PRO/II® CASEBOOK Ammonia Synthesis
ABSTRACT Over 140 million tons of ammonia is produced worldwide each year. The rewards for reducing costs, increasing efficiency and improving the profitability of ammonia plants a re enormous. Computer simulation of the plant is the first step towards identifying which parameters control the conversion rate, product purity, energy usage, and the production rate of an existing facility. This casebook demonstrates the use of PRO/II in the simulation of an ammonia synthesis process. The entire plant is modeled, from the reforming of the hydrocarbon feedstream to synthesis gas through its purification to its conversion to ammonia in a synthesis reactor. The ammonia synthesis loop involves a large recycle compared to the feed and product rates. In addition there are several thermal recycles and a nd two control loops. Special thermodynamics are used to ensure the accurate prediction of the separation of ammonia from the other materials. The casebook also outlines o utlines the use of the simulation for parametric studies in the evo lution of a control strategy. Ammonia Synthesis Rev. 3 PRO/II® is a registered mark of Invensys plc. Copyright 2006, all rights reserved
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INTRODUCTION Chemically combined nitrogen is essential for the growth of all living organisms. Animals and plants (with one or two exceptions) are unable to assimilate free nitrogen from the air, so so they depend upon nitrates, ammonium salts or other nitrogen compounds found in the soil. The natural supplies of fixed nitrogen were adequate for many centuries to satisfy the normal th processes of nature. However, by the beginning of the 19 century, the increase in world population along with the growth of big cities created a demand from the more industrialized countries for supplemental sources of fixed nitrogen. This supplement was first found in imported guano and sodium nitrate and later in ammoniacal solutions and ammonium sulphate by-products from the carbonization of coal in gas-works and coke ovens. th
By the start of the 20 century the demand for fertilizer nitrogen again outstripped supply. Three different processes for the fixation of nitrogen were designed and put in commercial operation to ensure adequate supplies of fertilizer nitrogen. One of those three was the direct synthesis of ammonia from nitrogen and hydrogen by the HaberBosch process. This was developed in Germany between 1905 and 1913 and virtually all fixed nitrogen is now produced by this process. The total world production in 2005 was over 140 million tons of fixed nitrogen. Global demand for ammonia is expected to increase about 2% per year for the foreseeable future. More than 90% of this is produced as ammonia and an d about 85% of the total fixed nitrogen production is used in fertilizers.
Alternative Routes to Ammonia There are several licensors of ammonia synthesis processes. All produce ammonia from hydrocarbon feedstocks and air. The hydrocarbon feedstock is usually a natural gas although others, such as naphtha, are used where natural gas is not locally available. The processes are all fundamentally similar: the feed hydrocarbon gas is desulfurized then converted to synthesis gas in a reforming process followed by a CO shift and methanation reactors. The resulting syngas, after purification, is a mixture of hydrogen and nitrogen in stoichiometric quantities (3:1). This is is converted to ammonia in a synthesis reactor which is located in a recycle loop, because of a low conversion per pass. Kellogg Brown & Root (KBR) is the largest licensor of Ammonia Synthesis Technology with over 170 Kellogg plants and over 25 Braun (now owned by KBR) plants. These plants have accounted 1 for more ammonia capacity worldwide than any other licensor.
1
“Ammonia”, by Kellogg Brown & Root, 2000
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The Kellogg Process
At over 170 units worldwide, the Kellogg process has been installed in more ammonia plants than any other process. The simulation considered in this Casebook is a generic Kellogg process. A full flowsheet may be found in Appendix A. Although the Kellogg process is a licensed process, operational benefits can be gained from changing a number of o f the variable parameters: temperatures, recycle rates, air and steam quantities and so on. Recently there has been some work on lowering the steam to carbon ratio to the primary reformer and modifying the catalysts used in the shift converters.The emphasis in ammonia plants today is for lower energy usage. New Kellogg plants claim 6.85Gcal/short ton for all energy requirements over one year.
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PROCESS OVERVIEW The Kellogg Ammonia process is a single train process, divisible into four stages, which operate sequentially. In stage 1 the natural gas undergoes catalytic reforming to produce hydrogen from methane and steam. The nitrogen required for the ammonia is introduced at this stage. In stage 2, the resulting syngas is purified by the removal of carbon monoxide and carbon dioxide in a MEA plant or similar process. Stage 3 consists of compression of the syngas up to the pressure required in stage 4, the ammonia loop.
