted at the operating conditions established in Task 3 (440 °C, 60 min, with

1% Mo introduced as Molyvan L). The heat-up time for the run was ~120 min. Compared t o

~270 min previously required, whic h was reduced to ~130 min by the installation of a substantial

amount of additional insulation. It took appro ximately 270 min, with no additional insulation. This

suggests that the temper atures in previous runs were much higher than intended. Although the

resid conversion in Run 263J-2,300CSTR was 46% and was similar to that achieved before the

modifications were made (43%), the re was a substantial reduction in the C -C hydrocarbon gas1 3

yield (3.4% vs 18.4%). There was a corresponding increase of ~17% in the yield of liqui d

products. The hydrogen-to-carbon ratio for the liquid product, 0.99, compare d with 0.79 previously

achieved (0.86 for the Wilsonville Run 263J feedstock), also shows the improved performance

of the CSTR in its current configuration. Similarly, the previously reported increase in IOM from

feedstock to product of 225 % (prior to modifi cations) was not observed. The IOM content of the

product was only 67% as great as th at in the feed, confirming the reduction in product molecular

weight.

Feedstock Type. Three runs were completed using different feedstocks: Wilsonville 263 J

interstage ashy resid, the distillation residue of the filtrate of the combined first-stage product s

from Runs 28-LA, 28a-LA, and 28b-LA (28oabLAFDR) and the combined products from first -

stage Runs 41-LA and 41a-LA (41oaLA). Mass balance, conversion, elemen tal details, simulated

distillation, feed, and product distributions are given in Tables 16-21.

The first run (263J-2,300CSTR) using Wilsonville 263J interstage resid is discussed above. The

second run (28oabLAFDR-1,30 0CSTR) using the combined 28oabLA filtered distillation residue

(FDR) as the feedstock produced high conversions of resi d and IOM material of 71.9% and 77%,

respectively. These values ar e significantly higher than those found in the base-case test using

Wilsonville Run 263J interstage material, where under the same conditions the resid and IO M

conversions were 46.0% and 33%, respectively (Figur e 7). Other indications of high conversion

and high activity of this feedstock include a significantly increased hydrogen uptake from 23 to

50 mg/g MAF feed, and th e elemental and boiling point distributions of the product materials. In

comparison to the results for Wilsonville Run 263J feedstock, the results from hydrotreatin g

28oabLAFDR show a greater reduction in the oxygen and corresponding increase in H/C ratio

35

between feed and products (Table 19). There is a pronounced shift to lower boiling components

for the 28oabLA product s compared to the 263J material. It should be noted that the Wilsonville

Run 263J material is the product of multiple recycling, whereas 28oabLA is a once-throug h

product.

The third run (41oaLA-1,300CSTR) used the combined product of first-stage Runs 41-LA an d

41a-LA, which had been distilled to remove the <250 °C fraction, but had not been filtered. The

low H/C ratio of 0.74, and the high IOM of 17.7% are indicators of th e mild conditions used in their

preparation. However, high conversions for resid and IOM of 72.6% and 74% respectively ,

indicate that these first-stage products have a high reactivity.

Figure 7 shows the resid and IOM contents of the feedstocks before and aft er hydrotreatment and

gives the hydrogen uptake for the three runs. The marked increase in catalytic hydrocrackin g

activity of the hydride ion derived resids over the reference recycled Wilsonville Run 263J resid

is evident. For these feedstocks, the resid conversion is approximately the same as the IO M

conversion. Figure 10 shows the shift in the boiling point distribution reinforcing the indication of

the high activity of the no vel concept feedstocks. All three, however, show the same trend, with

material converted from the >800 °C fraction being principally shifted into the 150 to 450 ° C

boiling range with only small net yields of material boiling in the 450 to 800 °C range.

Catalyst Type and Concentration. Two 300 mL CSTR runs were made to investigate catalyst

loading and type. Mass balance, conversion, elemental data, simulated distillation, feed, an d

product distribution details for the two runs are given in Tabl es 22-27, together with results for the

previously reported run (41oaLA-1,300CSTR), which used 1 % Mo (introduced as Molyvan L).

The operating conditions were the same as for previous tests (440 ° C, 60 min, and 2500 psig H ).2

The first run (41oaLA-2,300CSTR) examined the effect of reducing the Mol yvan L catalyst loading

from 1% Mo to 0.1% Mo. A mass balance closure of 98% was achieved with individual carbon

and hydrogen balances of 103% and 98%, respectively. Conversions of resid and IOM wer e

57.6% and 75%, respectively. C -C gas yield was 2.4% and hydrogen uptake was 38 mg/ g1 3

MAF feed. The second run (41bcLA-1,300CSTR), was made us ing a catalyst loading of 0.1% Mo

introduced as Molyvan A. A mass balance c losure of 99% was achieved. Individual carbon and

hydrogen balances were 100% and 97%, respectively. Conversions of resid and IOM wer e

36

54.9% and 76%, respectively. C -C gas yield was 1.1% and hydr ogen uptake was 35 mg/g MAF1 3

feed.

Figure 8 shows that the resid conversions in the tests made with 0.1% Mo were significantly lower

in both runs than that obtained when a catalyst loading of 1.0% Mo was used. Under the same

operating conditions, but wi th the higher catalyst loading, the resid conversion was 72.6%. This

confirms the microautoclave tests, in which a reduced Mo concentration (0.01% vs. 1.0%) also

resulted in a marked decrease in resid conversi on. Other indicators of lower conversion are the

decrease in hydrogen uptake, reduced oxygen removal, and lower H/C ratio between feed and

products. Surprisingly, the IOM conversion for all three runs was high (~75%) while Q I

conversions for all three tests were ~52%, in dicating that hydrogen addition to the larger species

is occurring to render the majority soluble in THF (and to a lesser amount in quinoline), but that

molecular size is not reduced sufficiently to significantly lower the boiling point. Figure 1 1

highlights the smaller shift of high boiling material for the reduced catalyst loadings.

Residence Time. Two 300 mL CSTR runs were completed in Task 4 to investigate the effects

of residence time on conversion. Mass balan ce, conversion, elemental and simulated distillation

data, feed and product distribution deta ils for the two runs including the results for the previously

reported run (41bcLA-1,300CSTR) us ing a 60 min at 440 °C residence time, are given in Tables

28-33. Molyvan A was used as the catalyst precursor. The operating conditio ns for the tests were

0.1% Mo and 2500 psig H .2

The first run (41bcLA-2,300CSTR) was completed using a residence time of 30 min at 440 °C.

A mass balance closure of 96% was obtained with individual carbon and hydrogen balances of

99% and 96%, respectively. Resid conversion was 42.9%, IOM conversion was 73%, and Q I

conversion was 50%. C -C gas yield was 0.4%. These values compare to 55% resi d1 3

conversion, 76% IOM conversion, 54% QI conversion and a C -C gas yield of 1.1%, for th e1 3

corresponding 60 min residence time test (41bcLA-1,300CSTR).

The second run (41bcLA-3,300CSTR) was made utilizing the product from the 30 min ru n

(41bcLA-2,300CSTR). Following the removal of the reactor product for weighing and sampling,

the material was returned to the reactor and processed for an additional 120 min at 440 C,

250 psig H ; no additional catalyst was added. The original 30 min residence time plus th e2

additional 120 min gave a total residence time of 150 min. Combining the results on a pro rata

37

basis gave a mass balance closure of 93%. Individual c arbon and hydrogen balances were 97%

and 89%. Resid conversion was 73.1%, IOM conversion was 86%, QI conversion was 78%, C -1

C gas yield was 2.6%. The conversion obtained is very similar to the 72.6% achieved whe n3

operating conditions of 440 °C, 2500 psig, a residence time of 60 min and a catalyst loading of

1% Mo, were used (Run (41bcLA-2,300CSTR). Figure 9 gives hydrogen uptake, C -C gas1 3

make, resid, IOM and QI conversions. The results suggest that if the longer residence time in

the reactor can be equated to recycle around a loop with short residence time, then in a system

with second stage resid recycle, resid conversions of >70 % can be achieved. However, this was

not found to be the case in the earlier microautoclave tests. Figure 12 shows a greater shift of

heavy to light boiling material, with an increase in residence time.

CONCLUSIONS AND RECOMMENDATIONS

In Task 4, changes to the process conditions were introduced to improve the overall economics

of the process. The effects on co al conversion, filtration characteristics, and catalytic upgrading

reactor were examined. The se verity of the conditions employed in the first stage were reduced

by decreasing the concentration of hydride ion source, catalyst loading, system pressure, an d

solvent-to-coal ratio. Reduction of severity in the first stage resulted in the inter-stage filtration

being adversely affected. Some improvement in filtration rate was achieved by the incorporation

of a filter aid. Repositioni ng the filter to a position downstream of the second-stage hydrotreater

resulted in a marked decrease in filtration rate. This was partially allev iated by the addition of filter

aids and increasing the severity of the conditions in the second stage (by increasing th e

residence time). At this filtration rate, if it can be shown that removal of a large portion of th e

distillate solvent could be recovered without increasing the viscosity too much, filtration cost s

would be under $1/bbl (Confidential Appendix 3).

In the second-stage 300 mL CSTR, filtered resid and unfiltered, first-stage product conversions

were consistently demonstrated to be higher than in tests made with the base case Wilsonville

Run 263J material. On recycle, second-stage resid product conversions were low. The impli-

cation of this is that the reaction on the first pass converted near th e maximum amount of material

that is convertible under the second-stage conditions used in these tests.

The reduced severity operating con ditions resulted in improved capital costs (from the projected

costs based on work done under Task 3) and a reducti on in the required selling price of a refined

gasoline product. However, as compared to the Wilsonville pilot plant Run 263J base case, the

38

required selling price of a barrel of gasoline produced by the Novel Concept process is $0.8 6

higher. It is, therefore, concluded that in order to be a viable alternative liquefaction process ,

costs need to be further reduced. It is likely that further c ost reductions for this process can occur

if the first-stage operating parameters app roach those suggested in the hypothetical paper study

performed at the conclusion of Task 3 that was used to guide the work of Task 4 (Table 6) .

However, cost reductions must be achieved without significantly affecting the extent of coa l

deoxygenation, product yield, and product quality. The costs associated with the first-stag e

reaction system can be reduced by reducing the following operating parameters: reacto r

residence time, steam/coal rati o, catalyst loading, and further reducing the solvent-to-coal ratio.

The cost of hydride ion source regeneration also must be addressed. Therefore, it i s

recommended that a systematic experimental investigation be carried out to assess the extent

to which the Novel Concept process can be improved.

39

REFERENCES

1. Burke, F. P.; Winschel, R. A.; Brandes, S. D.; Lancet, M. S.; Derbyshire, F. J.; Kimber ,G. M.; Anderson, R. K.; Carter, S. D.; Peluso, M. "Exploratory Research on Novel Coa lLiquefaction Concept, Management Plan", DOE/PC 95050, August 1995.

2. Brandes, S. D.; Winschel, R . A.; Derbyshire, F. J.; Kimber, G. M.; Anderson, R. K.; Carter,S. D.; Rantell, T. D.; Vego, A.; Peluso, M. "Exploratory Research on Novel Coal Liquefac-tion Concept, Task 2 - Evaluation of Process Steps, Topical Report", DOE/PC 95050-22,May 1997.

3. Brandes, S. D.; Winschel, R. A.; Derbyshire, F. J.; Kimber, G. M.; Anderson, R. K. ;Jacques, D. N. ; Rantell, T. D.; Peluso, M. "Exploratory Research on Novel Coal Liquefac-tion Concept, Task 3 - Flow Sheet Development - T opical Report", DOE/PC 95050-59, May1998.

4. S. A. Moore. "Filtration of CSTL Atmospheric Still Bottoms", British Coal Report to HR Iunder DOE Contract. Report No ECR 001 July 1991.

40

TABLE 1

ANALYSES OF BLACK THUNDER MINE SUBBITUMINOUS COAL

Ultimate, wt % Dry

C 70.32

H 4.68

N 1.04

S 0.50

O (by diff) 17.89

SO -free ash 5.543

Major Ash Elementals

SiO 31.482

Al O 15.762 3

TiO 1.142

Fe O 5.482 3

CaO 21.34

MgO 4.30

Na O 0.482

K O 0.492

P O 0.962 5

SO 17.263

Und 1.31

Moisture, % 22.40

41

TABLE 2

ANALYSES OF REILLY INDUSTRIES ANTHRACENE OIL AND RECYCLED DISTILLATE FROM RUN 23-LA SERIES

Ultimate Analysis, wt % Anthracene OIl 23-, 23a-, 23c-, 23d-, 23f-, 23g-LAReilly Industries Composited Distillate from Runs

C 91.21 90.00

H 5.77 6.41

N 1.00 0.83

S 0.55 0.52

O (by diff) 1.47 2.24

TABLE 3

ANALYSES OF MOLYVAN L, MOLYVAN A, AND DIMETHYLDISULTFATE

Ultimate Analysis, wt % Molyvan L Molyvan A DMDS

C 50.9 29.5 27.1

H 9.3 5.1 6.3

N 1.4 2.5 1.8

S 11.7 20.0 58.1

O (by diff) 1.2 0.9 6.8

Ash 25.6 42.0 0.0

42

TABLE 4

BOILING RANGE DISTRIBUTION OF MOLYVAN L

Boiling Point Range, C Molyvan Lo

<150 0.0

151-200 0.0

201-250 0.3

251-300 1.7

301-350 17.4

351-400 25.7

401-450 10.1

451-500 5.2

501-550 0.0

551-600 0.0

601-650 0.0

651-700 0.0

701-750 0.0

751-800 0.0

>800 39.6

Total 100.0

>566°C 39.6

TABLE 5

TASK 4 - FIRST-STAGE ONE-LITER AUTOCLAVE TESTS SUMMARY

Black Thunder Mine Subituminous Coal; HI "E"; 375 C

Run 23 24 25 26 27 28 29 30 31 32 33 34 35 36 37 38 39 40 41 42b,e a b b

Coal Loading, g 100 100 114 100 108 114 114 113 114 114 101 101 101 101 104 101 114 114 114 114

