W.L.Nelson, “Petroleum Refinery Engineering”, 4th Edition, McGraw Hill, 1958.
Overview ... Crude petroleum consists of thousands of chemical species.
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
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 Section End
59 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.
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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. 2. 3. 4.
Shrieve, “Chemical Process Industries”, Chapter 37, 5th Edition, McGraw Hill, 1984. J.A.Moulijn, M.Makkee, A.Van Diepen, “Chemical Process Technology”, Wiley, 2001. G.L.Kaes, “Refinery Process Modelling”, Athens Printing Company, 1st Edition, March 2000. 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:
Descriptor -ane -ene -ol -one
Aliphatic or open chain hydrocarbons (hc) Properties Class Formula saturated hc, unreactive paraffins CnH2n+2 unsaturated hc, forms olefines CnH2n additive compounds acetylenes CnH2n-2 reactive, OH replaced alcohols, phenols RCH3OH additive compounds ketones RR1.CO
Member C2H6 ethane C2H4 ethylene C2H2 acetylene C2H5OH ethyl alcohol (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 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 Overhead Bottoms Component 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:
Column Height Column Diameter Number of Trays Weir Length (side) Weir Length (centre) Liquid Flowpath Length Active Area Downcomer Area (side)
Key Tray Specifications Downcomer Area (centre) m 51.8 Tray Spacing m 2.9 Hole Diameter no 74 Total Hole Area m 1.859 Outlet Weir Height m 2.885 Tray Thickness m/pass 0.967 2 Number of Valves m 4.9 m2 0.86 Free Fractional Hole Area
m2 m mm m2 mm mm no/tray %
0.86 0.6 39 0.922 51 2 772 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. Area 0.43 0.335 1.859 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:
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Section 6 Hydrotreater Case/File Name R6.01
Hydrotreater Simulation Description 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.
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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.
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Section 7 Catalytic Reformer Case/File Name R7.01
Reformer Simulation Description 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 Inlet temperature °F Inlet pressure psia Measures ∆T (Typical) °F Recycle MMSCFD Catalyst Volume ft3
Reactor 1 937 413 60 (90-130) 0.5-1% Naphtha Feed 274
Reactor 2 937 394 35 (40-60)
Reactor 3 937 394 25 (1) (10-20)
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 Temperature °F Pressure psia H2 Purity
Separator 100 290 0.79
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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 Number of stages Feed tray Feed temperature °F Tray 1 temperature °F Bottom temperature °F Partial condenser pressure psia
Stabilizer 36 19 297 257 446 (Simulation 488) 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.
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Case 7.01 Catalytic Reformer Flowsheet is shown:
Thermodynamics selected:
K-Peng Robinson
The configuration for catalytic Reactor 3 is shown.
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H-Peng Robinson
Section 8 Amine Treatment Case/File Name R8.01
Vacuum Still Simulations Description 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.
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Section 9 Miscellaneous Application Case/File Name R9.01
Miscellaneous Application Simulations Description 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:
DE E
DE + E
VE
DS
DS + DE + E
VP
Where: DE E VE DS VP
Diesel in Methyl Ester Flow (m3/h) Methyl Ester Flow (m3/h) Methyl Ester Volume Fraction Diesel Flow from Ship (m3/h) Bio-diesel Product Volume Fraction
E E and VP = E + DE E + DE + DS E 1 − VE and substituting for DE leads to the following: DE = VE DE + E 1 V P DS and = E = (V E V P − 1) 1 − VP VE DS VE =
We have:
(
Rearranging gives:
(
Methyl Ester % 100
80
)
)
Ester Blend to Ship Flow Ratios Product Blend VE / VP % 15 10 5 15 10 5
6.667 10.0 20.0 5.33 8.0 16.0
Ship to Ester Blend Flow Ratio 0.176 0.111 0.053 0.231 0.143 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.
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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 Diesel Methyl Ester !00% Methyl Ester 80%
Fluid Physical Property Data at 15ºC Density kg/m3 807.15 876.43 864.74
Viscosity cps 2.78 7.48 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.
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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: 141.5 = 1 0.159! + 131.5 Where 1 bbl = 0.159 m3 A metric ton of West Texas Intermediate 39.6° API would contain about 7.6 barrels. 26
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:
"# = MeABP is calculated from:
$%& + $'& + $(& + $)& + $*& 5
+ # = "# ∆
Where ∆ is given by:
∆ = 0.94402 0.008650"# 321&.) + 2.997913&.''' 3 =
$*& $%& 90 10
MeABP (°R) is used in the definition of the Watson K which is given by:
4=
+ # %/'
Factors Note 1 10-12 10-9 10-6 10-3 10-2 10-1 101 102 103 106 109 1012
Prefix pico nano micro milli centi deci deca hecto kilo mega giga tera
E-12 E-09 E-06 E-03 E-02 E-01 E01 E02 E03 E06 E09 E12
Symbol p n μ m c d da h k M Note 2 G 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
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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) Reasonable Velocities Pressure Drop Service Conditions Fluid 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 Liquids pumped General Service 1.0 to 3.0 3.2 to 10 0.05 Non viscous Saturated Steam Heating Short Lines 20 to 30 65 to 100 0 to 1.7 bar Saturated Steam Process piping 30 to 60 100 to 200 >1.7 bar Superheated Steam Boiler and turbine leads 30 to 100 100 to 325 14 and up 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 Pump discharge line for Steam or gas
d in d in d in
(d/6 + 1.3) ft/s (d/3 + 5) ft/s 20d ft/s
d mm d mm d mm
(d/500 + 0.4) m/s (d/250 + 1.5) m/s 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 Liquid Pump discharge Steam or gas
(1.3 + d/6) ft/s (5.0 + d/3) ft/s (20d) ft/s
0.4 psi /100 ft 2.0 psi /100 ft 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.
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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. Products
Feeds Recycles
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: 29
1. 2. 3. 4. 5.
6.
Set up the column: number of stages, condenser, reboiler, operating pressure. Generate TPxy and RCM plots to verify that the target is thermodynamically feasible with the selected VLE K model. On the SPECIFICATIONS page, set ‘loose’ specifications such as ‘Reflux Ratio’ and ‘ Bottoms Flowrate’ or Reboiler Heat Input. Run the column and converge. Change the specifications if necessary. 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. 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.
2. 3. 4.
5.
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. 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. 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. 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. 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.
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