process simulation in refineries sampler
DESCRIPTION
Process Simulation Brief DescriptionTRANSCRIPT
PROCESS SIMULATION
REFINERY PROCESSES
SAMPLER
Modelling and Optimization
John E. Edwards Process Simulation Engineer, P & I Design Ltd
First Edition, June 2013 P&I Design Ltd
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Released by P & I Design Ltd 2 Reed Street, Thornaby TS17 7AF www.pidesign.co.uk Private distribution only Copyright © P & I Design Ltd 2012 [email protected] Printed by Billingham Press Ltd, Billingham TS23 1LF
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Process Simulation Refinery Processes
Contents Section 1 Refinery Processes 5 Section 2 Thermodynamics 9 Section 3 Crude Column 13 Section 4 Vacuum Still 23 Section 5 Splitting and Product Purification 27 Section 6 Hydrotreater 43 Section 7 Catalytic Reformer 47 Section 8 Amine Treatment 53 Section 9 Miscellaneous Applications 57 Section 10 General Engineering Data 59 Section End 70
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Preface The process industry covers a broad spectrum of activities that involve the handling and treatment of gases, liquids and solids over a wide range of physical and processing conditions. This manual provides a comprehensive review of the fundamentals, definitions and engineering principles for the study of processes encountered in hydrocarbon processing using steady state simulation techniques. Applications are presented for a wide range of processing units involving design and operations. Process simulations are carried out using CHEMCAD™ software by Chemstations, Inc. of Houston. This manual has been developed with the full support of Chemstations simulation engineers based in Houston. The simulation of crude distillation at atmospheric pressure, vacuum distillation and sour gas amine treatment is covered in Section 13 Process Measurement and Control of the book “Chemical Engineering in Practice” by J.E.Edwards. This manual includes these topics and extends the study to other refinery processes including splitters, stabilizers, hydrotreaters and reformers. Thermodynamics are reviewed with special reference to the application of pseudocomponent curves and crude oil databases Each topic is in the form of a condensed refresher and provides useful practical information and data. Each section is numbered uniquely for contents and references, with the nomenclature being section specific. The references are not a comprehensive list and apologies for unintended omissions. Reference is made to many classic texts, industry standards and manufacturers’ data. Information has been mined from individual project reports and technical papers and contributions by specialists working in the field. .
The Author
http://uk.linkedin.com/pub/john-edwards/1b/374/924 John E.Edwards is the Process Simulation Specialist at P&I Design Ltd based in Teesside, UK. In 1978 he formed P&I Design Ltd to provide a service to the Process and Instrumentation fields. He has over fifty years’ experience gained whilst working in the process, instrumentation and control system fields.
Acknowledgements A special thanks to my colleagues at Chemstations, Houston, who have always given support in my process simulation work and the preparation of the articles that make up this book: N.Massey, Ming der Lu, S.Brown, D.Hill, A.Herrick, F.Justice and W.Schmidt of Germany Also thanks to my associate P.Baines of Tekna Ltd for help with the organic chemistry topics.
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Section 1
Refinery Processes
References
1. Shrieve, “Chemical Process Industries”, Chapter 37, 5th Edition, McGraw Hill, 1984. 2. J.A.Moulijn, M.Makkee, A.Van Diepen, “Chemical Process Technology”, Wiley, 2001. 3. G.L.Kaes, “Refinery Process Modelling”, Athens Printing Company, 1st Edition, March 2000. 4. W.L.Nelson, “Petroleum Refinery Engineering”, 4th Edition, McGraw Hill, 1958.
Overview Most refinery products are mixtures separated on the basis of boiling point ranges. The block diagram, by API, shows overall relationship between the refining processes and refined products.
Refining is a mature, complex and highly integrated operation. Columns with a wide variety of internals are used in many stages of the process. Fractional distillation under vacuum and pressure conditions is used to separate components. Light ends are steam stripped and the heavy ends are vacuum distilled at reduced the temperatures. Stabilizers are used to remove light ends, including LPG, to reduce the vapor pressure for storage and subsequent processes. Absorbers and strippers are used to remove unwanted components such as sulphur. Simple distillation processes do not produce sufficient gasoline above the minimum required octane number. This is achieved by converting heavy to light hydrocarbons using catalytic processes including fluidic catalytic cracking (FCC), hydrotreating, hydrocracking, catalytic reforming and alkylation.
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Crude petroleum consists of thousands of chemical species. The main species are hydrocarbons but there can be significant amounts of compounds containing sulphur (0-6%), oxygen (0-3.5%) and nitrogen (0-0.6%). The main groups are: Aliphatic or open chain hydrocarbons as detailed in the table:
Aliphatic or open chain hydrocarbons (hc)
Descriptor Properties Class Formula Member
-ane saturated hc, unreactive paraffins CnH2n+2 C2H6 ethane
-ene unsaturated hc, forms additive compounds
olefines CnH2n C2H4 ethylene
acetylenes CnH2n-2 C2H2 acetylene
-ol reactive, OH replaced alcohols, phenols RCH3OH C2H5OH ethyl alcohol
-one additive compounds ketones RR1.CO (CH3)2.CO acetone
n-paraffin series or alkanes (CnH2n+2) This series has the highest concentration of isomers in any carbon number range but only occupy 20-25% of that range and make low octane gasoline. Most straight run (distilled directly from the crude) gasolines are predominately n-paraffins. The light ends primarily consist of propane (C3H8), n-butane (C4H10) together with water which are defined as pure components. iso-paraffin series or iso-alkanes (CnH2n+2) i-butane (C4H10) is present in the light ends but these compounds are mainly formed by catalytic reforming, alkylation or polymerization. olefine or alkene series (CnH2n) This series is generally absent from crudes and are formed by cracking (making smaller molecules from larger molecules). They tend to polymerize and oxidize making them useful in forming ethylene, propylene and butylene.