Reaction Time, 60 60 60 60 60 60 60 60 30 30 30 60 30 30 45 80 60 60 60 70min

System 4100 4250 4375 N/A 4600 4325 4170 3150 4170 4020 3650 3270 3620 3560 3590 3580 3570 4300 3000 3140Pressure, psi

HI "E"/dry Coal 1.0 1.0 1.0 1.0 1.0 1.0 1.0 0.7 1.0 1.0 0.75 0.60 0.75 0.75 0.75 0.75 0.75 10 0.70 0.70

Solvent/dry 1.9 1.9 1.4 1.0 1.0 1.4 1.4 1.4 1.4 1.2 1.0 0.8 1.4 1.4 1.4 1.4 1.4 1.4 1.4 1.4Coal, g/g

Water/dry Coal,g/g 0.6 0.6 0.7 0.4 0.4 0.5 0.3 0.4 0.35 0.35 0.35 0.08 0.35 0.35 0.35 0.35 0.35 0.50 0.35 0.35

KOH, mmoles/kg dry 202 111 111 111 111 111 111 111 50 50 50 40 50 100 100 100 100 111 111 170coal

Coal Conversionto THF Solubles, 85.5 N/A N/A N/A N/A N/A 79.4 72 49.0 50.0 32.0 36.7 52.0 57.0 59.0 62.6 54.6 84.5 72 N/Awt%

c d d

(a) not completed(b) multiple runs made under these conditions and run number (c) N/A = not available(d) approximate(e) made under Task 3

44

TABLE 6

TASK 4 - FIRST-STAGE SYSTEM-CONDITION GOALS

Case Established at Conclusion of GoalsTask 3 (Run 23-LA series) Improved Novel Concept Case

Interstage Solids Separation Post Second-Stage Separation

Coal Conversion: 90 wt % or better Coal Conversion: 70-80 wt % acceptable

Residence Time: ca. 60 min Residence Time: ca. 24 min

Temperature: 375 C Temperature: 375 C

System Pressure: ca. 4100 psi System Pressure: ca. 3000 psi

HI 'E"/MF coal = 1.0 (relative ratio) HI "E"/MF coal = 0.37 (relative ratio)

H O/MF coal, g/g = 0.64 H O/MF coal, g/g = 0.252 2

Catalyst: 202 mmoleKOH/kg MF coal Catalyst: 50 mmoleKOH/kg MF coal

Solvent: Reilly anthracene oil Solvent: Reilly anthracene oil

Solvent/MF coal, g/g =1.9 Solvent/MF coal, g/g = 1.0

TABLE 7

FIRST-STAGE MICROAUTOCLAVE TESTS EFFECT OF KOH CATALYST LOADING ON COAL CONVERSION

Black Thunder Mine Subbituminous CoalReilly Industries Anthracene Oil, HI "E", 375 C, 60 minHI "E"/coal = 1, Solvent/coal =1.8, g/g

Run No. Catalyst, mmole/kg dry coal Coal Conversion, wt %

228 500 92.0

237 222 89.5

238 111 89.4

241 83 90.4

244 83 90.3

239 55 87.7

240 45 88.2

45

TABLE 8

FIRST-STAGE MICROAUTOCLAVE TESTS AT REDUCED SOLVENT/DRY COAL RATIOS

Black Thunder MIne Coal375 C, 60 min

Run No. dry coal dry coal wt %Solvent / Catalyst , mmole/kg Coal Conversion, a b

242 1.5 83 88.8

245 1.25 83 92.1

246 1.0 83 88.4

247 1.0 50 89.4

a. solvent: Reilly Industries anthracene oilb. catalyst: KOH

TABLE 9

FIRST-STAGE MICROAUTOCLAVE TESTS - EFFECT OF RECYCLE SOLVENT ON COAL CONVERSIONREILLY INDUSTRIES ANTHRACENE OIL vs.

RECYCLED DISTILLATE FROM RUN 23-LA SERIES

Black Thunder Mine Subbituminous CoalHI "E", 375 C, 60 min

Run Solvent/ dry coal, KOH catalyst, Coal Conversion, No. Oil g/g mmole/kg coal wt %

247 Fresh 1.00 50 89.4

248 Recycled 1.00 50 89.2

TABLE 10

TASK 4 OFF-LINE FILTRATION TESTS

Filter membrane = GFA

Run No. No. Filter feed ( C) (mPa s) (%) (kPa) (m/kg x 10 ) (kg/m )NCF Temp Viscosity feed Pressure resistivity 30min.

o

Solids in Cake after

10

Total flow

2

n/a 70 Coal/oil slurry 100 1.0 15 210 4 320

23g LA 71 First stage product 250 0.8 4.6 210 34 225

27LA 72 First stage product 305 0.4 65 410 n/m (a) n/m

28LA 73 First stage product 245 1.8 6.6 210 320 50

30LA 74 First stage product 320 0.3 7.2 220 730 62

41bcLA (b) 79 First stage product 330 0.6 10.8 210 100 102b

41oaLA-1&2 75 Second stage product 150 1.0 9.1 210 4490 11

41bcLA-1 77 Second stage product 115 2.0 5.2 210 5530 11

41bcLA-2cont 78 Second stage product 130 1.0 5.4 210 1080 34

(a) n/m = not measured.(b) includes 2.5% filter aid

TABLE 11

MASS BALANCE DATA

Run No.Yields, % % % Coal Conversion, % daf enrich., % daf

Cake Insolubles, Solids in Feed, from ashd

CoalConversion

Dist. Filtrate Cake Total THFI QI THFI QI (1) (2) (3) (4) (2) (4)

23o LA 0.8 86.1 11.5 98.4 45.7 43.4 7.9 5.3 87.3 93.8 93.8 94.6 90.4 91.2

23g LA 4.9 84.4 9.2 98.5 42.1 40.8 7.6 4.6 86.1 96.1 94.4 96.4 90.5 91.0

27LA 0.2 11.8 88.0 100 58.6 52.0 50.5 47.3 n/m n/m n/m n/m 41 48a b c c

28LA 0.7 80.2 15.5 96.4 47.1 42.4 7.3 6.6 n/m n/m n/m n/m 88.5 90.1

30LA 4.6 74.7 18.4 97.7 45.4 40.2 17.0 7.2 n/m n/m n/m n/m - -(a) n/m = not measured(b) by difference(c) estimate(d) Coal conversions determined on an SO free basis from: (1) THF insolubles in reactor product (2) THF insolubles in the filter cake (3) Quinoline insolubles3

reactor product (4) Quinoline insolubles in the filter cake.

TABLE 12

FILTER CAKE VACUUM DRYING

Feedstock: filter cake from NCF73 (28LA)Conditions: 240 C at 2.8mmHg (equivalent atmospheric temp of 536 C)o o

Filter cake composition, % Yields, % Dry cake composition (calc), % Dry cake composition (meas), %

QI THFI dry cake distillate QI THFI QI THFI

42.4 47.1 55.4 44.6 76.5 85.0 89.5 94.2

TABLE 13

MICROAUTOCLAVE HYDROTREATING TESTS ON DISTILLATION RESIDUES FROM RUN 23h-LA FILTRATE

Run No. 50MA-1 50MA-2 50MA-3 50MA-4 50MA-5 50MA-6 50MA-7

Feedstock DR1 DR1 DR1 DR2 DR2 DR2 DR3 DR3 DR3 DR4 DR4a

Feedstock, g - 2.00 2.00 - 2.00 2.00 - 2.00 2.00 - 2.00

Catalyst precursor - Molyvan L Molyvan L - Molyvan L Molyvan L - Molyvan L Molyvan L - Molyvan L

Reaction time, min - 60 60 - 60 60 - 60 60 - 60

Reaction Temperature, C 440 440 440 440 440 440 440o

Charge Pressure,MPa - 10.1 10.1 10.1 10.1 10.1 10.1 10.1H : H S 98 : 2 98 : 2 98 : 2 98 : 2 98 : 2 98 : 2 98 : 22 2

Mo, % feed - 1.00 1.00 - 1.00 1.00 - 1.00 1.00 - 1.00

Products, wt % MAF feed

C -C gas - 4.0 3.7 - 5.4 5.5 - 5.1 4.3 - 6.21 3

C + vapor - 2.0 2.1 - 3.5 3.6 - 4.7 2.8 - 2.74

CO + CO - 0.2 0.2 - 0.5 0.5 - 0.2 0.2 - 0.72

< 566 C % 75.8 79.3 77.2 33.9 61.2 59.3 63.7 74.5 74.6 21.0 54.6 o

> 566 C % 24.2 14.5 16.8 66.1 29.4 31.1 36.3 15.5 18.1 79.0 35.8 o

Total 100 100 100 100 100 100 100 100 100 100 100

Derived values

Resid conversion, - 40.1 30.6 - 55.6 53.0 - 57.2 50.1 - 54.7> 566 C wt % MAF feed o

H uptake, mg/ g MAF feed - 27.0 24.6 - 35.4 32.2 - 39.7 35.3 - 37.02

(a) Feedstocks are distillation residues (DR) of filtrate (F) of Run 23h-LA first-stage product

49

TABLE 14

MICROAUTOCLAVE CATALYST TESTS ON THE DISTILLATION RESIDUE OF RUN 23h-LA FILTRATE

Run No. 50MA-7 50MA-8 50MA-9 50MA-10

Feedstock 23hLAFDR4 23hLAFDR4 23hLAFDR4 23hLAFDR4 23hLAFDR4a

Feedstock, g - 2.0024 2.0030 2.0047 2.0010

Catalyst precursor - Molyvan L Molyvan L - Molyvan A

Reaction time, min - 60 60 60 60

Reaction temperature, C 440 440 440 440o

Charge Pressure, MPa - 10.1 10.1 10.1 10.1H : H S 98 : 2 98 : 2 98 : 2 98 : 22 2

Mo, % feed - 1.00 0.01 0.00 1.00

Products, wt % MAF feed

C -C gas - 6.2 8.0 6.3 5.71 3

C + vapor - 2.7 2.5 2.0 3.34

CO + CO - 0.7 2.0 1.1 0.72

< 566 C, % 21.0 54.6 38.3 43.6 55.3 o

> 566 C, % 79.0 35.8 49.1 47.0 35.0 o

Total 100 100 100 100 100

Derived values

Resid conversion,> 566 C wt % MAF feed - 54.7 37.9 40.5 55.7 o

H uptake, mg/g MAF feed - 37.0 29.4 21.0 40.32

(a) 23hLAFDR4 = Fourth aliquot of the distillation residue (DR) of the filtrate (F) of first-stageproducts from Run 23h-LA

TABLE 15

MICROAUTOCLAVE TESTS OF RESIDUE RECYCLE

Run No. R4 Residue 50MA-7 50MA-11 50MA-12 50MA-13 50MA-14 50MA-15 50MA-16 50MA-17

1st Pass FeedstockFeedstock 28oabLAFDR23hLAFD 300CSTR

Second Pass

First Pass Feedstock, g - - 2.0024 - - - 1.2010 1.2027 1.2025 1.2028

Second Pass Feedstock, g - - - 2.0009 2.0029 2.0042 0.8034 0.8021 0.8008 0.8015

Catalyst precursor - - Molyvan L Molyvan L Molyvan L - Molyvan L Molyvan L - -

Reaction time, min - - 60 60 60 60 60 60 60 60

Reaction Temperature, C 440 440 440 440 440 440 440 440o

Charge Pressure, MPa - - 10.1 10.1 10.1 10.1 10.1 10.1 10.1 10.1H : H S 98 : 2 98 : 2 98 : 2 98 : 2 98 : 2 98 : 2 98 : 2 98 : 22 2

Mo, % feed - - 1.0 1.0 1.0 - 0.6 0.6 - -

Products, wt% MAF feed

C -C gas - - 6.2 4.1 3.8 3.5 5.0 5.4 4.7 5.21 3

C + vapor - - 2.7 3.6 2.5 2.5 3.1 5.3 2.7 2.14

CO + CO - - 0.7 0.0 0.0 0.0 0.2 0.2 0.2 0.22

< 566 C- % 21.0 29.2 54.6 17.3 28.1 24.8 44.0 41.0 52.2 50.0 o

> 566 C+ % 79.0 70.8 35.8 75.0 65.6 69.2 47.7 48.2 40.4 42.5 o

Total 100 100 100 100 100 100 100 100 100 100

Derived values

Resid conversion, - - 54.7 -5.9 7.4 2.3 37.0 36.4 46.7 46.2> 566 C wt % MAF feed o

H uptake, mg/g MAF feed - - 37.0 10.0 12.3 15.7 28.0 34.1 25.8 34.02

51

TABLE 16

MASS BALANCE DATA FOR FEEDSTOCK TESTS, 300 mL CSTR RUNS

Run Conditions: 440 C, 60 min @ 17.0 MPa H .2

Catalyst: 1% Mo in Molyvan L

Run No. 263J-2,300CSTR 1,300CSTR 41oaLA-1,300CSTR28oabLAFDR-

Feedstock 263J Interstage Resid 28oabLAFDR 41oaLA>250 C0

wt In, g

Feed 80.07 96.42 97.78

Catalyst 9.89 11.85 12.27

DMDS 7.42 6.89 6.15

H 22.49 23.17 26.322

Total In (g) 119.87 138.33 142.52

Wt Out (g)

Reactor Contents 70.95 92.56 91.84

Light Distillate 10.12 8.03 15.25

Water 3.41 6.95 1.50

C -C 's 3.38 7.26 3.151 3

C4 0.86 1.65 0.40+

CO+CO 0.27 0.10 0.112

H S 3.28 2.36 1.542

H 20.92 18.42 21.902

Total Out 113.19 137.33 135.74

% Mass Balance 94.43 99.28 95.24

TABLE 17

CONVERSION DATA FOR FEEDSTOCK TESTS, 300 mL CSTR RUNS

Run Conditions : 440 °C, 60 min @ 17.0 MPa H , 1% Mo as Molyvan L 2

Run No. Feedstock feed feed MAF feed feed THFI %

>566°C in >566°C H uptake IOMfeedstock, resid conv., C -C gas by diff., IOM in conversionwt% MAF wt% MAF yield, wt% mg/g MAF feedstock THFI

1 3

a

2

b

c

263J-2, 263J 66.5 46.0 1.2 23 16.6 33300CSTR Interstage Resid

28oabLAFDR-1, 28oabLAFDR 82.4 71.9 5.2 50 17.1 77300CSTR

41oaLA-1, 41oaLA 32.8 72.6 1.1 35 16.3 74300CSTR

a. Excludes C from DMDS1

b. Includes the hydrogen associated with the decomposition of DMDS.c. THFI wt % MAF feed .