Feedstocks and products Main Feeds
The main feedstock for this ammonia process is 6 million SCFD of natural gas at a temperature of o 60 F and a pressure of 340 psig. Its composition is shown in Table 1. Table 1. Natural Gas Feed Component Carbon Dioxide Nitrogen Methane Ethane Propane Butane Pentane
Mole % 2.95 3.05 80.75 7.45 3.25 2.31 0.24
Before entering the primary reformer, the natural gas is mixed with superheated steam at 334 psig and 950 F. Nitrogen is supplied from the air which is fed to the secondary reformer at 289 psig and 330 F. The composition of air is displayed in Table 2. Table 2. Air Feed Component Oxygen Nitrogen Argon
Mole % 21.00 78.05 0.95
There is also a water feed which is used to saturate the syngas in the MEA plant. Products
The main product of an ammonia plant is, of course, ammonia. The ammonia product stream must have a purity greater than 99.5%. The major impurities are hydrogen, nitrogen, argon and methane. The other product streams are:
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o letdown gas from the final separator consisting of approximately 45% ammonia, 30% hydrogen with the rest being nitrogen, argon and methane o purge gas from the ammonia loop consisting of approximately 8% ammonia, 57% hydrogen plus nitrogen, argon and methane Ammonia is recovered from the letdown and purge gases in an absorber. Detailed descriptions of the product streams may be found at the end of the Simulation section in Table 3.
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Stage 1: Catalytic Reforming
Following sulfur removal, the primary steam reformer converts about 70% of the hydrocarbon feed into raw synthesis gas in the presence of steam using a nickel catalyst. The main reforming reactions are: CH4 + H2O CO + H2O
↔ ↔
CO + 3H2 CO2 + H2
In the secondary reformer, air is introduced to supply the nitrogen. The heat of combustion of the partially reformed gas raises the temperature and supplies the energy to reform most of the remaining hydrocarbon feed. The reformer product stream is used to generate steam and to preheat the natural gas feed.
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Stage 2: Shift and Methanation
The shift conversion is carried out in two stages. The first uses a high temperature catalyst and the second uses a low temperature one. The shift converters remove the carbon monoxide produced in the reforming stage by converting it to carbon dioxide by the reaction: CO + H2O
↔
CO2 + H2
This reaction also creates additional hydrogen for the ammonia synthesis. Shift reactor effluent is cooled and the condensed water is separated. The gas is then passed to the purification section where carbon dioxide is removed from the synthesis gas in any one of a number of systems such as hot carbonate, MEA, Selexol etc. After the purification stage, the last traces of carbon monoxide and carbon dioxide are removed in the methanation reactions: CO + 3H2O CH4 + H2 CO2 + 4H2 CH4 + 2H2O ↔
↔
The methanation reaction is necessary not only to remove the carbon monoxide and carbon dioxide, but also to create water which can then be removed in the compression stage. This will add efficiency to the process in that any water that is removed will not have to be circulated in the large recycle loop in the ammonia synthesis stage.
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Stage 3: Compression
The purified synthesis gas is cooled and the condensed water is removed. The gas is then compressed in a three stage unit. The centrifugal compressors are driven by steam turbines using steam generated in the plant itself, reducing overall power consumption.
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Stage 4: Conversion
The compressed synthesis gas is dried, mixed with a recycle stream and introduced into the synthesis loop after the recycle compressor. The gas mixture is chilled and liquid ammonia is removed from the secondary separator. The vapor is heated and passed to the ammonia converter. The feed is preheated inside the converter before entering the catalyst bed. The ammonia synthesis reaction is: N2 + 3H2
↔
2NH3
Very high pressures (typically in excess of 300 atmospheres) are required in order to obtain a reasonable conversion. The conversion of hydrogen per pass is still less than 30% and so a large recycle of unreacted gases is necessary. The converter vapor product is cooled by ammonia refrigeration in the primary separator to condense the ammonia product. A purge is removed from the remaining gases to prevent the build up of inerts in the loop. The molar concentration of inerts (argon and methane ) in the converter feed is maintained at 12%.