Ring compounds Naphthene series or cycloalkanes (CnH2n) These compounds are the second most abundant series of compounds in most crudes. The lower members of this group are good fuels and the higher members are predominant in gas oil and lubricating oils separated from all types of crude. Aromatic series Only small amounts of this series occur in most common crudes but have high antiknock value and stability. Many aromatics are formed by refining processes including benzene, toluene, ethyl benzene and xylene. Lesser Components Sulfur has several undesirable effects including its poisonous properties, objectionable odour, corrosion, and air pollution. Sulfur compounds are removed and frequently recovered as elemental sulfur in the Klaus process. Nitrogen compounds cause fewer problems and are frequently ignored. Trace metals including Fe, Mo, Na, Ni and V are strong catalyst poisons and cause problems with the catalytic cracking and finishing processes and methods are used to eliminate them. Salt, which is present normally as an emulsion in most crudes, is removed to prevent corrosion. Mechanical or electrical desalting is preliminary to most crude processing. Crude oil is classified on the basis of density as follows: Light less than 870 kg/m3 >31.1° API Medium 870 to 920 kg/m3 31.1° API to 22.3° API Heavy 920 to 1000 kg/m3 22.3° API to 10° API Extra-heavy greater than 1000 kg/m3 <10° API Bitumen Heavy or extra-heavy crude oils, as defined by the density ranges given, but with viscosities greater than 10000 mPa.s measured at original temperature in the reservoir and atmospheric pressure, on a gas-free basis
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Natural Gas Light hydrocarbon mixture that exists in the gaseous phase or in solution in crude oil in reservoirs but are gaseous at atmospheric conditions. Natural gas may contain sulphur or other non-hydrocarbon compounds. Natural Gas Liquids Hydrocarbon components recovered from natural gas as liquids including ethane, propane, butanes, pentanes plus, condensate and small quantities of non- hydrocarbons. Atmospheric and vacuum distillations produce the different fractions as detailed in the table below.
Crude Petroleum Fractional Distillation
Temperature <30ºC 40-70 ºC 70-120ºC 120-150 ºC 150-300 ºC >350 ºC Residue
Description Gaseous
Hydrocarbon Gas oil Naptha Benzene Kerosene
Heavy oils
Asphalt or
Bitumen
Density 0.65 0.72 0.76 0.8
Composition C3H8, C4H10 C5H12, C6H14
C6H14, C7H16 C8H18
C8H18, C9H20
C10H22, C11H24
C12H26 to C18H36
C18H38 to C28H58
Applications
Gas fuel or enrichment
General solvent, aviation
spirit
Solvent for oils, fats &
varnishes
Solvent for oils, fats & varnishes
Home heating Jet fuel
Diesel, fuel oils
Roads, Wax
paper
Gasoline, contains C6H14, C7H16, C8H18 40-180 ºC
Further fractionation of the 70 to 150ºC cut is required to obtain the naptha and benzene cuts. Vacuum distillation of the topped crude is required to obtain Light Vacuum Gas Oil (LGVO) and Heavy Vacuum Gas Oil (HVGO) When the difference in volatility between components is small a solvent of low volatility is added to depress the volatility of one of the components. This process is known as extractive distillation. Butenes are separated from butanes using this method with furfural as the extractant. When a high volatility entrainer is used the process is known as azeotropic distillation. Anhydrous alcohol is formed from 95% aqueous solution using benzene to free the azeotrope and high purity toluene is separated using methyl ethyl ketone as the entrainer.
Typical Crude Oil Products Profile Ref EIA March 2004 Data
Product Refined gallons/barrel (gal/bbl)
Gasoline 19.3
Distillate Fuel Oil (Inc. Home Heating and Diesel Fuel) 9.83
Kerosene Type Jet Fuel 4.24
Residual Fuel Oil 2.10
Petroleum Coke 2.10
Liquified Refinery Gases 1.89
Still Gas 1.81
Asphalt and Road Oil 1.13
Petrochemical Feed Supplies 0.97
Lubricants 0.46
Kerosene 0.21
Waxes 0.04
Aviation Fuel 0.04
Other Products 0.34
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Section 2
Thermodynamics
References
1. G.L.Kaes, “Refinery Process Modelling”, Athens Printing Company, 1st Edition, March 2000 2. American Society for Testing and Materials (ASTM International), Standards Library Global K and H Models The following table gives a summary of suitable K and H models for common refinery processes.