TABLE 18

OVERALL ELEMENTAL BALANCE FOR FEEDSTOCK TESTS, 300 mL CSTR RUNS

Run No. C, wt % H, wt % N, wt % S, wt % O, wt % diff ash, wt % (SO free) Total, wt %3

263J-2,300CSTR 94 99 81 85 98 93 94

28oabLAFDR-1,300CSTR 102 96 97 61 107 108 99

41oaLA-1,300CSTR 102 97 81 49 46 85 95

53

TABLE 19

ELEMENTAL ANALYSES FOR FEEDSTOCK TESTS, 300 mL CSTR RUNS

Run No 263J-2,CSTR

C, H, N, S, O, wt % Ash, H/Cwt % wt % wt % wt % (diff) wt % Total (atomic)

IN

263J Resid. 73.8 5.3 1.1 1.1 4.3 14.4 100.0 0.86

Molyvan L 50.9 9.3 1.4 11.7 1.2 25.6 100.0 -

DMDS 27.1 6.3 1.8 58.1 6.8 0.0 100.0 -

OUT

Reactor Contents 71.5 5.9 1.0 2.4 0.8 18.4 100.0 0.99

Light Distillate 78.1 12.5 0.8 4.1 4.8 0.0 100.0 1.92

Run No 28oabLAFDR-1,300CSTR

C, H, N, S, O, wt % Ash, H/Cwt % wt % wt % wt % (diff) wt % Total (atomic)

IN

28oabLAFDR 86.1 5.9 1.4 0.3 5.2 1.1 100.0 0.82

Molyvan L 50.9 9.3 1.4 11.7 1.2 25.6 100.0 -

DMDS 27.1 6.3 1.8 58.1 6.8 0.0 100.0 -

OUT

Reactor Contents 85.1 7.8 1.2 0.8 0.3 4.8 100.0 1.10

Light Distillate 82.6 11.7 1.3 0.2 4.1 0.1 100.0 1.70

Run No 41oaLA-1,300CSTR

C, H, N, S, O,wt % Ash H/Cwt % wt % wt % wt % (diff) wt % Total (atomic)

IN

41oaLA 84.6 5.2 1.1 0.7 5.8 2.6 100.0 0.74

Molyvan L 50.9 9.3 1.4 11.7 1.2 25.6 100.0 -

DMDS 27.1 6.3 1.8 58.1 6.8 0.0 100.0 -

OUT

Reactor Contents 83.8 8.0 1.0 0.9 1.3 5.0 100.0 1.15

Light Distillate 82.4 10.6 1.3 3.2 1.8 0.7 100.0 1.54

TABLE 20

BOILING RANGE DISTRIBUTION FOR FEEDSTOCK TESTS, 300 mL CSTR RUNSwt %, AS DETERMINED

Boiling CatalystPoint

Range, Co

263J-2,300CSTR 28oaLAFDR-1,300CSTR 41oaLA-1,300CSTR

Feed Products Feed Products Feed Products

263J Reactor Light 28oabLA Reactor Light Reactor LightResid Contents Distillate FDR Contents Distillate 41oaLA Contents Distillate Molyvan L

<150 0.0 1.3 49.5 0.0 5.4 37.6 0.0 1.4 41.2 0.0

151-200 0.0 0.4 15.8 0.0 4.1 15.4 0.3 2.4 10.4 0.0

201-250 0.0 0.6 2.1 0.1 3.4 10.5 1.9 2.5 14.5 0.3

251-300 0.0 1.0 6.7 0.1 4.6 11.8 2.6 5.6 11.5 1.7

301-350 0.7 2.6 8.4 0.1 8.3 13.0 15.0 22.6 12.6 17.4

351-400 0.2 4.6 8.4 0.2 11.4 8.3 22.7 26.1 6.1 25.7

401-450 2.1 9.1 5.5 2.1 13.0 3.4 13.7 13.1 2.3 10.1

451-500 10.1 12.9 3.6 8.8 11.8 0.0 6.4 6.3 1.4 5.2

501-550 12.1 11.6 0.0 4.8 7.3 0.0 2.3 3.7 0.0 0.0

551-600 9.8 9.1 0.0 4.4 6.2 0.0 1.7 3.2 0.0 0.0

601-650 7.2 6.2 0.0 3.7 5.0 0.0 1.3 2.5 0.0 0.0

651-700 4.9 4.0 0.0 2.7 3.6 0.0 0.3 1.7 0.0 0.0

701-750 2.6 1.5 0.0 1.6 1.8 0.0 0.0 0.0 0.0 0.0

751-800 1.6 0.0 0.0 2.1 0.3 0.0 0.0 0.0 0.0 0.0

>800 48.7 35.1 0.0 69.3 13.8 0.0 31.8 8.9 0.0 39.6

Total 100.0 100.0 100.0 100.0 100.0 100.0 100.0 100.0 100.0 100.0

>566°C 71.7 53.0 0.0 82.4 28.7 0.0 34.6 15.3 0.0 39.6

55

TABLE 21

FEED AND PRODUCT DISTRIBUTION FOR FEEDSTOCK TESTS,300 mL CSTR RUNS

Run Conditions : 440 °C, 60 min, 17.0 MPa H and 1% Mo as Molyvan L2

Run Number 263J-2,300CSTR 28oabLAFDR,300CSTR 41oaLA-1,300CSTR

Feed In Product Out Feed In Product Out Feed In Product Out

(g) (%) (g) (%) (g) (%) (g) (%) (g) (%) (g) (%)

Feed 263J resid 28oabLAFDR 41oaLA

Feed Wt 80.07 66.8 96.4 69.7 97.8 68.6

Catalyst Molyvan L Molyvan L Molyvan L

Catalyst Wt 9.89 8.3 11.9 8.6 12.3 8.6

Boiling Pt. Range, °C

IBP-150 0.0 0.0 5.9 5.2 0.0 0.0 8.0 5.8 0.0 0.0 7.6 5.6

151-200 0.0 0.0 1.9 1.7 0.0 0.0 5.0 3.6 0.3 0.2 3.8 2.8

201-250 0.0 0.0 0.6 0.5 0.1 0.1 4.0 2.9 1.9 1.3 4.5 3.3

251-300 0.2 0.2 1.4 1.2 0.3 0.2 5.2 3.8 2.8 2.0 6.9 5.1

301-350 2.3 1.9 2.7 2.4 2.2 1.6 8.7 6.3 16.8 11.8 22.7 16.7

351-400 2.7 2.3 4.1 3.6 3.2 2.3 11.2 8.2 25.3 17.7 24.9 18.3

401-450 2.7 2.3 7.0 6.2 3.2 2.3 12.3 9.0 14.6 10.2 12.4 9.1

451-500 8.6 7.2 9.5 8.4 9.1 6.6 10.9 7.9 6.9 4.8 6.0 4.4

501-550 9.7 8.1 8.2 7.2 4.6 3.3 6.8 5.0 2.2 1.5 3.4 2.5

551-600 7.8 6.5 6.5 5.7 4.2 3.0 5.7 4.2 1.7 1.2 2.9 2.1

601-650 5.8 4.8 4.4 3.9 3.6 2.6 4.6 3.4 1.3 0.9 2.3 1.7

651-700 3.9 3.3 2.8 2.5 2.6 1.9 3.3 2.4 0.3 0.2 1.6 1.2

701-750 2.1 1.8 1.1 1.0 1.5 1.1 1.7 1.2 0.0 0.0 0.0 0.0

751-800 1.3 1.1 0.0 0.0 2.0 1.4 0.3 0.2 0.0 0.0 0.0 0.0

>800 (*) 42.9 35.8 24.9 22.0 71.5 51.7 12.8 9.3 36.0 25.2 8.2 6.0

DMDS 7.4 6.2 6.9 5.0 6.2 4.3

H 22.5 18.8 20.9 18.5 23.2 16.8 18.4 13.4 26.3 18.4 21.9 16.12

C -C 3.4 3.0 7.3 5.3 3.2 2.41 3

C 0.9 0.8 1.7 1.2 0.4 0.34+

CO+CO 0.3 0.3 0.1 0.1 0.1 0.12

H S 3.3 2.9 2.4 1.7 1.5 1.12

H O 3.4 3.0 7.0 5.1 1.5 1.12

TOTAL 119.9 100.0 113.2 94.4 138.4 100.0 137.3 99.2 142.6 100.0 135.7 95.2

* Inclusive of ash and IOM

56

TABLE 22

MASS BALANCE DETAILS FOR CATALYST TESTS, 300 mL CSTR RUNS

Run Conditions : 440 C, 60 min @ 17.0 MPa H 2

Run No. 41oaLA-1,300CSTR 41oaLA-2,300CSTR 41bcLA-1,300CSTR

Feedstock 41oaLA>250 C 41oaLA>250 C 41bcLA>250 C

Catalyst 1% Mo Molyvan L 0.1% Mo Molyvan L 0.1% Mo Molyvan A

Wt In (g)

Feed 97.78 99.80 99.80

Catalyst 12.27 1.28 0.41

DMDS 6.15 7.42 8.48

H 26.32 29.16 26.672

Total In (g) 142.52 137.66 135.36

Wt Out (g)

Reactor Contents 91.84 89.30 87.32

Light Distillate 15.25 11.37 12.50

Water 1.50 2.29 2.30

C -C 's 3.15 4.83 3.961 3

C 0.40 0.37 0.484+

CO+CO 0.11 0.24 0.212

H S 1.54 1.55 3.282

H 21.90 25.51 23.332

Total Out 135.69 135.46 133.38

% Mass Balance 95.21 98.40 98.54

TABLE 23

HYDROTREATING RESULTS FOR CATALYST TESTS, 300 mL CSTR RUNS

Run Conditions : 440 °C, 60 min @ 17.0 MPa H 2

Run No. Feedstock Catalyst feed feed feed feed feedstock % THFI %

>566°C in >566°C C -C gas H uptake IOMfeedstock, resid conv., yield, by diff., QI IOM in conversionwt % MAF wt % MAF wt % MAF mg/g MAF QI in conversion feedstock THFI

1 3

a

2

b c

d

41oaLA-1, 41oaLA 1% Mo 32.8 72.6 1.1 46 7.9 50 17.4 73300CSTR >250 °C Molyvan L

41oaLA-2, 41oaLA 0.1% Mo 32.8 57.6 2.4 38 7.9 54 17.4 75300CSTR >250 °C Molyvan L

41bcLA-1, 41bcLA 0.1% Mo 31.3 54.9 1.1 35 6.1 54 16.3 76300CSTR >250 °C Molyvan A

a. Excludes C from DMDS1

b. Includes the hydrogen associated with the decomposition of DMDS.c. QI wt % MAF feed.d. THFI wt % MAF feed.

TABLE 24

OVERALL ELEMENTAL BALANCE FOR CATALYST TESTS, 300 mL CSTR RUNS

Run no. C, H, N, S, O, ash, wt % (SOfree) Total, wt %

3

wt % wt % wt % wt % wt % (diff)

41oaLA-1,300CSTR 102 97 81 49 46 85 95

41oaLA-2,300CSTR 103 98 90 55 89 67 98

41bcLA-1,300CSTR 100 97 109 100 102 65 99

58

TABLE 25

ELEMENTAL ANALYSES FOR CATALYST TESTS, 300 mL CSTR RUNS

1% Mo Molyvan L (Run No 41oaLA-1,300CSTR)

C, H, N, S, O, wt % Ash, H/Cwt % wt % wt % wt % (diff) wt % Total (atomic)

IN

41oaLA 84.6 5.2 1.1 0.7 5.8 2.6 100.0 0.74

Molyvan L 50.9 9.3 1.4 11.7 1.2 25.6 100.0 -

DMDS 27.1 6.3 1.8 58.1 6.8 0.0 100.0 -

OUT

Reactor Contents 83.8 8.0 1.0 0.9 1.3 5.0 100.0 1.15

Light Distillate 82.4 10.6 1.3 3.2 1.8 0.7 100.0 1.54

0.1% Mo Molyvan L (Run No 41oaLA-2,300CSTR)

C, H, N, S, O, wt % Ash, H/Cwt % wt % wt % wt % (diff) wt % Total (atomic)

IN

41oaLA 84.6 5.2 1.1 0.7 5.8 2.6 100.0 0.74

Molyvan L 50.9 9.3 1.4 11.7 1.2 25.6 100.0 -

DMDS 27.1 6.3 1.8 58.1 6.8 0.0 100.0 -

OUT

Reactor Contents 86.3 7.1 1.1 0.4 2.9 2.2 100.0 0.99

Light Distillate 73.5 8.8 1.8 8.9 7.0 0.0 100.0 1.44

0.1% Mo Molyvan A (Run No 41bcLA-1,300CSTR)

C, H, N, S, O,wt % Ash H/Cwt % wt % wt % wt % (diff) wt % Total (atomic)