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ENERGY INTEGRATION The process features a high level of energy integration. Nearly all the power and heating requirements of the process are met by the heat available elsewhere in the process.
In the Reformer The primary reformer passes the natural gas and steam mixture through catalyst-packed tubes in a furnace. The furnace exhaust gases are used to generate steam in a series of boilers. This steam is then used to drive the compressors. The effluent from the secondary reformer is used in a waste heat boiler and then to preheat the feed to the desulfurizer.
In the Shift and Methanators The methanator feed is preheated by exchanging with the methanator product and then further heated by exchange with the high temperature shift reactor effluent. Energy from this effluent is further recovered in a waste heat boiler. The effluent from the low temperature shift reaction is used in another reboiler.
In the Conversion The products from the ammonia converter are mixed with the synthesis gas from the compressors and cooled for ammonia separation by exchange with the converter feed. The converter feed is preheated by the products inside the converter vessel.
MATERIAL RECYCLE The final stage, the Synthesis Loop, is a recycle operation. The conversion of nitrogen and hydrogen to ammonia in the converter is very low - less than 30% of the hydrogen is converted per pass. Therefore there has to be a large recycle in order to convert all the feed and ensure that the final ammonia product reaches the required purity. The ratio of recycle to ammonia product is of the order of 3:1
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PROCESS SIMULATION The PRO/II input for the simulation may be found in Appendix C. This section will explain the details of the PRO/II simulation used in this case study. The screen shots are meant to illustrate specific data entry for points of interest. The “Simulation Flowsheets” shown below differ from the previous process flowsheets in that they include stream identifiers and show the way the simulation is solved. This becomes especially important for recycle solutions, both for thermal recycles and material recycles. A full flowsheet may be found in Appendix A.
General Data Because stage 4 of this process involves a loop which has a large recycle rate in comparison to the product rates, it is essential that the flowsheet is in mass balance. The normal stream component recycle convergence tolerance must therefore be tightened - in this case to 0.05%. The stream Temperature and Pressure recycle tolerances were also tightened to 0.1 and 0.001 respectively. The threshold mole fraction limit for trace components in the recycle is also reduced to 0.0001 in order to ensure that all components are checked for the loop convergence. The maximum allowed number of recycle trials is increased to 150 to ensure that the number of recycle trials is sufficient to converge the flowsheet. The Wegstein acceleration option is used to help speed up convergence. The graphic below shows the data entry window for the recycle convergence and acceleration options.
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PRO/II is also asked to report the overall flowsheet mass balanc e in the Miscelleous Report Options menu.
Component Data All the components in the simulation are available in the PRO/II data bank.
Thermodynamic Data For most of the units in the flowsheet the Soave-Redlich-Kwong (SRK) equation of state is an excellent predictor of phase equilibrium and thermal properties. However, in stage 4 where the ammonia is being separated from the recycle gas in flash units, a more accurate prediction is needed to represent the interaction between ammonia and the other components in the streams. For
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this, a second method is used with user-defined SRK binary interaction parameters defined in Tab le 3 below. Table 3. Binary Interaction Parameters for Thermodynamic Method 2
Component i Component j k ija k ijb k ijc UOM
H2 N2 0.085 0.0 0.0 R
H2 Argon 0.0004 0.0 0.0 R
H2 C1 -0.2079 0.0 0.0 R
N2 C1 0.0204 0.0 0.0 R
H2 NH3 0.276 0.0 0.0 R
N2 NH3 0.31 0.0 0.0 R
Argon C1 NH3 NH3 0.3383 0.18 0.0 0.0 0.0 0.0 R R
Invoking the use of the second thermodynamic set for the ammonia separators is as simple a s selecting the second thermodynamic set from the Thermodynamic System drop-down window in the Flash drum data entry window. The graphic below depict this drop down list with the available thermodynamic methods.
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36
34
39
37 D-7 DM1
37A 40
32A
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Stream Data Feed Streams
The main feed streams - natural gas, steam, and air - are defined in the normal way by rate, composition, temperature, and pressure. In addition there is a water feed to the MEA plant which ensures that the syngas leaves the plant saturated. Recycle Streams
There is a recycle stream in stage 4. The large flowrate of this recycle stream relative to the flowsheet feeds and products along with the fact that the ammonia species is created in the recycle, necessitates an initial estimate of the rate and composition of the recycle stream. The stream initial estimates are entered into stream 31R below:
Other Streams
There are several thermal recycles in this flowsheet. These may be simulated in one of several different ways: They can be input as they are and allowed to converge naturally They can be replaced by simpler units and solved rigorously after the flowsheet has solved Or they may be circumvented by using the REFERENCE STREAM facility.