Refinery Processing Thermodynamic Models Summary
Process K Model H Model (Forced)
Crude Atmospheric Distillation Grayson Streed Lee Kessler
Vacuum Distillation ESSO Lee Kessler
Hydrotreater SRK SRK
Sour Gas Treatment Amine Amine
FCC Gas Treatment Peng Robinson Peng Robinson
Propylene Splitter Peng Robinson Peng Robinson
Compression BWRS BWRS
Grayson Streed K model is primarily applicable to systems of non-polar hydrocarbons. It is good for modelling hydrocarbon units, depropanizers, debutanizers, or reformer systems. The approximate range of applicability is as follows: Temperature Range Pressure Range 0 to 800°F < 3000 psia -18 to 430°C < 20000 kPa ESSO K model predicts K-values for heavy hydrocarbon materials at pressures below 100 psia. The average error for pure hydrocarbons is 8% for p* > 1 mmHg, and 30% for p* between 10E-06 and 1 mmHg according to API Technical Data Book Vol 1. It is good for modelling vacuum towers. Lee Kessler H model is good for hydrocarbon systems. AMINE K model is based on the Kent Eisenberg method to model the reactions with diethanolamine (DEA), monoethanolamine (MEA), methyl diethanolamine (MDEA) being included. The chemical reactions in the CO2-Amine system are described by the following reactions:
RR'NH2+ ↔ H+ + RR'NH RR'NCOO + H2O ↔ RR'NH + HCO3 CO2 + H2O ↔ HCO3- + H+ HCO3- ↔ CO3- - + H+ H2O ↔ H+ + OH-
Where R and R' represent alcohol groups. The reaction equations are solved simultaneously to obtain the free concentration of CO2. The partial pressure of CO2 is calculated by the Henry's constants and free concentration in the liquid phase. The AMINE K Model in CHEMCAD treats the absorption of CO2 in aqueous MEA as a fast chemical reaction, in other words, gas film controlled implying a very low stripping factor. However it is known that this process is liquid film controlled since Henry’s Law controls the diffusion of CO2 into the liquid prior to chemical reaction taking place.
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Section 3
Crude Column
Crude Column Simulations
Case/File Name Description
R3.01 Crude Column Feed
References
1. H.Kister, “Distillation Design”, McGraw-Hill, ISBN 0-07-034909-6
2. G.L.Kaes, “Refinery Process Modelling”, Athens Printing Company, 1st Edition, March 2000
3. W.L.Nelson, “Petroleum Refinery Engineering”, 4th Edition, McGraw Hill, 1958
Process Description The simplified process flow diagram shows the basic layout for the crude and vacuum distillation units. Desalted crude is preheated with the pump around and topped crude heat exchangers prior to being heated to ~620ºF in the direct fired furnace. Above this temperature thermal decomposition (cracking) will take place resulting in increased light ends and fouling of heat exchange surfaces due to carbon based deposits. The following initial guidelines are suggested: 1. For paraffin based crudes at moderate furnace temperatures, an estimated cracked gas
rate of 5.0 SCF/bbl (42 gal/bbl) crude oil is reasonable. 2. For asphalt based crude oil a cracked gas production of 2.5 SCF/bbl crude oil is suggested. 3. The cracked gas may be given an arbitrary composition as follows:
50 mol% methane, 40 mole% ethane, and 10 mole% propane. The feed to the atmospheric crude tower is a mixed vapor-liquid phase of ~0.4 vapor fraction. The vapours flow upwards and are fractionated to yield the products. Crude towers are typically 4m diameter, 20–30m in height with 15–30 trays.
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A typical Process Flow Diagram for a crude unit, including pump-around circuits and side strippers, is shown. The column is modelled on the basis of theoretical stages, as opposed to actual trays, so it is necessary to apply tray efficiency η to translate the actual trays NA to theoretical trays NT where η=NT/NA. Note that commercial simulators provide various tray efficiency models, which are not suitable for crude distillation columns. Tray efficiency η should be based on experience. The relationships between NA and NT are indicated in the diagram.
The liquid product sidestreams contain light hydrocarbons which must be removed to meet the initial boiling point specification for the products. The liquid sidestreams are fed to strippers that use either a reboiler or steam to strip out these light components which are returned to the crude tower. Current preference is to use reboiled side strippers for the lower boiling products to reduce the heat load on the crude tower condenser and the sour water stripper. Side strippers are typically 1-2m diameter, 3m in height with 4–8 trays representing 2–3 theoretical stages. Height limitations can be met by using structured packing which has high capacity and low HETP values as compared to trays. Pumparound cooling circuits provide reflux to remove the latent heat from hot flash zone vapors and condense the side products. A pump-around zone may be considered equivalent to an equilibrium flash where equilibrium liquid is recirculated. The large flow of pump-around liquid creates a region of constant liquid composition that eliminates fractionation. The heat removed preheats the crude feed.
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Section 4
Vacuum Still
Vacuum Still Simulations
Case/File Name Description
R4.01 Vacuum Unit
Vacuum distillation is used to separate the high boiling bottoms from the crude column. The Vacuum Unit process flow diagram is shown with distillation UnitOp 1 selected as Tower+.