IN

41bcLA 84.6 5.4 1.2 0.5 5.7 2.6 100.0 0.77

Molyvan A 29.5 5.1 2.5 20.0 0.9 42.0 100.0 -

DMDS 27.1 6.3 1.8 58.1 6.8 0.0 100.0 -

OUT

Reactor Contents 85.8 6.9 1.4 0.4 3.4 2.1 100.0 0.97

Light Distillate 64.0 8.1 1.8 16.7 9.4 0.0 100.0 1.52

TABLE 26

BOILING RANGE DISTRIBUTION FOR CATALYST TESTS, 300 mL CSTR RUNS wt %, AS DETERMINED

Boiling Point 1% Mo Molyvan L 0.1% Mo Molyvan L 0.1% Mo Molyvan ARange, Co

41oaLA-1,300CSTR, 41oaLA-2,300CSTR, 41bcLA-1,300CSTR,

Feed Products Feed Products Feed Products Catalyst

41oaLA Contents Distillate 41oaLA Contents Distillate 41bcLA Contents Distillate Molyvan LReactor Light Reactor Light Reactor Light

<150 0.0 1.4 41.2 0.0 1.3 30.1 0.0 0.8 41.4 0.0151-200 0.3 2.4 10.4 0.3 0.6 11.6 0.7 1.0 8.7 0.0201-250 1.9 2.5 14.5 1.9 1.8 21.7 2.8 1.5 23.6 0.3251-300 2.6 5.6 11.5 2.6 4.2 16.6 2.3 4.0 13.5 1.7301-350 15.0 22.6 12.6 15.0 20.7 16.0 15.1 19.9 9.5 17.4351-400 22.7 26.1 6.1 22.7 27.6 4.0 23.1 28.1 3.3 25.7401-450 13.7 13.1 2.3 13.7 14.8 0.0 13.1 14.7 0.0 10.1451-500 6.4 6.3 1.4 6.4 7.0 0.0 6.4 7.3 0.0 5.2501-550 2.3 3.7 0.0 2.3 3.6 0.0 2.2 3.8 0.0 0.0551-600 1.7 3.2 0.0 1.7 3.5 0.0 3.7 3.5 0.0 0.0601-650 1.3 2.5 0.0 1.3 2.7 0.0 0.0 2.8 0.0 0.0651-700 0.3 1.7 0.0 0.3 0.0 0.0 0.0 2.0 0.0 0.0701-750 0.0 0.0 0.0 0.0 0.0 0.0 0.0 1.3 0.0 0.0751-800 0.0 0.0 0.0 0.0 0.0 0.0 0.0 0.0 0.0 0.0

>800 31.8 8.9 0.0 31.8 12.2 0.0 30.6 9.3 0.0 39.6Total 100.0 100.0 100.0 100.0 100.0 100.0 100.0 100.0 100.0 100.0

>566°C 34.6 15.3 0.0 34.6 17.3 0.0 33.1 17.8 0.0 39.6

60

TABLE 27

FEED AND PRODUCT DISTRIBUTION FOR CATALYST TESTS, 300 mL CSTR RUNS

Run Conditions: 440 °C, 60 min @ 17.0 MPa H 2Run Number 41oaLA-1,300CSTR 41oaLA-2,300CSTR 41bcLA-1,300CSTR

1% Mo as Molyvan L 0.1% Mo as Molyvan L 0.1% Mo as Molyvan A

Feed In Product Out Feed In Product Out Feed In Product Out

(g) (%) (g) (%) (g) (%) (g) (%) (g) (%) (g) (%)

Feed 41oaLA 41oaLA 41bcLA

Feed Wt 97.8 68.6 99.8 72.5 99.8 73.7

Catalyst Molyvan L Molyvan L Molyvan A

Catalyst Wt 12.3 8.6 1.3 0.9 0.4 0.3

Boiling Pt. Range, °C

IBP-150 0.0 0.0 7.6 5.6 0.0 0.0 4.6 3.4 0.0 0.0 5.9 4.4

151-200 0.3 0.2 3.8 2.8 0.3 0.2 1.9 1.4 0.7 0.5 2.0 1.5

201-250 1.9 1.3 4.5 3.3 1.9 1.4 4.1 3.0 2.8 2.1 4.3 3.2

251-300 2.8 2.0 6.9 5.1 2.6 1.9 5.6 4.1 2.3 1.7 5.2 3.9

301-350 16.8 11.8 22.7 16.7 15.2 11.0 20.3 15.0 15.1 11.2 18.6 13.9

351-400 25.3 17.7 24.9 18.3 23.0 16.7 25.1 18.5 23.1 17.1 24.9 18.7

401-450 14.6 10.2 12.4 9.1 13.8 10.0 13.2 9.7 13.1 9.7 12.8 9.6

451-500 6.9 4.8 6.0 4.4 6.5 4.7 6.3 4.6 6.4 4.7 6.4 4.8

501-550 2.2 1.5 3.4 2.5 2.3 1.7 3.2 2.4 2.2 1.6 3.3 2.5

551-600 1.7 1.2 2.9 2.1 1.7 1.2 3.1 2.3 3.7 2.7 3.1 2.3

601-650 1.3 0.9 2.3 1.7 1.3 0.9 2.4 1.8 0.0 0.0 2.4 1.8

651-700 0.3 0.2 1.6 1.2 0.3 0.2 0.0 0.0 0.0 0.0 1.7 1.3

701-750 0.0 0.0 0.0 0.0 0.0 0.0 0.0 0.0 0.0 0.0 1.1 0.8

751-800 0.0 0.0 0.0 0.0 0.0 0.0 0.0 0.0 0.0 0.0 0.0 0.0

>800 * 36.0 25.2 8.2 6.0 32.2 23.4 10.9 8.0 30.5 22.5 8.1 6.1

DMDS 6.2 4.3 7.4 5.4 8.5 6.3

H 26.3 18.4 21.9 16.1 29.2 21.2 25.5 18.8 26.7 19.7 23.3 17.52

C -C 3.2 2.4 4.8 3.5 4.0 3.01 3

C 0.4 0.3 0.4 0.3 0.5 0.44+

CO+CO 0.1 0.1 0.3 0.2 0.2 0.12

H S 1.5 1.1 1.6 1.2 3.3 2.52

H O 1.5 1.1 2.3 1.7 2.3 1.72

TOTAL 142.6 100.0 135.7 95.2 137.7 100.0 135.5 98.4 135.4 100.0 133.4 98.5

* Inclusive of ash and IOM

61

TABLE 28

MASS BALANCE DATA FOR RESIDENCE TIME TESTS, 300 mL CSTR RUNS

Run Conditions : 440 °C @ 17.0 MPa H 2

Catalyst : 0.1% Mo Molyvan A

Run No. 41bcLA-2,300CSTR 41bcLA-1,300CSTR 41bcLA-3,300CSTR

Feedstock 41bcLA>250 C 41bcLA>250 C 41bcLA>250 C

Time (mins) 30 60 150

wt In (g)

Feed 115.3 99.80 97.19

Catalyst 0.48 0.41 0.40

DMDS 8.48 8.48 16.69

H 20.54 26.67 54.112

Total In (g) 144.8 135.36 168.39

Wt Out (g)

Reactor Contents 107.2 87.32 74.29

Light Distillate 6.38 12.50 18.40

Water 1.92 2.30 3.47

C -C 's 3.33 3.96 8.121 3

C 0.49 0.48 1.004+

CO+CO 0.18 0.21 0.192

H S 2.74 3.28 6.552

H 17.4 23.33 44.392

Total Out 139.64 133.38 156.41

% Mass Balance 96.44 98.54 92.89

TABLE 29

CONVERSION DATA FOR RESIDENCE TIME TESTS, 300 mL CSTR RUNS

Run Conditions: 440 C @ 17.0 MPa H 2

Catalyst: 0.1% Mo Molyvan A

Run No. Feedstock Time, min feed feed MAF feed feed feedstock % THFI %

>566 C in >566 C H uptake IOMfeedstock, resid conv., C -C gas by diff., QI IOM in conversionwt % MAF wt % MAF yield, wt % mg/g MAF QI in conversion feedstock THFI

1 3

a

2

b c

d

41bcLA-2, 41bcLA 30 31.3 42.9 0.4 28 6.1 50 16.3 73300CSTR >250 C

41bcLA-1, 41bcLA 60 31.3 54.9 1.1 35 6.1 54 16.3 76300CSTR >250 °C

41bcLA-3, 41bcLA 150 31.3 73.1 2.6 103 6.1 78 16.3 86300CSTR >250 °C

a. Excludes C from DMDS.1

b. Includes the hydrogen associated with the decomposition of DMDS.c. QI wt % MAF feed.d. THFI wt % MAF feed

TABLE 30

OVERALL ELEMENTAL BALANCE FOR RESIDENCE TIME TESTS, 300 mL CSTR RUNS

Run No. C, wt % H, wt % N, wt % S, wt % O, wt %, diff ash, wt % (SO free) Total, wt %3

41bcLA-2,300CSTR 99 96 96 80 86 72 96

41bcLA-1,300CSTR 100 97 109 100 102 65 99

41bcLA-3,300CSTR 97 89 86 82 81 100 93

63

TABLE 31

ELEMENTAL ANALYSES FOR RESIDENCE TIME TESTS, 300 mL CSTR RUNS

30 mins at 440 °C (Run No 41bcLA-2,300CSTR)

C, H, N, S, O, wt % Ash, H/Cwt % wt % wt % wt % (diff) wt % Total (atomic)

IN

41bcLA 84.6 5.4 1.2 0.5 5.7 2.6 100.0 0.77

Molyvan A 29.5 5.1 2.5 20.0 0.9 42.0 100.0 -

DMDS 27.1 6.3 1.8 58.1 6.8 0.0 100.0 -

OUT

Reactor Contents 85.9 6.6 1.3 0.6 3.4 2.2 100.0 0.92

Light Distillate 59.9 8.2 1.6 20.8 9.5 0.0 100.0 1.64

60 mins at 440 °C (Run No 41bcLA-1,300CSTR)

C, H, N, S, O, wt % Ash, H/Cwt % wt % wt % wt % (diff) wt % Total (atomic)

IN

41bcLA 84.6 5.4 1.2 0.5 5.7 2.6 100.0 0.77

Molyvan A 29.5 5.1 2.5 20.0 0.9 42.0 100.0 -

DMDS 27.1 6.3 1.8 58.1 6.8 0.0 100.0 -

OUT

Reactor Contents 85.8 6.9 1.4 0.4 3.4 2.1 100.0 0.97

Light Distillate 64.0 8.1 1.8 16.7 9.4 0.0 100.0 1.52

150 mins at 440 °C (Run No 41bcLA-3,300CSTR)

C, H, N, S, O,wt % Ash H/Cwt % wt % wt % wt % (diff) wt % Total (atomic)

IN

41bcLA 84.6 5.4 1.2 0.5 5.7 2.6 100.0 0.77

Molyvan A 29.5 5.1 2.5 20.0 0.9 42.0 100.0 -

DMDS 27.1 6.3 1.8 58.1 6.8 0.0 100.0 -

OUT

Reactor Contents 86.3 7.0 1.2 0.7 1.1 3.7 100.0 0.98

Light Distillate 72.5 8.9 1.7 9.4 7.5 0.0 100.0 1.48

TABLE 32

BOILING RANGE DISTRIBUTION FOR RESIDENCE TIME TESTS, 300 mL CSTR RUNS,wt % AS DETERMINED

Boiling Point 30 min @440 °C 60 min @440 °C 150 min @440 °CRange, Co

41bcLA-2,300CSTR, 41bcLA-1,300CSTR, 41bcLA-3,300CSTR,

Feed Products Feed Products Feed Products

41bcLA Reactor Light 41bcLA Reactor Light 41bcLA Reactor LightContents Distillate Contents Distillate Contents Distillate

<150 0.0 0.0 50.6 0.0 0.8 41.4 0.0 0.0 29.1

151-200 0.7 1.9 11.7 0.7 1.0 8.7 0.7 0.8 14.1

201-250 2.8 1.9 24.8 2.8 1.5 23.6 2.8 1.0 22.0

251-300 2.3 3.9 7.6 2.3 4.0 13.5 2.3 3.1 17.5

301-350 15.1 18.6 4.0 15.1 19.9 9.5 15.1 24.0 14.8

351-400 23.1 27.0 1.3 23.1 28.1 3.3 23.1 30.3 1.9

401-450 13.1 13.8 0.0 13.1 14.7 0.0 13.1 14.2 0.6

451-500 6.4 7.2 0.0 6.4 7.3 0.0 6.4 6.9 0.0

501-550 2.2 3.8 0.0 2.2 3.8 0.0 2.2 4.1 0.0

551-600 3.7 3.2 0.0 3.7 3.5 0.0 3.7 3.7 0.0

601-650 0.0 3.2 0.0 0.0 2.8 0.0 0.0 2.7 0.0

651-700 0.0 2.5 0.0 0.0 2.0 0.0 0.0 1.8 0.0

701-750 0.0 1.8 0.0 0.0 1.3 0.0 0.0 0.0 0.0

751-800 0.0 0.3 0.0 0.0 0.0 0.0 0.0 0.0 0.0

>800 30.6 10.9 0.0 30.6 9.3 0.0 30.6 7.4 0.0

Total 100.0 100.0 100.0 100.0 100.0 100.0 100.0 100.0 100.0

>566°C 33.1 20.9 0.0 33.1 17.8 0.0 33.1 14.4 0.0

65

TABLE 33

FEED AND PRODUCT DISTRIBUTION FOR RESIDENCE TIME TESTS, 300 ML CSTR RUNS

Run Conditions: 440 °C, 17.0 MPa H and 0.1% Mo as Molyvan A2

Run Number 41bcLA-2,300CSTR 41bcLA-1,300CSTR 41bcLA-3,300CSTR30 mins 60 mins 150 mins

Feed In Product Out Feed In Product Out Feed In Product Out

(g) (%) (g) (%) (g) (%) (g) (%) (g) (%) (g) (%)