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The reference stream option was chosen in this simulation to circumvent the thermal recycles. Thermal recycles can be broken using reference streams because the composition of the stream remains the same through the thermal recycle loop. The only thing that changes within the thermal recycle loop is temperature and pressure. If the temperature an pressure is know elsewhere in the loop, a reference stream can be used to break the thermal recycle loop. This will speed up the convergence time of the simulation without sacrificing accu racy in the results. In stage 2, the feed to the low temperature shift reactor (stream 13) is referenced to the h igh temperature shift product (stream 11). This defines all the properties of stream 13 as the same as o stream 11. Temperature and pressure are over-written and defined as 400 F and 274 psig. Similarly, o stream 21 is referenced to stream 19 with a temperature of 675 F and a pressure of 254 psig. These reference streams allow the heat exchangers to be solved separately after converging the air flowrate controller loop.
Stream 13 is not defined as a product from a unit operation. However, in stage 4, stream 38B is a unit operation product. Nevertheless, in order to eliminate a thermal recycle around the Ammonia Converter, it is referenced to stream 38. This means that stream 38B takes its composition and rate at all times from stream 38 but its temperature and pressure are calculated by the heat exchanger from which it is a product.
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Unit Operations for Stage 1 - Catalytic Reforming
Input
Since there is no sulfur in the input, the Desulfurizer RX1 is modeled as an isothermal flash to set the temperature and pressure of the reformer feed. Both the reformers, RX2 and RX3, are modeled as Gibbs Free Energy reactors with typical temperature approaches specified. The rate of air feed to the secondary reformer must be such that the syngas product from stage 2 has the correct nitrogen:hydrogen ratio. This is achieved by putting a controller on the stage 2 Methanator product and varying the air feed rate to the secondary reformer. This means that stages 1 and 2 are in a controller loop. To satisfy the primary reformer, the molar ratio of steam (stream 5) to natural gas (stream 1) should be 6:1. The flow rate of the steam could be set to the correct value in the Stream Data Section. However, that would involve calculating the molar rate of natural gas (the stream is known only in volume units) before the simulation begins. It is much e asier to let PRO/II perform that calculation. Furthermore, allowing PRO/II to perform this calculation allows the natural gas stream to be PRO/II CASEBOOK
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changed at any time (for a turndown case or a different design case) and the steam rate to be automatically recalculated.
The secondary reformer product (stream 8) is used to preheat the desulfurizer feed (stream 2). This thermal recycle is not a simple one in that the process demands a fixed temperature for both the products from exchanger X-1. This is achieved by dividing the exchanger into two halves: the cold side is X-1A and is solved before the desulfurizer; the secondary reformed product is cooled in H-1, a combination of waste heat boiler WHB1 and the hot side of X-1. After stages 1 and 2 have solved, H-1 is divided into WHB1 and the hot side of X-1 (X-1B). This can be achieved because PRO/II allows the use of reference streams to link two streams. The define feature is similar to reference stream feature except that it links unit o peration parameters. The define feature was used to link the duties from one exchanger to another. In this case stream 8_R1 was referenced to stream 8 and the duties for X1-B and WHB1 were defined as the duties for X-1A and H-1 respectively. The graphics below illustrate the use of the define feature to line the duties together.
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WHB1's duty is calculated as the difference between the duties of H-1 and X-1A. X1-B's duty is simply the same as that for X1-A except with the opposite sign. Results
The stream calculator determines the steam flowrate as 3952. 7 lb mole/hr which is six times the natural gas flowrate. Virtually all of the C2 and higher hydrocarbons are broken down in the primary reformer and converted to hydrogen, carbon monoxide and carbon dioxide. 58% of the methane in the feed is also converted. 95% of the residual methane is converted in the secondary reformer along with the remaining traces of ethane and propane. All of the oxygen is consumed in the secondary reformer. The duties of WHB1 and X-1 are 53.62 and 4.17 MM Btu/hr respectively.