The thermodynamic selection is K Model ESSO and H Model Lee Kessler. The feed is defined by the following specification: Feed rate 360 m3/day Bulk gravity 0.9168 specific gravity Feed temperature 150ºF Feed pressure 58 psia Distillation curve volume % based on TBP at 1 atm
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The column specifications are:
Vacuum Column Data
Description Specification
Number of strippers 0
Number of pumparounds 2
Number of exchangers 1
Number of side products 2
Stages Theoretical 8 Feed 8
Column pressures Top 30 mmHg dP 35 mmHg
Stripping Steam condition 335ºF and 115 psia
Bottom steam flow 166.67 lb mol/h
Condenser Total
Reboiler None
Pumparound 1
Stages Draw-3 Return-1
Flow 276218 kg/h Phase liquid
Duty 0 MJ/h
Pumparound 2
Stages Draw-5 Return-4
Flow 538139 kg/h Phase liquid
Duty 0 MJ/h
Side Product Draw 1
Stage 3
Flow 72 m3/h Phase liquid
Side Product Draw 2
Stage 5
Flow 213 m3/h Phase liquid
Side Heat Exchanger
Stage 8 No duty (Feed stage)
Stage Specifications
Stage 3 1 kmol/h Liquid flow
Stage 5 85 m3/h Liquid Flow
Stage 8 69 m3/h Liquid Flow
Pseudocomponent Curves allow group plots to be generated for the streams:
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Section 5
Splitting and Product Purification
Splitting and Product Purification Simulations
Case/File Name Description
R5.01 Deethanizer
R5.02 Debutanizer Depropanizer
R5.03 Debutanizer Reflux Depropanizer
R5.04 C3 Splitter
R5.05 C4 Splitter
R5.06 C4 Splitter Tray Column
R5.07 Kerosene Splitter
References
1. H.Kister, “Distillation Design”, McGraw-Hill, ISBN 0-07-034909-6 2. G.L.Kaes, “Refinery Process Modelling”, Athens Printing Company, 1st Edition, March 2000 3. G.L.Kaes, “Practical Guide to Steady State Modelling of Petroleum Processes
4. H.Kazemi Esfeh and I.Aalipour mohammadi, “Simulation and Optimization of Deethanizer Tower”, 2011 International Conference on Chemistry and Chemical Process, Singapore
Introduction A primary activity in hydrocarbon processing involves the fractionation and purification of light ends using columns, the most common being stabilizers, deethanizers, debutanizers and depropanizers. A typical purification plant schematic is shown:
Deethanizer The deethanizer removes ethane (C2H6) and lighter components which may be fed to the olefines unit for production of ethylene (C2H4) or polyethylene or polypropylene products. Bottoms are fed to the debutanizer. Design for C2 mole fraction or C2/C3 mole ratio in the bottoms. Debutanizer The debutanizer separates mixed LPG product (mostly C3’s and C4’s) and a stabilized condensate (C5+). Design for RVP in bottoms with 12 psia being typical and reflux ratio 0.5 – 1.0 Depropanizer The depropanizer separates propane (C3’s) as overheads from the butane (C4) to the bottoms. Stabilizer Stabilizers are used to remove light ends (mainly C4’s) from condensate to meet Reed Vapour Pressure (RVP) specification for future processing or to allow storage in floating roof tanks. Design for RVP in bottoms with 12 psia being a typical maximum value All purification units use the bottom tray or reboiler temperature and reflux for control. The stabilizer uses bottom tray or reboiler temperature alone as there is no condenser for reflux control. Using these parameters in process simulation allows predicted product properties to be compared against actual process conditions. Simulation parameters can be adjusted to match current behaviour to provide a powerful troubleshooting tool.
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Splitters are used extensively in hydrocarbon processing, including C2’s, C3’s, C4’s and Naphtha. The process simulation methods used are similar to those for the purification process with the CHEMCAD SCDS UnitOp being used.
Tray Column Industry Practice and Efficiencies (1)
Process Actual Trays Overall Efficiency Theoretical Trays
Naphtha Splitter 25 - 35 70 - 75 18 - 25
C2 Splitter 110 - 130 95 - 100 105 - 125
C3 Splitter 200 - 250 95 - 100 190 - 240
C4 Splitter 70 - 80 85 - 90 60 - 68
C2 Splitter (C2H6 – C2H4) This involves the separation of ethylene from ethane using low temperature distillation. The splitter is normally operated at high-pressure, utilizing closed-cycle propylene refrigeration. The objective is
to obtain a high % recovery of high purity ethylene. This process is a high energy user and costly.
C3 Splitter (C3H8 – C3H6) This involves the separation of propylene form propane. High pressure, typically 220 psia, is needed to condense the propylene vapor at ambient temperatures around 40°C and allows the use of cooling water on the condenser. C4 Splitter (iC4H10 – nC4H10) This involves the separation of i-butane form n-butane. Naphtha Splitter Full Range Naphthas (FRN) feed is taken from the crude unit overheads and the splitter separates the light from the heavy naphtha. Light naphtha from the overheads is cooled against the incoming FRN and then condensed in air fin fan coolers and used as reflux or routed to the light naphtha stabilizer column for stabilization and recovery of light ends LPG. The column uses a forced circuit fired reboiler system. The splitter bottoms are pumped via a heat exchanger to recover heat from the Naphtha Hydrotreater hot reactor effluent into a fired furnace to provide the desired reboiler duty to effect the separation of the light and heavy naphthas. Heavy naphtha from the column bottoms is fed to the Naphtha Hydrotreater section and subsequently the Catalytic Reformer feedstock. The thermodynamics suitable for simulating these hydrocarbon mixtures are the equation of states Soave – Redlich - Kwong (SRK) for pressures >1 bar and Peng Robinson for pressures >10 bar.
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Case R5.01 Deethanizer (4)
Ethane is the primary component in the feed to olefin plants for the production of unsaturated hydrocarbons such as ethylene. ���� = ���� + �� Methane and ethane are to be separated from propane using 48 theoretical stages with the feed being introduced on tray15. Integral condenser is stage 1 and reboiler stage 48. Column top pressure is 18.33 bar but tray pressure drop was not included. Feed composition and conditions are shown in Stream 1. Suitable thermodynamics are SRK or Peng Robinson. Reference (4) indicated good agreement with both methods but the prediction, by both methods, of ethane composition in tower bottoms was inaccurate leading to a higher ethane recovery than on plant The column operating conditions are to be established to achieve the following separation.
C2 Splitter Operating Targets
Component Overhead Bottoms
mole fraction mole fraction
methane 0.241 0
ethane 0.738 0.0022
propane 0.0106 0,9914
i-butane 0 0.00555
n-butane 0 0.00085
H2S 0.000047 0
The SCDS convergence parameters were set for a distillate propane composition 1.4 reflux ratio and a bottoms ethane composition 0.0022 mass fraction. The column converged with a reboiler duty of 3461 MJ/h. Tray composition profile is shown.