Feed 41bcLA 41bcLA 41bcLA

Feed Wt 115.3 79.6 99.8 73.7 97.2 57.7

Catalyst Molyvan A Molyvan A Molyvan A

Catalyst Wt 0.5 0.3 0.4 0.3 0.4 0.2

Boiling Pt. Range,°C

IBP-150 0.0 0.0 3.2 2.3 0.0 0.0 5.9 4.4 0.0 0.0 5.4 3.5

151-200 0.8 0.6 2.8 2.0 0.7 0.5 2.0 1.5 0.7 0.4 3.2 2.0

201-250 3.2 2.2 3.6 2.6 2.8 2.1 4.3 3.2 2.7 1.6 4.8 3.1

251-300 2.7 1.9 4.7 3.4 2.3 1.7 5.2 3.9 2.2 1.3 5.5 3.5

301-350 17.4 12.0 20.2 14.5 15.1 11.2 18.6 13.9 14.7 8.7 20.6 13.2

351-400 26.6 18.4 29.0 20.8 23.1 17.1 24.9 18.7 22.5 13.4 22.9 14.6

401-450 15.1 10.4 14.8 10.6 13.1 9.7 12.8 9.6 12.7 7.5 10.7 6.8

451-500 7.4 5.1 7.7 5.5 6.4 4.7 6.4 4.8 6.2 3.7 5.1 3.3

501-550 2.5 1.7 4.1 2.9 2.2 1.6 3.3 2.5 2.1 1.2 3.0 1.9

551-600 4.3 3.0 3.4 2.4 3.7 2.7 3.1 2.3 3.6 2.1 2.7 1.7

601-650 0.0 0.0 3.4 2.4 0.0 0.0 2.4 1.8 0.0 0.0 2.0 1.3

651-700 0.0 0.0 2.7 1.9 0.0 0.0 1.7 1.3 0.0 0.0 1.3 0.8

701-750 0.0 0.0 1.9 1.4 0.0 0.0 1.1 0.8 0.0 0.0 0.0 0.0

751-800 0.0 0.0 0.3 0.2 0.0 0.0 0.0 0.0 0.0 0.0 0.0 0.0

>800 (*) 35.3 24.4 11.7 8.4 30.5 22.5 8.1 6.1 29.7 17.6 5.5 3.5

DMDS 8.5 5.9 8.5 6.3 16.7 9.9

H 20.5 14.2 17.4 12.5 26.7 19.7 23.3 17.5 54.1 32.1 44.4 28.42

C1-C3 3.4 2.4 4.0 3.0 8.2 5.2

C4+ 0.5 0.4 0.5 0.4 1.0 0.6

CO+CO 0.2 0.1 0.2 0.1 0.2 0.12

H S 2.7 1.9 3.3 2.5 6.6 4.22

H O 1.9 1.4 2.3 1.7 3.5 2.22

TOTAL 144.8 100.0 139.6 96.4 135.4 100.0 133.4 98.5 168.4 100.0 156.4 92.9

* Inclusive of ash and IOM

66

Figure 1. Simplified Block Flow Diagram, Liquefaction System, Original Novel Concept Case.

67

Figure 2. Simplified Block Flow Diagram, Conceptual Commercial PlantOriginal Novel Concept Case.

68

Figure 3. Simplified Block Flow Diagram, Liquefaction System, Improved Novel Concept Case.

69

Figure 4. Effect of HI "E" on Coal Conversion

70

Figure 5. Effect of Boiling Point Distribution on Resid Conversion.

71

Figure 6. Effect of Catalyst Concentration on Conversion.

263J 28oab LAFDR 41oa LA20

30

40

50

60

70

80

20

30

40

50

60

70

80

Feedstock

Catalyst : Molyvan L (1% Mo)Reactor : CSTRReaction conditions : 60min @ 440°C

resid conversion

IOM conversion

H2 uptake

263J 28oab LAFDR 41oa LA0

20

40

60

80

100

Feedstock

resid content

IOM content

72

Figure 7. Feedstock Conversion and Hydrogen Uptake.

73

Figure 8. The Influence of Catalyst Type and Concentrationon the Conversion of Feedstock 41LA.

20 40 60 80 100 120 140 16040

50

60

70

80

90

Residence time, min

Feedstock : 41bc LA (>250°C)Catalyst : Molyvan A (0.1%Mo)Reactor : CSTRReaction temperature : 440°C

20 40 60 80 100 120 140 1600

20

40

60

80

100

120

0

2

4

6

8

10

12

Residence time, min

resid conversion

IOM conversion

QI conversion

H2 uptake

C1-C3 gas yield

74

Figure 9. The Influence of Residence Time on the Conversion of Feestock 41LA.

75

FigFigure 10. Feedstock Hydrocracking: Effect on Boiling Point Distribution

76

FigurFigure 11. Catalyst Testing: Effect on Boiling Point Distribution.

77

FigFigure 12. Residence Time Testing: Effect on Boiling Point Distribution.

A-78

APPENDIX 1

HYDRIDE ION SOURCE PRODUCTIONAND HYDRIDE ION COAL LIQUEFACTION STUDIES

ANNOTATED BIBLIOGRAPHY(Completed September 1996)

M. S. Lancet; S. D. BrandesCONSOL Inc.

Research & Development4000 Brownsville Road

Library, PA 15129

M. PelusoLDP Associates

32 Albert E. Bonacci DriveHamilton Square, NJ 08690

A-79

HYDRIDE ION SOURCE PRODUCTION

Aguilo, A.; Horlenko, T. "Formic Acid", Hydrocarbon Proc. 1980, 120-130.

This is an excellent review of the production and industrial uses of formic acid in 1980. All 7 5million lb of formic acid production in the U.S. at that time came as a by-product of the liqui dphase oxidation of n-butane to acetic acid. Main U.S. producers were Union Carbide an dCelanese. Formic acid is prod uced in Europe via formamide since formamide is widely used forthe production of HCN in Europe. The authors predict that future formic acid capacity is likely tobe provided by hydrolysis of methyl formate made from methanol carbonylation (Leonar dProcess) or dehydrogenation. Other formic acid synthesis routes discussed are: production fromformate salts, direct synthesis from CO and H O, dehydrogenation of MeOH to methyl formate2

and direct synthesis of formamide. Seventy references are cited in this article.

Ammon, C., U.S. Patent No. 2,070,503, February 9, 1937. "Method for Converting Carbo nMonoxide into Formic Acid".

This is the first U.S. Patent that deals with the production of formic acid by the hydrolysis o fmethyl formate with recyc le of co-produced methanol to the methyl formate production step. Anapparatus to accomplish the hydrolysis also is described. This technique has never bee ncommercialized.

Badische Anilin, Soda Fabrik, British Patent No. 252,848, June 10, 1926, "Improvements in theManufacture and Production of Alkyl Esters of Formic Acid".

This patent describes a method to synthesize methyl formate and ethyl formate by the reactionof CO with methanol and ethanol at 30 to 140 C and 200 to 1000 atm in the presence of sodiumalcoholates (sodium methoxide and sodium ethoxide). They mention the need for sulfur-free andmoisture-free CO in this process.

Braca, G.; Raspolli G., A. M.; Laniyonu, N. J.; Sbrana, G,; Micheli, E.; DiGirolamo, M. ;Marchionna, M. "Hydrogenolysis of Methyl Formate by H /CO Mixtures with CuO/ZnO/Al O2 2 3

Based Methanol Synthesis Catalysts". Ind. Eng. Chem. 1995 Rev. 34, 2358-2363.

This paper addresses the hydrogenolysis (second) step in the following two-step methano lsynthesis (CH OH + CO HCOOCH ) followed by (HCOOCH + 2H 2CH OH). The authors3 3 3 2 3

claim to have solved the problem of catalyst poisoning by CO during the hydrogenolysis. Thisallows the use of impure H (i.e., hydrogen gas tha t contains residual unreacted CO from the first2

step) in this reaction. They report that CuO/ZnO/Al O based methanol synthesis catalysts are2 3

active for hydrogenolysis under mild conditions (150-185 C, 5-10 MPa) both in gas and liquidphases with high selectivity to MeOH (80-90%) in a hydrogen feed containing CO with no catalystpoisoning or decrease in selectivity to MeOH.

Burgemeister, T.; Kastner, F.; Leitner, W. "(PP) RhH] and [(PP) Rh][O CH] Complexes as Models2 2 2

for the Catalytically Active Intermediates in the Rh-Catalyzed Hydrogenation of CO to HCOOH"”2

Angew Chem. Int. Ed. Engl. 1993, 32, 739-741.

A-80

The rhodium catalyzed transf er-hydrogenation of formates also is mentioned and referenced. Apaper by Tsai (see below) also is referenced. This technique does not appear to be practical.

Carpenter, G., U.S. Patent No. 1,949,825, March 6, 1934. "Process for the Production o fCarboxylic Acids".

This is a second DuPont U.S. Patent (see Vail, No. 1,895,238 below) for catalysts to produc eformic acid from CO and steam. Two non-vol atile acidic elements from Groups III through VI areused (e.g. phosphomolybdic acid on activated carbon) at pressures from 25 to 900 atmospheresand temperatures from 200 to 300 C. This information does not appear to be relevant for theNovel Concept project, but the use of phosphomolybdic acid is interesting since Exxon ha spatents for heavy oil and coal conversion that use this same material.

Carpenter, G. B., U.S. Patent No. 1,991,732, Feb. 19, 1935. "Process for the Production o fFormic Acid".

Three U.S. Patents (Nos. 1,991,732; 2,001,659; 2,023,003), issued to DuPont in 1935, ar eclosely related and appear to differ only in the catalyst choice. This scheme is unlikely to b euseful in the Novel Concept Liquefaction Process.

This patent describes a catalytic process for the formation of formic acid from CO and steam athigh pressure (25 to 900 atm) at 100 to 400 C. One or more metal halides are used as th ecatalyst, but zinc chloride seems to be the preferred catalyst as it alone is mentioned in th eclaims. The synthesis requires relatively high-purity CO; however, it is suggested that inert gasessuch as nitrogen or carbon dioxide can be used to control the temperature of the exothermi creaction at the expense of formic acid yield. (See also related patents by Carpenter an dWoodhouse.)

Carpenter, G. B., U.S. Patent No. 2,023,003, Dec. 3, 1935. "P rocess for the Production of FormicAcid".

The process described in this patent is similar to two others that are described (Carpenter andWoodhouse). The main difference is that a non- volatile acid salt of acidic oxides such as As, W,Mo, U, Cr, V, B, Si and Zr is used as the catalyst. Other conditions are approximately the same(i.e., high-purity CO, up to 900 atm., 100 to 400 C).

Chang, B-H.; Grimm, R. A.; Trivedi, B. C., U.S. Patent 4 661 623, 1987. "Method of ProducingMethyl Formate from Methanol and Carbon Monox ide Using Anionic Group VIII Metal Catalysts",Ashland Oil, Inc.

This patent describes the production of methyl fo rmate from methanol and CO at <3000 psia and<200 C. This process requires the use of anhydrous methanol an d CO. The catalyst is the cruxof the invention. This catalyst is an alkali metal salt of various Group VIII metal anionic ligands.Selectivity to methyl formate is very high (>99%) and the authors claim economic benefits overconventional catalysts (alkaline or alkali metal methoxides). It appears to be a good method for

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production of high purity methyl formate, but hydride ion reagent purity may not be critical to itsapplicability to the novel concept liquefaction process.

Coteau, W.; Ramioulle, J.; UCB, British Patent No. 1,511,961, May 24, 1978, "Process for th eProduction of Methyl Formate".

This patent describes improvements to the methanol carbonylation processes for producin gmethyl formate in the presence of metal methoxides. The process reacts CO with a recycle dstream of liquid (MeOH + MF). They claim that this results in lower operating pressure (20 to 110atm) at 70 to 110 C without a need to recycle CO. The CO used in this process does not needto be high purity CO and it can contain hydrogen, nitrogen or other process-inert gases. Th eprocess still requires both the CO and MeOH to be essentially moisture-free to prevent catalystprecipitation.

Czaikowski, M. P.; Bayne, A. R. "Make Formic Acid from CO", Hydrocarbon Processing 1980 59,103-106.

This is a very good review of commercial formic acid production with particular focus on th eScientific Design/Bethlehem formic acid process. Some economics and capacity data ar epresented, however, the data are at least 15 years old. The authors are Scientific Desig nemployees.

Darensbourg, D.; Gray, R.; Ovalles, C.; Pala, M. "Homogeneous Catalysis of Methyl FormateProduction from Carbon Monoxide and Methanol in the Presence of Metal Carbonyl Catalysts",Journal of Molecular Catalysis 1985, 29, 285-290.

This paper presents experimental results for methyl formate production using tungsten an druthenium carbonyl catalysts. The results show that the methyl formate production per unit o fcatalyst is 3 to 6 times greater for these catalysts compared to the conventional potassiu mmethylate catalyst. The cost of conventional make-up catalyst per pound of methyl formate i sprobably in the $.50 to $.75 range. This factor, togethe r with the expected large price differencesbetween the metal carbonyl catalysts a nd the potassium catalyst, probably explains why metalliccarbonyl catalysts are not used currently in commercial methyl formate production.

Dreyfus, H., U.S. Patent No. 2,028,764, January 26, 1936. "Synthesis of Formic Acid".

This patent describes the production of formic acid from CO and water via the presence of a nacid such as a lower fatty type, acet ic or hydrochloric. The acid can be in a solid, liquid or vaporform. Also a salt such as cuprous chloride can be used to adsorb the CO. Reaction conditionsare temperatures of 150 to 250 C at pressures up to 20 atmospheres.

Drury, D., European Patent Application No. 0-095-321, DOP November 30, 1983. "Productio nof Formate Salts".

The patent deals with the production of formate salts of nitrogenous bases containing tertiar ynitrogen atoms (e.g., trialkylamm onium formate). The nitrogenous base is reacted with CO and2

hydrogen using a soluble metal catalyst such as ruthenium trichloride. This concept does not

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appear to be useful to us but reference is made to a Japanese Patent A pplication (No. 53-46820)in which a somewhat similar system is used to produce esters of formic acid.