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Unit Operations for Stage 2 - Catalytic Shift and Methanation
Input
Both the High Temperature Shift Reactor (RX4) and the Low Temperature Shift Reactor (RX5) are modeled by an equillibrium reactor unit operation using the built-in shift reaction equilibrium data. They operate adiabatically and the products are assumed to be at equilibrium. The MEA plant removes the bulk (99.92%) of the carbon dioxide from the shift reactor product and is modeled as a stream calculator (T-1). The treated gas is saturated with water in flash unit SAT and excess water discharged. The final part of stage 2 is the Methanator (RX6), modeled by an equillibrium reactor unit operation using the built-in methanator and shift reaction equilibrium data. The reactor operates adiabatically and equilibrium is assumed to be achieved. The syngas must have a hydrogen:nitrogen molar ratio of 3:1 to satisfy the stoichiometry of the ammonia reaction. This is accomplished by inserting a controller to measure that ratio in the Methanator product and vary the air feed to the Secondary Reformer until the required ratio is achieved. Thus there is a significant loop involving most of the units of stages 1 and 2. This makes it more important for any thermal recycles within that loop to be eliminated if at all possible. The thermal recycle in stage 2 of the plant is a complex one, involving preheating the Methanator feed first with its own product and then with the RX4 product. This recycle is eliminated by referencing streams 13 and 21 to bypass these exchangers completely. PRO/II CASEBOOK
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Stream 13 is the same as stream 11 except for its temperature and pressure, both of which are known. Therefore stream 13 is set in the Stream Data Section and referenced to stream 11. It gets its rate and composition from stream 11 and its conditions from the Stream Data Section. In the same way, stream 21 is referenced to stream 19. After the control loop has solved, exchangers X-2, WHB2 and X-3 are solved in the normal way. Results
The controller solves after 3 trials with an air flowrate of 1008 lb mole/hr. This produces the required hydrogen:nitrogen ratio in the product from the methanator. 85% of the carbon monoxide is converted in RX-4 and 93% of the remainder is converted in RX-5. This gives a concentration of 0.055% carbon monoxide and 11.1% carbon dioxide in the exit gas from RX-5. The concentration of carbon dioxide is reduced to 0.02% in the MEA Plant. The final removal of all the carbon monoxide and carbon dioxide is carried out in the methanator reactor.
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Unit Operations for Stage 3- Compression
Input
The Condensate Separator (D-2) is modeled as a flash unit, decanting water in a declared water stream and using the PRO/II built-in water solubility data. The compressors are modeled as single stage isentropic compressions with a specified adiabatic efficiency of 95%. Each one has an aftercooler reducing the product temperature to 95 F. The first and second compressors also have water decantation streams to knock out any liquid water that may have condensed in the aftercoolers. Finally, all the remaining water is removed in a drying unit (SEP1), modeled as a stream calculator. Results
The compressor work for the three stages is 1752, 1831 and 1776 HP respectively. The corresponding aftercooler duties are 4.73, 4.70 and 4.35 MM Btu/hr. Most of the water is removed after the first compressor with a small amount removed after the second compressor. The final separator removes all remaining water from the synthesis loop feed.
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Unit Operations for Stage 4 – Conversion
Input
The fourth compressor (CP-4) is in the ammonia recycle loop. The loop starts at that unit, stream 31R being the main recycle stream. An initial estimate for this stream is needed, as discussed in the Stream Data Section above. The Ammonia Converter feed is preheated by exchanging inside the converter with its product stream. This exchanger is modeled separately in the flowsheet as FDEF. The exchangers before the separators do not appear on the simulation flowsheet as they are combined with the separators. There are two thermal recycles in the loop. The loop involving the convertor and feed preheater is split by referencing the convertor feed (stream 38B) to the product from exchanger X-4 (stream 38) in the Stream Data Section. FDEF is then solved after the convertor when both feed streams have been calculated. The loop involving exchanger X-4 is solved explicitly as it appears in the flowsheet. This involves introducing an inner loop within the ammonia recycle loop. This loop could also be eliminated by using the devices mentioned above. The Ammonia Converter (RX-7) is modeled using an equilibrium reactor model with supplied equilibrium data.