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Case R5.06 C4 Splitter Tray Column The previous data has been based on an industrial fractionator; reference: Klemola and Ilme, Ind.Eng.Chem.35, 4579 (1996) with tray specification as follows:
Key Tray Specifications
Column Height m 51.8 Downcomer Area (centre)
m2 0.86
Column Diameter m 2.9 Tray Spacing
m 0.6
Number of Trays no 74 Hole Diameter
mm 39
Weir Length (side) m 1.859 Total Hole Area
m2 0.922
Weir Length (centre) m 2.885 Outlet Weir Height
mm 51
Liquid Flowpath Length m/pass 0.967 Tray Thickness
mm 2
Active Area m2 4.9 Number of Valves
no/tray 772
Downcomer Area (side) m2 0.86 Free Fractional Hole Area % 18.82
SCDS simulation model is now changed to Tray Column Mass Transfer and the tray details are entered as shown. Tray efficiency profiles were not entered but 85 to 90% is typical. The side weir dimension is as shown in the diagram below and is not to be confused with side weir length. Note that the Downcomer side area shown in the table is for 2 passes. The total hole area is shown as 0.922m2 which is in ratio to the active area of 4.9m2 giving the free fractional hole area of 18.82%. The simulation is now shows the following results:
1.859
0.335
Area 0.43
2 Passes
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Section 6
Hydrotreater
Hydrotreater Simulation
Case/File Name Description
R6.01 Hydrotreater
References
1. J.A.Moulijn, M.Makkee, A.Van Diepen, “Chemical Process Technology”, Wiley, 2001. 2. G.L.Kaes, “Refinery Process Modelling”, Athens Printing Company, 1st Edition, March 2000. 3. W.L.Nelson, “Petroleum Refinery Engineering”, 4th Edition, McGraw Hill, 1958. Process (1, 2)
Hydrotreaters are used to selectively remove undesirable elements and hydrogen saturate unsaturated components. Reactor pressures vary 500-1000psig and temperatures 550-700°F. Hydroteating involves reaction with hydrogen to remove mainly sulfur, nitrogen and oxygen with some hydrogenation of double bonds and aromatic rings taking place. Hydrotreating is always applied as a pre-treatment to naphtha reforming to protect the catalyst against S-poisoning. Hydrotreating of heavy residues is not considered here. H2 reacts with mercaptans (1H2), thiophenes (3H2) and benzothiphenes (5H2) produce H2S H2 reacts with pyridine (5H2) produces NH3 H2 reacts with phenols (1H2) produces H2O
Hydrotreater Hydrogen Usage and Losses (2)
Units scf/barrel fresh feed sm3/m3 fresh Remarks
Reactions(basis fresh feed 100 – 500 18 -89 Unsaturated components >H2
Solubility(basis fresh feed) 10 - 20 1.8 – 3.6
Purge 40 - 100 7.2 - 18 Depend H2 in makeup gas
Recycle 500 - 1500 89 - 267 Maintain H2/hydrocarbon ratio
The flow sheet is similar for all hydrotreating operations. The liquid feed stock is mixed with a hydrogen-rich gas and preheated by exchange with the reactor effluent. The warm feed is brought to the desired reaction temperature in a furnace and fed to the hydrotreating reactor. The reactor effluent is cooled and the hydrogen-rich gas is separated from the liquid product. The separator liquid is sent to a fractionator for removal of dissolved light hydrocarbon liquids and gases.
19
Case R6.01 Hydrotreater Unit Simulation flowsheet
Simulation Parameters Thermodynamics selection is K-SRK and L-SRK. Pseudocomponents are created for the feed (FEEDC6+) and product (PRODC6+) streams based on predicted molecular weight, API gravity and normal boiling point.
20
Section 7
Catalytic Reformer
Reformer Simulation
Case/File Name Description
R7.01 Catalytic Reformer
References
1. G.L.Kaes, “Refinery Process Modelling”, Athens Printing Company, 1st Edition, March 2000. 2. A.Askari et al, “Simulation and Modelling of Catalytic Reforming Process”, Petroleum & Coal
ISSN 1337-7027. Process Description A traditional reforming process uses three fixed bed reactors in series. Endothermic dehydrogenation reactions take place in the first two reactors requiring fired inter-heaters to raise the temperature for the following reactor with hydrocracking reactions being significant in the final reactor. Reforming catalysts are subject to poisoning by hydrogen sulphide and other sulfur compounds, nitrogen, and oxygen which are removed from the naphtha by mild hydrotreating. The primary reformer feed stock is virgin (uncracked) naphtha from the crude distillation process and other naphtha stocks of suitable boiling point range are acceptable after hydrotreating. The reactions do not occur evenly through the reactors so it is the convention in simulation work to consider all the reaction taking place in the last reactor. Reactors 1 and 2 are set up for mass transfer with the pressure drop being entered and the isothermal mode being used to set the outlet temperature. Using the initial reactor inlet composition for the inter-furnace duty calculations does not result in significant inaccuracies. The final reactor is set up in adiabatic mode with the kinetic reactions specified. Pre-treated naphtha is combined with recycle gas with H2 composition in range 75 to 85 mole % and preheated by exchange with the effluent from reactor 3. Typical reactor pressures and temperature drops are shown:
Operating Data Reactor 1 Reactor 2 Reactor 3
Inlet temperature °F 937 937 937
Inlet pressure psia 413 394 394
Measures ∆T (Typical) °F 60 (90-130) 35 (40-60) 25 (1) (10-20)
Recycle MMSCFD 0.5-1% Naphtha Feed
Catalyst Volume ft3 274 411 910
Note 1 Simulation temperature drop is much larger due to the example reactions considered The temperature drops across the reactors are monitored to track catalyst activity. Separator parameters The separator feed is cooled to 90 - 100°F using air and water coolers and the flash drum pressure was run at 290 psia and with isentropic flash. The hydrogen rich gas stream is used in other refinery operations and compressed and remainder recycled to the process where it combines with the naphtha feed prior to the feed/effluent heat exchanger.