Edwards, J.; Nicolaidis, J. ; Cut llip, M.; Bennett, C. "Methanol Partial Oxidation at Low Tempera-ture", Journal of Catalysis 1977, 50, 24-34.

This is a study of formaldehyde formation wher e, at low temperature, methyl formate forms as anintermediate compound. This is not of much interest unless conditions could be determine dunder which methyl formate would be favored. An interesting sidelight is the use of ferri cmolybdate and molybdenum trioxide in formaldehyde synthesis.

Eversole, J. F., U.S. Patent No. 2,160,064, May 30, 1939. "Manufacture of Formic Acid".

A process for the production of relatively pure formic acid from methanol is described. In thi sprocess, vapor phase methanol is dehydrogenated to methyl formate and hydrogen at 150 t o400 C over a reduced copper oxide catalyst. The gases are removed and the methyl formateis distilled from the unreacted methanol and fed into a boiling aqueous solution of formic acid thatcontains a small amount of hydrolysis catalyst. The methyl formate is hydrolyzed to formic acidand methanol. Pure fo rmic acid is recovered as product by a dual azeotrope distillation and themethanol is recycled to the dehydrogenation unit. The methyl formate synthesis part of th escheme appears to be the same as methanol dehydrogenation processes dating back at leastto 1921.

Forster, D. "The Commercial Utility of Syn thesis Gas and Methanol as Building Blocks - PresentSituation and Future Prospects", Am. Chem. Soc. Pet. Div. Preprints 1986, 31, 69-73.

This is a rather cursory review of the state of C chemistry in 1986. It briefly covers synthesis gas1

production and its uses in the manufacture of chemicals. The author mentions formic aci dproduction via methanol carbonylation as a significant new use for methanol.

Furusaki, S. Manada, N., Yamashina, H., Matsuda, M., U.S. Patent No. 4,420,633 December 13,1983. "Process for the Preparation of an Ester of Formic Acid".

This patent describes preparation of esters of formic acid (e.g., methyl formate) by the vapo rphase reaction of CO, hydrogen, and a nitrous acid (e.g., methyl nitrite) over a platinum o rpalladium supported catalyst. A simple separation of the product is claimed as one of th eadvantages of the process. In the proposed methyl formate system, by-product methanol fromthe first stage liquefaction reactor will be recovered and recycled to the methyl format eregeneration system.

Gassner, F.; Leitner, W. "Hyd rogenation of Carbon Dioxide to Formic Acid Using Water-SolubleRhodium Catalysts", J. Am. Chem. Soc., Chem. Commun. 1993, 1465-1466.

This paper describes the direct hydroge nation of CO to formic acid in aqueous solution of water2

soluble rhodium catalysts in the presence of amines. Considerably higher turnover rates (up to

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~290 mol HCOOH per mol cata lyst per hour) were observed than were found in various organicsolvents (THF, DMSO, benzene). However, the rate of formic acid format ion was only about 20%of that found in supercritical CO 2.

Higdon, B. W.; Hobbs, C. C.; Onore, M. J., U.S. Patent No. 3,812,210, May 21, 1974. "Vapo rPhase Decomposition of Methyl Formate to Yield Methanol".

This patent describes a catalytic thermal decomposition process to synthesize methanol fro mmethyl formate. In this process, methyl formate is contacted with an alkaline earth metal oxide(preferably barium oxide) at 200 to 500 C to form methanol and CO. This patent reinforces thefact that methyl formate is thermodynamically unstable above about 200 C.

Hiratani, T.; Noziri, S. "C Chemistry Based on Methyl Formate", Chem. Econ & Eng. Rev. 19851

17, 21-24.

This is an interesting concept to make use of the large quantities of CO that are produced in basicoxygen converters in steel plants. The CO off gas is rea cted with import methanol to form methylformate: CH OH + CO HCOOCH . The synthesized methyl formate is an easily transported3 3

chemical feedstock that can be used as a bui lding block for many products. This process, whichemphasizes C chemistry that is not dependent on naphtha, is of particular interest i s1

resource-poor countries like Japan, its country of origin.

Ikarashi, T. "New MGC Process for High-Purity Carbon Monoxide Production and Review o nDerivatives of Methyl Formate", Chem. Econ & Eng. Rev. 1980 12, 31-34.

This paper describes the Mitsubishi Gas Chemical (MGC) process to produce high-purity C Ofrom methanol. It is proposed as an alternative to cryogenic separation of the gas produced bysteam reformation of natural gas or naphtha or by the partial oxidation of heavy oil. In thi sprocess, methanol is hydrogenated to methyl formate over a Cu-Zr-Zn or Cu-Zr-Zn-Al catalyst.The methyl formate is subsequently pyrolized to CO and methanol over an alkali earth metal ,activated carbon or zeolite catalyst. The product yields from the MGC process, particularly o fmethyl formate, are several hundred times as high as earlier dehydrogenation processes.

Jessop, P. G.; Ikariya, T.; Noyory, R. "Homogeneous Catalytic Hydrogenation of Supercritica lCarbon Dioxide" Nature 1994, 368, 231-233.

The authors discuss the production of formi c acid by the hydrogenation of supercritical CO over2

a ruthenium catalyst in the presence of basic agents such as tertiary amines. They report thatcertain Ru(II)-phosphine-complex catalysts are highly active for the hydrogenation of supercriticalCO at 50 C and 205 atm in the presence of triethylamine (reaction rates are >18 times higher2

than the same reaction over the same catalyst at the same conditions in THF). Turnover ratesas high as 1400 mol HCOOH produced per mol of catalyst per hour are reported. The authorsindicate that further refinement is needed to commercialize this process.

Kaplan, L. "Formic Acid from CO Containing Gases", Chemical Engineering 1982, 71-73.

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The information in this article was used to develop preliminary costs for methyl format eproduction. It reviews the pre-1980 difficulties of hydrolyzing methyl formate to formic acid anddiscusses the two then-new process es (Leonard & Scientific Design/Bethlehem) for overcomingthis difficulty. Both processes start with CO and methanol to produce methyl formate an dintegrate this step with the downstream hydrolysis step where methanol is p roduced and recycled.The methyl formate producti on step is said to be "well-established", although improved versionsof the sodium methylate catalyst are said to be used. The capital investment and operatin grequirements for both processes are given. It is interesting to note that the inventor of one of theprocesses, Mr. Leonard, predicted that his process would eventually produce formic acid for 5to $.06/lb versus the then current price of $.29/lb.

Kim, K. M.; Woo, H. C.; Cheong, M.; Kim, J. C.; Lee, K. H.; Lee, J. S.; Kim, Y. G. "Chemica lEquilibria and Catalytic Reaction of Gas-Phase M ethanol Synthesis from Methyl Formate", Appl.Catal. 1992, 83, 15-30.

This article discusses gas phase methanol synthesis from methyl formate at atmospheri cpressure. There is a great deal of good thermodynamics data in this article, however, at lowerpressure than in the novel concepts process. The principal finding is a nearly step change i nmethyl formate stability at ~500 K (2 27 C), above which methyl formate decomposes rapidly (at1 atm.) to methanol (up to ~300 C) and to CO and H above ~300 C.2

Kirk, R. E.; Othmer, D. F., Encyclopedia of Chemical Technology

• Toubes, B., Encyclopedia of Chemical Technology, Kirk, R. E., Othmer, D. F., Editors ;(1954) Vol 6, 875-883. "Formic Acid".

• Jordan, T. E., Stern, D., Encyclopedia of Chemical Tech nology, Kirk, R. E., Othmer, D. F.,Editors; (1954) Vol 5, pp 824-850. "Esters, Organic".

• Lee, D. D., Encyclopedia of Chemical Technology, Kirk, R. E., Othmer, D. F., Editors ;(1954) Vol 3, pp 179-191. "Carbon Monoxide".

These are short but informative synopses of the synthesis, production and uses for formic acid,carbon monoxide and organic esters. The 1954 publication date of th is edition of the Kirk-OthmerEncyclopedia limits its usefulness.

Lee, J. S.; Kim, J. C.; Kim, Y. G. “Methyl Formate as a New B uilding Block in C Chemistry”, Appl.1

Catal. 1990, 57, 1-30.

This is a major review of methyl format e chemistry, production and uses. Most of the productionmethods that have been identified in the literature search are discussed in this review article .Principal synthesis methods are dehydrogenation of methanol (2CH OH HCOOCH + 2H ) with3 3 2

copper, silver or tungsten carbide catalysts and carbonylation of methanol (CH OH + CO 3

HCOOCH ) in the presence of an alkaline metal methoxide catalyst. Much of the paper i s3

devoted to discussing the conversion of methyl formate to other chemicals.

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The Leonard Process Co., Inc., Hydrocarbon Proc. (Nov, 1979), 176. “Formic Acid”

This is a one page description of a continuous process (Leonard Process) to produce high purityformic acid from impure CO and H O. CO is reacted with MeOH in the presence of a sodiu m2

methoxide catalyst to form methyl formate. The methyl formate is then hydrolyzed to formic acidand MeOH in a second stage. The formed MeOH is recycled to th e first stage and the net MeOHconsumption is <4 lb/100 lb of formic acid product.

Liu, Z.; Tierney, J.; Shah, Y.; Wender, I. “Kinetics of Two-Step Methanol Synthesis in the SlurryPhase”, Fuel Processing Technology 1988, 18, 185-199.

This paper presents the results of one liter batch autoclave runs for methyl formate productionvia conventional technology. Slightly hi gher than normal reaction temperatures are explored (upto 110 C) in order to investigate the feasibility of combining methyl formate and methan eproduction in a single reactor. The results show the superiority of potassium versus sodiu mmethylate catalyst, the significant effect of cataly st loading and quantifies the increase in reactionrate with increasing temperature. However, current commercial methyl formate productio noperates at near equilibrium conditions where methyl formate concentration increases wit hdecreasing reaction temperature. See also Palekar Ph.D. thesis (199 3) for additional information.

Lower, J., U.S. Patent No. 2,373,583, April 10, 1945. “Conversion of Methyl Formate to FormicAcid”

This is a DuPont patent that seeks to overcome the difficulties of hydrolyzing methyl formate toformic acid by using acetolyses with an organic dicarboxylic acid (e.g., adipic acid) in th epresence of an inorganic acid catalyst to form anhydrous formic acid.

Lynn, J. B., Homberg, O. A., Singleton, A. H., U.S. Patent No. 3,907,884, 1975. “Formic Aci dSynthesis by Lower Alkyl Formate Hydrolysis”

This patent describes the autocatalytic hydrolysis of formate esters of C to C alcohols to formic1 4

acid and alcohol. This invention eliminates the need for a ddition of an inorganic acid catalyst andresults in up to a 50 fold reduction in times required to reach equilibrium in the hydrolysis o fmethyl formate to formic acid. Autogenic pressur es and temperatures of 20-150 C are listed forthis reaction. This autocatalytic hydrolysis of methyl formate to formic acid is an integral part ofthe SD/Bethlehem formic acid process (see Czaikowski and Bayne).

Meyer, K., Muller, J., U.S. Patent No. 1,522,257, January 6, 1925. "Production of Formic Acid".

This BASF patent discloses the use of ammonium formate, sul furic acid and an added formamideto produce formic acid.

Palekar, V. M., PhD Dissertation, Univ. of Pittsburgh (1993). "Liquid Phase Synthesis o fMethanol/Methyl Formate Under Mild Conditions".

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This PhD thesis is a very complete review and description of liquid phase methanol synthesis.The process that is the focus of th e work is the same multi-step single stage methanol synthesisvia methyl formate intermediate that is discussed in Tierney and Wender. Although DOE is notspecifically acknowledged in this thesis, it is likely that the work is part of that referred to i nTierney and Wender. The bibliography lists 102 references.

Reutemann, W.; Kiezcka, H. "Formic Acid Production", Ullmann's Ency. of Ind. Chem., 5thEd.,1989, A 12, 16-22.

This is a more recent encyclopedic description of the commercial production and uses of formicacid than can be found in the 1954 Kirk, Othmer work. For instance, the Leonard Process, aUSSR (Ukrainian) process and the SD/Bethlehem process, none of which existed in 1954 (thepublication date of our edition of the Kirk-Othmer Chemica l Encyclopedia), all are described. Thereference on hand is an excerpt and does not contain the bibliography listing.

Ritter, F. "High-Concentration Formic Acid", Industrial and Engr. Chemistry 1935, 1224-1225.

This article deals with a method for removing water from formic acid.

Rohde, K., U.S. Patent No. 1,848,664, March 8, 1932. "Production of Formic Acid".

This patent describes a method for preparing anhydrous formic acid from dilute formic acid b yheating the dilute formic acid in the presence of an acid anhydride. Particular mention is madeof phthalic acid anhydride and pyroboric acid anhydride. The anhydride is mixed with the diluteformic acid in a ratio that is high enough to take up the excess water from the formic acid an dheated to ~100 C for about 1 hour. Anhydrous formic acid is recovered with >99% yield.

Röper, M. "Oxygenated Base Chemicals from Synthesis Gas (Methyl Formate as Versatil eIntermediate)", Erdöl und Kohle - Erdgas - Petrochemie vereinigt mit Brennstoff-Chemin 1984, 37,506-511.

This 1984 paper is a very good, concise review of industrial production of and uses for methylformate. The author predicts that if and when syngas and methanol become competitive a sfeedstocks for the chemical industry, methyl formate based chemistry will become increasinglyimportant. An interesting use for methyl formate discussed in this paper, and mentioned i nseveral others as well, is as a storage and transport medium for syngas.

Schuchardt, U.; Sousa, F. M. B. "Oxalate as an Intermediate in the Base-Catalysed Water-GasShift Reaction", Fuel 1986, 65, 669-672.

The authors use an oxalate intermediate to explain observed base catalyzed water-ga sconversion. They suggest that carbonate, formed by reaction of the hydroxide with CO , reacts2

readily with CO to form oxalate. This oxalate intermediate then easily decomposes to CO and2

formate. Since the formate is stable to ~250 C and, therefore, does not easily decompose to

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carbonate and H below 300 C, the water-gas shift reaction becomes truly catalytic only above2

~300 C.