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The reference condition for heat of reaction data is given as vapor at 800 F and the stoichiometry represents the well-known ammonia synthesis reaction. The h eat of reaction is given as -45.18 thousand energy units per mole of base component reacted, in this case nitrogen. The equilibrium constant is a function of temperature according to the Arrhenius equation: Ln (K eq) = A + B/T The equilibrium constant is computed at a temperature equal to the reaction temperature minus the approach temperature difference given. The feed to the reactor has to contain 12 mole% of inerts (argon and methane). This is achieved by varying the rate of the recycle stream 31R using a controller (CTL1). This action of the controller could conflict with the recycle convergence because this stream is also the recycle stream for the main loop. This is avoided by solving the recycle before the controller is invoked. In order to maintain the inerts material balance, a purge stream is taken off the recycle stream. The rate of this stream is calculated in calculator (BD-1). This would act similar to a valve on the purge stream regulating the flow.
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The calculator builds a dummy stream 31X consisting of stream 28, the dry syngas feed to the loop, minus the products from the loop, streams 32 and 37. The splitter SP1 operates such that the argon rate in the purge stream 31P is equal to the argon rate in the dummy stream 31X - in other words equal to the net argon coming in to the loop. Thus there is an exact argon balance in the loop and there will be no build up of inerts as the flowsheet recycle solves.
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Results
The synthesis loop solves with a production of 1539.6 lb mole/hr of 99.7% purity ammonia. The product compositions are shown in Table 4. The overall conversion to ammonia is 98% with a reactor conversion per pass of 27.4% based on Nitrogen. The recycle stream rate is 10654 lb mole/hr giving a recycle:feed ratio of 3.3:1. The purge stream is 74.1 lbmol/hr which is 0.7 % of the total recycle flowrate. Table 4. Ammonia Plant Product Streams Component Hydrogen Nitrogen Argon Methane Ammonia Rate (lb mol/hr) Temperature (F) Pressure (psig)
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NH3 Product 1.69 0.43 0.56 1.92 1535.03 1539.72 79.4 350
Let Down Gas 21.45 6.15 4.05 7.68 33.3 72.65 79.4 350
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Purge 42.39 14.38 4.97 6.11 6.27 74.12 85 4660
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USE OF THE MODEL IN PROCESS CONTROL Very clear benefits of using a computer model of a Kellogg Ammonia plant were demonstrated in a paper published in Hydrocarbon Processing, November 1980. The paper focuses on the design of control strategies and uses a simulation model to derive the responses of the plant to different settings of various parameters. The parameters that were examined were: The ratio of synthesis loop recycle rate to fresh syngas In many plants the fresh syngas feed to the synthesis loop is limited by front-end restrictions such as insufficient synthesis compressor power or absorber capacity. The ratio of nitrogen to hydrogen in the syngas At constant syngas make-up rates, ammonia production increased as H/N decreased, but synthesis compressor loadings increased. Synthesis loop pressure Production of ammonia increases with increased loop pressure. The design pressure, constrained by mechanical considerations, is the limiting factor and the purge system should be operated the keep the operating pressure just below the design pressure. The temperature of the synthesizer feed Production is extremely sensitive to this parameter and an optimum temperature can be found. The control of this parameter is critical to the profitability of the plant. Methane content in syngas Ammonia production decreases with increasing methane. Improved primary reformer temperature control can alleviate this problem. The results of these simulations have been put into practice in a number of installations and these have consistently led to improved productivity and higher conversion efficiency.