Operating Data Separator
Temperature °F 100
Pressure psia 290
H2 Purity 0.79
21
Stabilizer parameters The liquid is feed to the stabilizer to remove the light ends. Reformer stabilizers generally have 30 to 36 actual trays with overall tray efficiencies in the range 70 to 75%. The primary function is to strip the n-butane from the reformate product. The distillate is sent to a gas recovery plant and the column bottoms product is stabilized reformate.
Operating Data Stabilizer
Number of stages 36
Feed tray 19
Feed temperature °F 297
Tray 1 temperature °F 257
Bottom temperature °F 446 (Simulation 488)
Partial condenser pressure psia 239
To indicate the principles of configuring the catalytic reformer the following stoichiometric equations have been used.
Refer to Section 1 for a more detailed analysis of refinery chemistry. It is recognised that the compositions do not represent typical conditions and is an over simplification of the number of species and the complexity of the kinetic reactions involved. In practices there are many reactant components and intermediate products which make it extremely complicated to study rigorously. To reduce this complication reactants are classified into definite groups as pseudocomponent streams. There are many models available for the study of reaction kinetics including Langmuir-Hinshelwood, Arrhenius and Smith. The Smith model considers the following groupings:
Naphthene + H2 → Paraffin Naphthenes → Aromatics + 3H2
Hydrocracking of paraffin Hydrocracking of naphthenes
Pseudocomponent groupings will include specific boiling ranges such as for C6 to C11 paraffins, C6 to C11 naphthenes, benzene, toluene and C8 to C11 aromatics.
22
Case 7.01 Catalytic Reformer Flowsheet is shown:
Thermodynamics selected: K-Peng Robinson H-Peng Robinson The configuration for catalytic Reactor 3 is shown.
23
Section 8
Amine Treatment
Vacuum Still Simulations
Case/File Name Description
R8.01 Sour Gas Treatment
References
1. A.Kohl and R.Nielsen, “Gas Purification”, Gulf Publishing , 5th Edition, 1997 Process Chemical absorption of CO2 and H2S with amines provides the most cost effective means of obtaining high purity vapor from sour gases in a single step. The process is well established for refinery gas sweetening which are carried out at high pressures. Several alkanolamines such as MEA (monoethanolamine), DEA (diethanolamine) and MDEA (methylydiethanolamine) have been used, with the selection being determined by the application. The speciality amine aqueous solution strength can vary in the range 15 to 50% and can have a significant effect on the process economics. Generic amines, such as MEA and DEA, are more corrosive and strengths are limited to 30%. The higher the solution strength the liquid circulation requirement is reduced. Steam consumption is highly dependent on this selection, with lower concentrations requiring more steam. The boiling point for regeneration increases at higher MEA concentrations which greatly increases the rate of corrosion of common metals. Also MEA tends to degrade as the temperature rises, increasing the replacement expense and involving the removal of degradation products. MEA has a substantially higher vapor pressure than other amines and a water wash at the top of the absorber can be used to minimize amine losses. The Lean Amine Feed temperature can have a significant influence on the amine losses. A typical gas sweetening flowsheet using amine solution is shown. It consists of an absorber in which cooled lean solvent flows downward contacting with the upward flowing gas to be treated. The rich liquid leaves the absorber at a higher temperature due to the heats of solution and reaction and is preheated with stripper bottoms prior to being fed to the reboiled stripping column. The overhead stripped gas is cooled to remove water vapor which is returned to the column to maintain the water balance. If the gas stream to be treated contains condensable hydrocarbons the lean amine temperature should be above the dew point temperature to prevent condensation of an immiscible hydrocarbon liquid which will promote foaming in the absorber.
24
Section 9
Miscellaneous Application
Miscellaneous Application Simulations
Case/File Name Description
R9.01 Biodiesel Blending
Biodiesel Blending Diesel is blended with Methyl Ester (ME) during ship offloading. The diesel composition in the ester can vary from 0 to 20%v/v and the product blend ester composition is in the range 5% to 15%v/v. The ship discharge flow can vary from 800m3/h to 1000 m3/h at a maximum pressure of 10 barg. The mass balance on the streams give:
Where: DE Diesel in Methyl Ester Flow (m3/h) E Methyl Ester Flow (m3/h) VE Methyl Ester Volume Fraction DS Diesel Flow from Ship (m3/h) VP Bio-diesel Product Volume Fraction
We have: DE
EV
E
E++++
==== and DDE
EV
SE
P++++++++
====
Rearranging gives: (((( ))))
V
V1ED
E
E
E
−−−−==== and substituting for DE leads to the following:
(((( ))))VV1
DVE
EP
SP
−−−−==== and
(((( ))))1VV
1
D
ED
PES
E
−−−−====
++++
Ester Blend to Ship Flow Ratios
Methyl Ester Product Blend VE / VP
Ship to Ester Blend
Flow Ratio % %
100 15 6.667 0.176
10 10.0 0.111
5 20.0 0.053
80 15 5.33 0.231
10 8.0 0.143
5 16.0 0.067
A process control system would set the blender flow ratios by entering the ME blend (VE) and Final Product blend (VP). The ME blend flow (E+DE) required for a “wild” Ship Discharge flow (DS) is calculated. The flow ratio controller would manipulate the ME flow to achieve the desired ratio. In the simulation the blend actual ME component standard liquid volume fraction and the product ME standard liquid volume fraction are set in the appropriate controllers.