Tagaev, O.; Zhavoronkov, N.; Moiseev, I. ; Pazdersky, J.; Kalechits, I.; Kouchubei, V. "Processfor the Production of Methyl Formate", Belgium Patent Application No. 893.511, December 14,1982.

This patent (in French) discusses th e use of olefin hydroxide ogligomers (e.g., diethylene glycol)in addition to the standard alkali metal alcoholate catalyst. A simpler process with highe rproductivity is claimed.

Tierney et al.,"Synthesis of Methanol via Methyl Formate Appears Commercially Promising" ,Chemical Engineering 1989, 21.

This three paragraph blurb in the "CHEMINATOR" sectio n of Chemical Engineering mentions thebenefits of the Tierney et al. (Univ. of Pittsburgh) two-step (concurrent in a single slurry reactor)methanol synthesis with a methyl-formate intermediate. A homogeneous catalyst is used in theformation of methyl formate from methanol and CO and a heterogeneous catalyst is used in thesubsequent hydrogenation of methyl form ate to methanol. The process is reported to operate atabout 100 C lower than other processes with little recycle. This process is explained in detailelsewhere (Tierney and Wender).

Tierney, J. W.; Wender, I, "A Novel Process for Methanol Synthesis - Final Report" ,DOE/PC/89786-T18 (DE94017111) (1994).

This is the final report to DOE on work done at the University of P ittsburgh to develop a multi-stepsingle stage methanol synthesis v ia methyl formate intermediate. The synthesis operates undermild conditions (100-180 C, 30-65 atm) with up to 90% per pass conversions and 98% selectivityto methanol. It uses typically a mixed catalyst system com prised of alkali compounds and copperchromite or alkali promoted copper chromite to catalyze both the carbonylation of methanol t omethyl formate and the hydrogenation of methyl formate to methanol. The net reaction is :2H + CO CH OH. The process is tolerant o f up to 1% CO and 3000 ppm H O in the system.2 3 2 2

This is an excellent review of the state of methyl formate and methanol synthesis in 1994. Fifty-five references are cited.

Tsai, J-C.; Nicholas, K. "Rhodium-Catalyzed Hydrogenation of Carbon Dioxide to Formic Acid",Journal of the American Chemical Society 1992, 114, 5117-5124.

This is a voluminous paper dealing with the production of formic acid from CO and hydrogen2

using what seems to be a rather exotic catalyst. (See also the Burgemeister article mentionedabove.)

Vail, W., U.S. Patent No. 1,895,238, January 24, 1933. "Process for the Preparation of FormicAcid".

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This is the first of the DuPont patents dealing with the production of formic acid. The acid i sproduced from CO and steam using a boron phosphate catal yst in the vapor phase at a pressureof 100 to 1,000 atmospheres and a temperature of 200 to 300 °C.

Wainwright, M. S. "Catalytic Process for Methanol Synthesis - Established and Future", MethaneConversion 1988, 95-108. Bibby, D.M.; Chang, C.D.; Howe, R.F.; Yurchak, S., Eds., Elsevier,Amsterdam.

This book chapter is a fairly comprehensive review of methanol syntheses. Of interest to th eNovel Liquefaction Concept is the discussion of methanol production through a methyl formateintermediate. The chemistry involved is the catalytic carbonylation of an alcohol to its format eester followed by catalytic hydrogenation of this ester to methanol and the original alcohol. Theauthor concludes that the alcohol of choice always will be methanol. He mentions the possibilityof performing both the carbonylation and hydrogenatio n in a single reactor, but he states that thiswill not work due to selective poisoning o f the hydrogenation catalyst by CO. Apparently Tierneyand Wender, at the University of Pittsburgh, have overcome this problem, since they claim t ohave successfully demonstrated this concept.

Wakamatsu, H.; Shimomura, K., U.S. Patent No. 3,816,513, June 11, 1974. "Process fo rProducing Methyl Formate".

This Ajinomoto Co. patent discloses an integrated process for producing methyl formate fro msynthesis gas in four steps including the internal production and recycle of methanol. The quotedimpurity levels of carbon dioxid e and water for methyl formate production are significantly higherthan those stated elsewhere, presumably due to the use of a potassium methylate catalys tinstead of the traditional sodium methylate catalyst.

Wietzel, R., U.S. Patent No. 1,843,434, February 2, 1932. "Concentration of Aqueous Formi cAcid".

This IG Farben patent discloses the use of formamide and a mineral acid to remove water.

Willkie, H. F., U.S. Patent No. 1,400,195, December 13, 1921. "Process of Making Methy lFormate".

This is the earliest patent located that describes the production of methyl formate b ydehydrogenation of methanol over a metallic (copper, nickel, chromium or iron) catalyst. Th eprocess appears to operate at atmospheric pressure, but no pressure is sp ecified. A temperaturerange of 350 to 450 F is specified. Current technology tends to favor production of methy lformate via methanol carbonylation rather than dehydrogenation.

Woodhouse, J. C., U.S. Patent No. 2,001,659, May 14, 1935. "Preparation of Formic Acid".

A catalytic process for the formation of formic acid from CO and steam at high pressure (100 to>1000 atm.) at 100 to 400 C is described. Halogens and other volatile non- metals that react withwater to form acids (S, Te, P, ammonium chloride, carbon tetrachloride, alkyl halides, etc) are the

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preferred catalysts. The synthesis requires relatively high-purity CO (see also two patents b yCarpenter).

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HYDRIDE ION COAL LIQUEFACTION STUDIES

Appell, H. R.; Wender, I. "The Hydrogenation of Coal with Carbon Monoxide and Water", Am.Chem. Soc. Fuel Div. Prepr. 1968, 2, 220-224.

This is one of the first reports in modern t imes (post 1925) of attempts to liquefy coal with carbonmonoxide and water. The authors found rapid conversion of lignite to benzene solubles (~90%in 10 min) with maximum conversion occurring at ~380 C. This work is the basis for mos tsubsequent CO and hydride ion coal liquefaction work.

Chan, J.; Jackson, W.; Marshall, M. "Reactions of Brown Coals with CO-H O in the Presence of2

Alkaline Catalysts in a Well-Stirred Autoclave with Hot-Charge Facility", Fuel 1994, 73, 1628-1631.

This article discusses a 300 mL autoclave study of coal conversion and asphaltene yields for 3coals treated with var ious WGS catalysts. This aqueous-phase work was conducted at a waterto dry coal ratio of 2.5 to 4. Slightly higher coal conversion (91 vs 88%) was obtained at 355 Cversus 375 C , but oil yields were significantly higher at the higher temperature. Hydrocarbongas make was approximately 0.5%. No significant differences in yields were observed despitesignificant differences in the carbon and hydrogen contents of the starting coals.

Del Bianco, A.; Girardi, E. "Studies o n Coal Reactivity by Using the CO/H O/Base System", Fuel2

Processing Technology 1989, 23, 205-213.

This paper deals with batch liquefaction tests with Illinois 6 coal in a 30 mL reactor without the useof an organic solvent. Reactions were conducted at 400 C for one hour to ensure a TH Fconversion of greater than 90%. The study focused on the relationship between the extent o fwater-gas-shift (WGS) reaction and the production of heptane and toluene soluble material s(i.e., oils and asphaltenes). It was found that oils and asphaltene yields were maximized at a40 to 80% range of WGS conversion. When the WGS conversion was increased even furthervia the use of a catalyst, oils and asphaltene yields did not improve. The authors conclude thatthe chemical nature of the coal plays an important role in its ability to utilize the formate-reactionintermediate produced by the WGS reaction.

Del Bianco, A.; Girardi, E.; Stroppa, F. "Liquefaction of Sulcis Subbituminous Coal in aCO/H O/Base System", Fuel, 1990, 69, 240-244.2

This paper deals with batch liquefaction tests with an Italian coal in a 30 mL reactor without theuse of an organic solvent. This coal contains 11.2% oxygen and 6.2% sulfur and thus differssignificantly from Black Thunder coal. Tests were conducted at 400 C at residence times o fzero (heat-up only) to two hours with a WGS catalyst. A 94% coal conversion was obtained at30 minutes residence time, while the oil yield increased steadily with res idence time. Interestingly,hydrocarbon gas yield was less than 3% at two hours residence time. As in the previou sDel Bianco paper, high coal conversion was achieved at a ra nge of WGS conversions. Oil yieldsdid not increase with increasing WGS conversion. This leads the author to re-emphasize th eimportance of the coal's ability to utilize the formate-reaction intermediate.

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Farnum, S. A.; Wolfson, A. C.; Miller, D. J.; Ga ides, G. E.; Messick, D. D. "Coal Liquefaction withC Labeled Carbon Monoxide", Am. Chem. Soc. Fuel Div. Prepr. 1985, 30, 354-358. 13

The authors used C labeled CO in the liquefaction of li gnite, subbituminous coal and bituminous13

coal at 1000 psig (cold) and 300 C, 340 C and 380 C for 1 hour. Although conversions to THFsolubles of up to 87% were observed, no incorporation of C into products other than CO was13

2

found. They conclude that these data are consistent with the formation of active “nascent ”hydrogen from the closely bound coal water via the water-gas shift reaction. The data also areconsistent with hydride ion coal liquefaction.

Gentzis,T.; Parker, R.; Simpson, P. "Liquefaction of Black Thunder Coal, Part 1: Petrographi cExamination of Residues", Fuel 1995, 74, 1599-1610.

This is the first of two papers by Alberta Research Council (ARC) personnel on their work withBlack Thunder coal as part of DOE's Advanced Concepts Program (see also Parker below). Itdeals with optical microscopy of the THF insolubles from ARC's batch and continuous benc hscale liquefaction tests. The authors, who determined the vitroplast, inertinite, mesophase andmineral matter content of various insoluble residues, conclude that petrography can be used to"guide process control an provides valuable information in coal liquefaction studies."

Horvath, I. T.; Siskin, M. "Direct Evidence for Formate Ion Formation During the Reaction of Coalswith Carbon Monoxide and Water", Energy & Fuels 1991, 5, 932-933.

The authors present direct high pressure C-NMR evidence for the existence of formate ion s13

during high-pressure (700 psi cold CO) liquefactio n of lignite and subbituminous coals at 315 C.

Hulston, C.; Redlich, P .; Jackson, W.; Larkins, F.; Marshall, M. "Reactions of Coals of DifferentRanks with CO-H Mixtures with and without Sodium Aluminate", Fuel 1995, 74, 1870-1874. 2

This article, discusses a microautoclave study of coal conversion and asphaltene yields fo reight coals (including Wyoming subbitu minous coal) treated with H -CO-water-sodium aluminate2

(a water gas shift catalyst) mixtures. Thi s aqueous phase work was conducted at a water to drycoal ratio of 2.5. Reactor conditions were held constant at 365 C, 30 minutes residence timeand 3 MPa cold gas pressure. The ratio of H to CO was the main study variable .2

Interestingly,some of the coals were acid washed to remove inorganics and minerals that couldhave catalytic effects. The Wyoming coal , which had a feed ash content of 3.2%, gave 50 %coal conversion with pure CO and only 25% conversion with pure H . The data in this article are2

not particularly useful to this project.

Jackson, W. R.; Larkins, F. P.; Stray, G. J. "Reac tions of Coal with Carbon Monoxide/ DeuteriumOxide in the Presence of Alkaline Catalysts. Establishment of Mineral Matter in the Coal as theOrigin of Kinetic Isotope Effects", Fuel 1992, 71, 343-345.

This article describes experiments in which fre sh coals and acid washed coals were reacted withCO and water (H O and DO) in the presence of sodium car bonate to verify kinetic isotope effects2

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that were reported by Ross, Green, Mansani and Ham No. 1. The isotope effect was verified insome of the fresh coals, but it was not observed in any of the acid washed coals. The authorsconclude that the observed kinetic isotope effects are probably cau sed by the presence of certainmineral matter (possibly oxidized pyrite species) in the coal and that the reason for the highe rtoluene soluble conversions found by Ross et al. with DO were a result of relatively lowe rconsumption of CO in the water-gas shif t reaction with DO due to this mineral matter catalysis ofthe shift reaction. This work does not, however, rul e out the formate ion liquefaction intermediateproposed by Ross et al.

Knudsen, C. L.; Willson, W. G.; Baker, G. G. "Hydrogen-Carbon Monoxide Reactions in Low -Rank Coal Liquefaction", Am. Chem. Soc. Fuel Div. Prepr. 1981, 26, 132-141.

The authors report results of one-liter batch autoclave studies in which North Dakota (Beulah )lignite was reacted in CO/H /water mixtures. They found that the CO and water are the kinetically2

favored reactants over pure H at all temperatures from 350 C to 480 C for the liquefaction of2

lignite. They found the amount of CO consumed to be related to the quantity of coal fed an dsurmised that the driving force for CO/water liquefaction is the number of reactive sites initiallypresent in the coal.

Lim, S.; Rathbone, R.; Rubel, A.; Givens, E.; Derbyshire, F. "Carbon Monoxide Pretreatment ofSubbituminous Coals", Energy & Fuels 1944, 8, 294-300.

This paper discusses low-temperature treatment (e.g., 300 C) with water, CO, and a base. Itwas intended to subsequently process the product from this step at normal coal liquefactio nconditions. In the low-temperature treatment step, low THF conversions (20%) were obtained.This work is not particularly relevant to our current high coal-conversion work.

Oelert, H; Siekmann, R. "Liquefaction of Coal and Related Materials", Fuel 1976, 55, 39-42.

This article discusses processing different feedstocks (i.e., coal, pine needles) with water an dhydrogen or CO at a temperature of 380 C for 30 minutes. No organic solvent is used. Production of benzene-soluble products increases with increasing H/C and O/C contents of thefeeds. A 54% benzene-soluble yield was achieve d with the lignite tested. THF coal conversionswere not measured.