DE
E
DS
VE
VP
DE + E
DS + DE + E
25
Case R9.01 Biodiesel Blender Simulation The physical property data used are shown. The ship diesel temperature can vary in the range 15ºC to 40ºC. Thermodynamics used were K- UNIFAC and L- Latent Heat.
Fluid Physical Property Data at 15ºC
Fluid Density Viscosity
kg/m3 cps
Diesel 807.15 2.78
Methyl Ester !00% 876.43 7.48
Methyl Ester 80% 864.74 5.72
The controllers are configured as shown:
The simulation results are in accordance with the theory developed above.
The maximum and minimum ester blend flowrates obtained are 230.7 and 42.1 m3/h.
26
Section 10
General Engineering Data
Contents Units Refinery Process Overview Commercial Steel Pipe ANSI B36.10:1970 & BS 1600 Part 2: 1970 Typical Overall Heat Transfer Coefficients Typical Fouling Resistance Coefficients Heuristics for Process Equipment Design Process Simulation Procedures and Convergence
References 1. Crane Co., “Flow of Fluids Through Valves, Fittings and Pipes”, Publication 410, 1988
Units Volume The basic measurement for crude oil liquid volume is referred to as a barrel (bbl). CHEMCAD unit converter feature by selecting “Fn f6” allows conversion between units as shown:
Gas Constant R 8.314 J/K-mol 1.986 Btu/R-lbmol 0.73 ft3 atm/ R lb mol API gravity formulae API gravity (API ρ) of petroleum liquids is determined from specific gravity (SG) at 60°F:
�� = 141.5�� � 131.5
The specific gravity of petroleum liquids can be derived from the API gravity:
�� = 141.5�� + 131.5
Heavy oil with a specific gravity of 1.0 (density water at 60°F) has an API gravity of:
141.5 � 131.5 = 10°� Crude oil is often measured in metric tons (1000kg). The number of barrels per metric ton for a given crude oil based on its API gravity is calculated from:
�������� = 1 � � 141.5
�� + 131.5 � 0.159! Where 1 bbl = 0.159 m3 A metric ton of West Texas Intermediate 39.6° API would contain about 7.6 barrels.
27
Reid Vapor Pressure (RVP) A measure of gasoline volatility being defined as the absolute vapor pressure exerted by a liquid at 100°F(37.8 °C) as determined by the test method ASTM-D-323. The test method applies to volatile crude oil and volatile non-viscous petroleum liquids. Cetane Index Based on the density and distillation range ASTM D86 of a hydrocarbon using two methods ASTM D976 and D4737 (ISO 4264). Cetane index in some crude oil assays is often referred to as Cetane calcule, while the cetane number is referred to as Cetane measure. Aniline Point Defined as the minimum temperature at which equal volumes of aniline and oil are miscible to give an estimate of the content of aromatic compounds in the oil. The lower the aniline point, the greater is the content of aromatic compounds. VABP and MeABP Petroleum fractions are cuts with specific boiling point ranges, API gravity and viscosity. Each cut can be divided into narrow boiling fractions called pseudo-components where the average boiling point can be estimated as either mid-boiling point or mid-percentage boiling point. The TBP curve is divided into an arbitrary number of pseudo-components or narrow boiling cuts. Since the boiling range is small both mid-points are close to each other and can be considered as the VABP or MeABP for that pseudo-component. Five different average boiling points can be estimated on the distillation curve. The volume average boiling point (VABP) and the mean average boiling points (MeABP) are the most widely used.
VABP is calculated from the ASTM D86 distillation and is the average of the five boiling point temperatures (°F) at 10, 30, 50, 70 and 90% distilled:
"�# = $%& + $'& + $(& + $)& + $*&5
MeABP is calculated from:
+��# = "�# �∆ Where ∆ is given by:
��∆= �0.94402 � 0.008650"�# � 321&.��) + 2.99791�3&.'''
�3 = $*& � $%&90 � 10
MeABP (°R) is used in the definition of the Watson K which is given by:
4 = +��#��
%/'
Factors Note 1 Prefix Symbol
10-12 E-12 pico p
10-9 E-09 nano n
10-6 E-06 micro μ
10-3 E-03 milli m
10-2 E-02 centi c
10-1 E-01 deci d
101 E01 deca da
102 E02 hecto h
103 E03 kilo k
106 E06 mega M Note 2
109 E09 giga G
1012 E12 tera T
Note 1 Tip for setting power, make equal to number 0’s so 0.00001 = 10-5 and 100000 = 105
Note 2 Refinery industry practice sometimes uses MM to signify 106
28
Heuristics for Process Equipment Design
In modelling, “Rules of Thumb” or heuristics based on experience, are used for estimating many parameters before more specific data is available. Piping Design Industry practice for initial design of piping systems is based on economic velocity or allowable pressure drop ∆P/100ft. Once detailed isometrics are available the design will be adjusted to satisfy local site conditions.
Reasonable Velocities for Flow of Fluids through Pipes (Reference Crane 410M)
Service Conditions Fluid Reasonable Velocities Pressure Drop
m/s ft/s kPa / m
Boiler Feed Water 2.4 to 4.6 8 to 15
Pump Suction & Drain Water 1.2 to 2.1 4 to 7
General Service Liquids pumped Non viscous
1.0 to 3.0 3.2 to 10 0.05
Heating Short Lines Saturated Steam 0 to 1.7 bar
20 to 30 65 to 100
Process piping Saturated Steam >1.7 bar
30 to 60 100 to 200
Boiler and turbine leads Superheated Steam 14 and up
30 to 100 100 to 325
Process piping Gases and Vapours 15 to 30 50 to 100 0.02%line pressure
Process piping Liquids gravity flow 0.05
Reasonable velocities based on pipe diameter Process Plant Design, Backhurst Harker p235
Pump suction line for d in (d/6 + 1.3) ft/s d mm (d/500 + 0.4) m/s Pump discharge line for d in (d/3 + 5) ft/s d mm (d/250 + 1.5) m/s Steam or gas d in 20d ft/s d mm 0.24d m/s
Heuristics for process design Reference W.D.Seader, J.D.Seider and D.R.Lewin, “Process Design Principles” are also given:
Liquid Pump suction (1.3 + d/6) ft/s 0.4 psi /100 ft Liquid Pump discharge (5.0 + d/3) ft/s 2.0 psi /100 ft Steam or gas (20d) ft/s 0.5 psi /100 ft
Air for combustion, unless otherwise stated, is at ISO conditions of 15°C, 1.013 bar and 60% relative humidity. Air for compression is defined at Free Air Delivery(FAD) conditions of 20°C, 1 bar and dry.
Pumps
Centrifugal pumps: single stage for 15-5000 gpm, 500 ft max head. Centrifugal pumps: multistage for 20 – 11,000 gpm, 5500 ft max head. Efficiency 45% at 100 gpm, 70% at 100 gpm and 80% at 10,000 gpm.
29
Process Simulation Procedures and Convergence Steady state simulation proves the capability to achieve stable and reproducible operating conditions with acceptable product purity, yield and cycle times to satisfy the commercial requirements and the safety and environmental issues for the regulatory authorities.
A process simulation involves taking the input stream flow rates, compositions and thermodynamic conditions, performing a series of iterative calculations as the streams are processed through Unit Operations and recycles, finally leading to the output stream flow rates, compositions and thermodynamic conditions. The chart below shows the basic steps involved in setting up a steady simulation.
It is recommended that the SCDS UnitOp is used for building fully integrated models because it has a greater number of connection points.
SCDS 1 and SCDS 21-24 icons are the most developed having built in dynamic vessels and control loops. However for our initial exercise we will use SCDS Column 1 icon.
For refinery operations the Tower UnitOp is more suitable as it includes pump around and stripping facilities.
The stage numbering convention in CHEMCAD is from top to bottom, 1 to N. A stage is considered the space above a plate. If a condenser is present it is stage 1; if a reboiler is present it is stage N. To model a column which has ten stages plus condenser and reboiler 12 stages (10+condenser+reboiler = 12) must be specified. If a condenser is present, the feed must not enter stage 1, as that is the reboiler. Top stage feeds should enter stage 2, the top stage (plate), if a condenser is present. Likewise, if a reboiler is present a bottom plate feed is connected to stage (N-1), not stage N. Typically the user has a product specification, mass fraction of a key component in either the bottoms or tops, for a column design or to achieve with an existing column. Converging a column model in simulation is similar to converging a column in the real world; it is difficult to go directly to high purity separation. It is best to start with an easy target, such as reflux ratio and bottoms flowrate. Once the column is converged to this simple specification, we ‘tighten’ the specifications toward the target specification. Use the following procedure:
Feeds
Recycles
Products
30
1. Set up the column: number of stages, condenser, reboiler, operating pressure. 2. Generate TPxy and RCM plots to verify that the target is thermodynamically feasible with
the selected VLE K model. 3. On the SPECIFICATIONS page, set ‘loose’ specifications such as ‘Reflux Ratio’ and ‘ Bottoms Flowrate’ or Reboiler Heat Input. 4. Run the column and converge. Change the specifications if necessary. 5. Go to the CONVERGENCE page of the column dialog. Set the initial flag to 0 Reload
Column Profile. This setting instructs CHEMCAD to use the current converged profile as its starting point (initial conditions) in iterative calculations.
6. On the SPECIFICATIONS page change to more tight specifications. Run the column. If the column converges, tighten the specifications and run again. If the column fails to converge, do not save the profile of the failed attempt. Relax the specifications and run the column again. Repeat from step 5 until you reach the target. Often, it is difficult to obtain the first convergence on a column. If the column is run with no condenser or reboiler, one does not have the option of ‘loose’ specifications. If the column has a condenser or reboiler, relaxing specifications does not always help. 1. On the convergence page of the column dialog, specify estimates if you can make
reasonable estimates. Note that a bad guess will make the column more difficult to converge than no estimate.
2. Remove non key components from the feed(s) to obtain the first convergence. Now set the initial flag to 0 Reload Column Profile, return the other components, and run the unit again.
3. Specify a larger number of iterations on the convergence page of the column dialog. The default is 50, but possibly 52 iterations will find the answer.
4. Try an alternate column model. If you are currently using the SCDS try the same separation with a TOWER or vice versa. The two models use different mathematical models; often one will find an answer in 10 iterations while the other is difficult to converge. It is not possible to obtain different answers with the columns; the models are numerical methods to find a stable composition profile.
5. Consider a partial condenser. If you have a condenser present but have a significant amount of light ends, you may have difficulty converging the column. The default condenser type, total, requires that no vapor leaves stage 1. If light ends are present, this may not be possible without cryogenic temperatures. Changing condenser mode to partial allows the light ends gases to slip past the condenser.