Parker, R.; Carson, D.; Gentzis, T. "Liquefaction of Black Thunde r Coal, Part 2: Process Severity,Gas Composition & Catalyst Selection", Fuel 1995, 74, 1611-1617,

ARC used Black Thunder coal, a one liter stirred autoclave, a Wilsonvi lle derived vehicle solventfrom Run 263, a CO-steam feed gas and operated at low temperatures (370 to 410 C). Theeffects of various water gas shift catalysts (including none) on coal conversion ,oil-asphaltene-preasphaltene yields, hydrocarbon gas yield and hydrogen consumption ar ereported. Hydrogen to CO ratio was also varied. The heat-up time for the autoclave was 60 to75 minutes. Seven new references are cited on CO-coal systems.

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Parnes, Z. N.; Lipovich, V. G.; Kalechits, I. V.; Kalinkin, M. I.; Kdchetygov, V. D.; Korobkov, V. "Liquefaction of Coal Under Conditions of Ionic Hydrogenation", English Translation of KhimiyaTverdogo Topliva 1985, 19, 80-84.

Compares results of traditional high H pressure thermal liquefaction with lower temperatur e2

(T <400 C) hydride ion liquefaction. These authors used orthophosphoric acid as a proto ndonating catalyst and tetralin as the hydride ion source. They report benzene solubl econversions (at 350 C) of ~50% with the orthophosphoric acid - tetralin system vs ~25% fo rthermal (high pressure hydrogen) dissolution. They conclude t hat such mild condition hydride ionliquefaction may be an attractive alternative to conventional coal liquefaction processes.

Rindt, J. ; Cisney, S. "Low Temperature Liquefaction of North Dakota Lignite", DOE Coal Lique-faction Contractors Conf., November 1983.

This paper is an interesting investigation of lignite liquefaction using CO, steam and a vehicl esolvent (anthracene oil). The effects of temperature, residence time and CO gas concentrationon coal conversion, gas, distillate and soluble resid yields are reported. No water gas shif tcatalyst is used.

Ross, D. S.; Blessing, J. E. "Isopropyl Alcohol as a Coal Liquefaction Agent", Am. Chem. Soc.Fuel Div. Prepr. 1977, 22, 208-213.

The work reported in this article showed that isopropyl alcohol can be an effective coal (IllinoisNo. 6) liquefaction agent that is promoted by the presence of bases. Pyridine solubl econversions of up to 97% were reported. A proposed mechanism for this lique faction was hydrideion transfer from the alcohol or alkoxide salt to the coal.

Ross, D. S.; Blessing, J. E. "Hydroconversion of Bituminous Coal with CO/Water", Fuel 1978, 57,379-380.

This article presents data on 300 mL autoclave studies of Illinois 6 coal liquefaction at 400 C and20 minutes using CO, water, and KOH. No organic solvent was used. A pyridine solubl econversion of approximately 98% was obtained with pure CO o r a 50-50 mixture of hydrogen andCO. Conversion to benzene solubles was higher (51 vs 40%) with pure CO. A hydride-transfer-agent mechanism is proposed. The authors note that almost all the hydrogen produced fromthe CO shift reaction shows up as hydrogen gas, indicating that little, if any, hydrogen has beenadded to the coal.

Ross, D.; Blessing, J.; Nguyen, Q.; Hum, G. "Conversion of Bituminous Coal in CO/Wate rSystems, Part 2: pH Dependence", Fuel 1984, 63, 1206-1210.

This article presents data on 300 mL autoclave studies of Illinois 6 coal liquefaction at 400 C and20 minutes using CO, wate r and KOH. No oil solvent was used. It was found that a starting pHof at least 12.6 was needed to achieve a 50% conversion to benzene solubles. A test was alsorun with potassium formate in the a bsence of CO to test the formate ion conversion mechanism.

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This test gave much poorer results than expected due to suspected premature format edecomposition.

Ross, D. S.; Hum, G. P. "Oxygen Loss and the Conversion of Coal", Proc. 1985 Intl. Conf. onCoal Science 1985, 146-148.

This is one of the earliest reports by Ross et al. that describes the use of CO/water liquefactionto determine the form of the coal oxygen. The work reported here was d one in a stirred autoclavewith Illinois No.6 coal at 400 C in CO and in N . It is relevant to the Novel Liquefaction Concepts2

work because it involves hydride-ion (CO/water) liquefaction. They conclude that oxygen lossand coal conversion are not directly related and that about half the O-loss occurs below 50 %conversion to toluene solubles. They also c onclude that, contrary to earlier belief, the remainderof the coal oxygen is not etheric. Ross, Green, Mansani and Hum No. 2 is a more recent updateof this work.

Ross, D. S.; Hum, G. P.; Miin, T-C.; Green, T. K.; Mansani, R. "Supercritical Water/C OLiquefaction and a Model for Coal Conver sion", Fuel Processing Technology 1986, 12, 277-285.

The authors studied the liquefaction of Illinois No.6 coal in water/CO and CO/D O at 400 C.2

Conversion was found to be limited by the reduction chemistry kinetics and they proposed amodel in which coal is partitioned during conversion between a reduction step and conversionto char. The also report that irreversible consumption of OH by CO is a limiting feature i n-

2

water/CO liquefaction.

Ross, D. S.; Green, T. K.; Mansani, R.; Hum, G. P. "Coal Conversion in CO/Water. 1. ConversionMechanism", Energy & Fuels 1987, 1, 292-294.

Liquefaction of Illinois No. 6 coal with CO/water in a 300 mL stirred autoclave at 400 C isdescribed. This work also confirms the superiority of CO/water over H as a conversion medium.2

A significant kinetic hydrogen isotope e ffect was observed that the authors ascribed to formationof a formate intermediate. This effect was later (Jackson, Larkins, and Stray) shown to resul tfrom catalysis of the water-gas shift reaction by certain inorganic constituents of the coal.

Ross, D. S.; Green, T. K.; Mansani, R.; Hum, G. P. "Coal Conversion in CO/Water. 2. OxygenLoss and the Conversion Mechanism", Energy & Fuels 1987, 1, 292-294.

This paper is a later and more complete discussion of CO/water liquefaction experiments usedto determine the source of the oxygen in Illinois No. 6 coal (Ross and Hum). The conclusion isthat much of the oxygen in these coals is present in tightly bound water and not etheric.

Severson, D.; Souby, A.; Baker, G. "Continuous Liquefaction of Lignite in a Process DevelopmentUnit", ACS Fuel Division Prepr. 1977, 22, 161-182.

This paper describes 0.6 T/D PDU test work at the University of North Dakota. A single stagereaction system with a CO-hydrogen gas system and a distillate recycle solvent was used in an

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effort to produce solvent refined lignite (SRL). While somewhat dated, this paper does containsome interesting data.

Severson, D.; Souby, A.; Owens, T. "Extended Operation of a Solvent Refined Lignite ProcessDevelopment Unit", Energy Sources 1982, 6, 173-192.

This paper appears to be an update of the prev ious work done at the University of North Dakota.Test work was done at a very high 3/1 solvent-to-coal ratio and there was a significant build-upof solids in the reactor.

Shibaoika, M.; Takemura, Y.; Russell, N; Ouchi, K. "Migration and Diffusion of Molybdenu mCatalyst During CO/Steam Coal Hydrogenation", Fuel 1986, 65, 362-366.

This paper discusses the liquefac tion of an Australian bituminous coal with various molybdenumcatalysts at 350 to 400 C using naphthalene as an organic solvent. The tests were made in a146 mL autoclave with a heat-up time of approximately 90 minutes and a 60 minute residencetime at reaction temperature. Conversions to benzene solubles was measured.

Sofianos, A.; Butler, A.; Louwrens, H. "Catalytic Liquefaction of South African Coals Using th eCarbon Monoxide/Water System, Part 1: Pyrite Catalysis", Fuel Processing Technology 1989,22, 175-188.

Four bituminous coals were liquefied in one- and two-liter autoclaves with 10 wt % pyrite and noorganic solvent. The use of pyrite increased coal conversion from 80% to 95% at 400 C and30 minutes residence time. Interestingly, coal conversion was lower at 450 C than at 400 C(89 vs 95%), but oil yield was higher at 450 C (31 vs 24%).

Sondreal, E.; Knudson, C.; Schiller, J.; May, T. "Development of the CO-Steam Process fo rLiquefaction of Lignite and Western Subbituminous Coals", Ninth Biennial Lignite Symposium ,Grand Forks, ND, 1977.

This is an early review paper describing the prog ress made on the CO-steam process from 1975to 1977 at the Grand Forks Energy Research Center. Recommended reactor conditions were460 C and 4,000 psig. This is dated information but there is some interesting information o nmethane formation from the solvent and the solvent-coal mixture versus reactor temperature andresidence time, and on the sodium content of the high molecular weight residual product as afunction of reaction conditions.

Sondreal, E.; Wilson, W.; Stenberg, V. "Mechanisms Leading to Process Improvements in LigniteLiquefaction Using CO and H S", Fuel 1982, 61, 925-938. 2

This paper, another "must read", gives the then current understanding of the SRL process .Bottoms recycle and the use of H S are two differences from previously reported work. A tutorial2

is presented on the characteristics of North Dakota lignite and its behavior in liquefaction. Theimportance of inherent coal moisture, CO as part of the reducing gas and the use of H S as a2

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liquefaction catalyst are discussed. More importantly , a discussion of liquefaction mechanismswith CO is presented ( 934-936). The author concludes that, "the formate intermediate, of andby itself, is insufficient in explaining the effectiveness of CO and steam for liquefying lignite".

Takemura, Y.; Ouchi, K. "Catalytic Liquefaction of Various Coals Using a Mixture of Carbo nMonoxide and Water", Fuel 1983, 62, 1133-1137.

This paper reports on microautoclave studies of the liquefaction of 20 different coals with CO ,water and a ground Co-Mo catalyst for water gas shift. Tests were made with and without avehicle solvent at 400 °C for 60 minutes. For an Aus tralian subbituminous coal, coal conversionwas much greater with the use of the vehicle solvent even though the extent of water gas shiftwas less.

Watanabe, Y.; Yamada, H.; Kawasaki, N.; Wada, K.; Mitsudo, T. "Fe(CO) - Sulfur-Catalysed5

Coal Liquefaction Using H O-CO as a Hydrogen Source", Chemistry Letters 1992,111-114.2

This paper (the first of three by these authors) discusses microautoclave studies of Yallour n(Victorian) brown and Wandoan subbitumin ous coals, in a reaction system of CO, water, methylnaphthalene, sulfur, and Fe(CO) or Fe O . The use of sulfur and Fe(CO) increased Yallourn5 2 3 5

coal conversion from 57 to 96% at 375 C and 60 minutes residence time. The conversion withFe O was 75% at the same conditions . Oil yields also were improved with the use of sulfur and2 3

Fe(CO) . Less dramatic improvements were achieved with Wandoan coal, and the difference5

in iron sources was much less critical. The iron dosage used was approximately 3% on coal .Coal conversion dropped from 96% to 89% when the CO pressure (cold) was lowered fro m7 to 5 MPa. Coal conversion also was ad versely affected by reaction temperatures greater than380°C and water to coal ratios less than 0.6/1 and greater than 1.1/1. No explanation for th ewater effect was given.

Watanabe, Y.; Yamada, H.; Kawasaki, N.; Hata, K.; Wada, K.; Mitsudo, T. "Fe(CO) -5

Sulfur-Catalyzed Coal Liquefaction of Yallourn Coal in Syngas-H O Systems", Chemistry Letters2

1993, 275-278.

This article is a follow-up on the previous paper. Results with syngas (equimolar CO and H )2

were slightly better at similar reaction conditions than with CO alone. The water-to-coal ratio ,again, had an effect on coal conversion. As m entioned in the review of the Watanabe Fuel 1996article, the cost of using Fe(CO) could be a significant factor.5

Watanabe, Y.; Yamada, H.; Kawasaki, N.; Hata, K.; Wada, K.; Mitsudo, T. "Fe(CO) -5

Sulfur-Catalysed Coal Liquefaction in H O-CO Systems", Fuel 1996, 75, 46-50.2

This paper discusses microautoclave studies of four coals, including Wyoming subbituminous,in a reaction system of CO, water, methyl naphthalene, sulfur and Fe(CO) . The use of sulfur5

and Fe(CO) increased Wyoming coal conversion from 61 to 87% at 375 C and 60 minutes5

residence time. The iron dosage used is approximately 3% on coal. The cos t of using a Fe(CO) 5

is not known but thought to be significant.

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Wu, M. M.; Winschel, R. A. "Coal Desulfurization and Coal Conversion with Formates" ,Proceedings, Second Annual Pittsburgh Coal Conference, Pittsburgh, PA, 1985, 115-126,

The use of formates as agents for chemical desulfurization of coal and as liquefaction agents isdiscussed. This work summ arizes some of CONSOL’s earlier work with hydride ion liquefactionagents, including formic acid, ammonium formate, formamide and methyl formate. It was reportedthat staged heating (rapid heat-up to ~300 C followed by slow heating to 400 C) drasticallyimproved conversion with lignite and methyl formate. Lignite and bituminous coal conversionsto THF solubles of over 90 wt % were reported with each hydride ion source.

Zemskov, V. V. "Brown Coal Li quefaction Using Ion Hydrogenation", Proceedings, InternationalConference on Structure and Properties of Coal, Warsaw, Poland, 1991, 129-130.

This presented paper compares hydride ion liquefaction of brown coal (lignite) in a 1:1.3:0. 7mixture of coal : tetralin : formic acid with therm al dissolution in tetralin alone at 250-500 C. Theauthor concludes that higher conversions are attained at lower temperatures (350 C or less) inthe formic acid system than by therma l treatment alone. He also concludes that formic acid onlypromoted coal dissolution below ~420 C because formic acid decomposes above ~350 C.Unfortunately, the author reports no data to support his conclusions.