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APPENDIX A - Complete Ammonia Plant Flowsheets Graphical Representation:
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PRO/II Flowsheet:
PG001
CO2
PG000
17
T-1
7
19A
5 6
4
WAT
16
13
SAT
15
S1
3
RX-2
RX-4
RX-3
XS
D-1
18
RX-5
10
8
RX-1
REB1
13X
11 21X
X-1A
1
CW-1
14
22
WHB2
2 H-1
21
GASHEATER
12 8_R1 9
CT1
10X
X-2 RX-6
WHB1
X1-B
20
19
Stage 1 - Catalytic Reforming
X-3
23
CW2
24
Stage 2 - Catalytic Shift and Methanation
PG002
PG003
38B WA3 27
26
CP-2
CP-1
WA1
WA2
31X
29A 28A
BD1 FDEF
X-4
CP-3
RX-7
38 SEP1
36A
25 28
REFC
S1 34
31P
DM2
29
M1
36
SP1
D-2
24W
31A
CTL1
31RA
33
31
Stage 3 - Compression
CAL1
39
37
31RB 31R
DM3 DUM1
D-7
CP-4
DM1
37A
DUM2
40 32 32A
D-6
D-8
DM4
Stage 4 - Conversion
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7 5 6
4 S1
3
RX-2
RX-3
10
8
RX-1
X-1A
1
2 GASHEATER
Hx Name Hx Description Duty
WHB1
X1-B
53.6202
4.1698
H-1 MM BTU/HR
8_R1 9 WHB1
10X X1-B
Stage 1 - Catalytic Reforming
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CO2
17
T-1
19A 10
WAT
16
13
SAT
15 CW-1
14 RX-4
XS
D-1
18
RX-5
REB1
13X
11 21X
22
WHB2
21 12 CT1 X-2 RX-6
20
19 X-3
23
CW2
24
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24 WA3 27
26
CP-2
CP-1
WA1
WA2
28A CP-3
SEP1
25 28
D-2
24W
Stage 3 - Compression
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Page 36
38B
31X
29A BD1 FDEF
X-4 RX-7
38 36A 28
REFC
S1
CAL1
34
31P
DM2
29
M1
36
SP1
31A
CTL1
31RA
33
31
39
37
31RB 31R
DM3 DUM1
D-7
CP-4
DM1
37A
DUM2
40 32 32A
D-6
D-8
DM4
Stream Name Stream Description Temperature Pressure Phase Flowrate
40 NH3 PROD F PSIG LB-MOL/HR
79.412 350.000 Liquid 1539.771
39 LETDWN GAS 79.412 350.000 Vapor 72.685
31P PURGE 84.999 4660.000 Vapor 73.284
Stage 4 - Conversion
PRO/II CASEBOOK
AMMONIA SYNTHESIS
Page 37
APPENDIX B - PRO/II SELECTED OUTPUT The chart below was taken from the PRO/II output file. This depicts the overall plant mass balanc e as calculated by PRO/II: OVERALL PLANT MASS BALANCE COMPONENT --------------------1 H2O 2 O2 3 CO 4 CO2 5 H2 6 N2 7 A 8 C1 9 C2 10 C3 11 NC4 12 NC5 13 NH3 TOTAL
- - - - - - - - - - - - - - - - - - LB/ HR - - - - - - - - - - - - - - - - - FEED +REACTI ON - PRODUCT =DEVI ATI ON ---------- ---------- ---------- ---------200171. 24 - 19308. 08 180863. 16 0. 00 6773. 16 - 6773. 16 0. 00 0. 00 11110. 16 0. 00 11110. 16 0. 00 51686. 75 32899. 52 84586. 27 0. 00 23765. 16 124. 05 23896. 18 - 6. 96 135607. 21 - 22050. 76 113588. 40 - 31. 95 2295. 14 0. 00 2295. 14 0. 00 9501. 45 - 8284. 17 1219. 00 - 1. 71 1475. 82 - 1475. 81 0. 01 0. 00 944. 14 - 944. 14 0. 00 0. 00 884. 53 - 884. 53 0. 00 0. 00 114. 08 - 114. 08 0. 00 0. 00 0. 00 26811. 15 26817. 75 - 6. 60 444328. 83
0. 00
444376. 06
- 47. 23
PERCENT DEV ------0. 00 0. 00 0. 00 0. 00 - 0. 03 - 0. 03 0. 00 - 0. 14 0. 00 0. 00 0. 00 0. 00 - 0. 02 - 0. 01
The chart below was taken from the PRO/II output file. This depicts the summary of RX-7,7, the ammonia converter. OPERATI NG CONDI TI ONS REACTOR TYPE DUTY, MM BTU/ HR TOTAL HEAT OF REACTI ON AT 800. 00 F, MM BTU/ HR
FEED VAPOR PRODUCT TEMPERATURE, F PRESSURE, PSI G
I NLET ----------38B 571. 29 4760. 0000
ADI ABATI C REACTOR 7. 08086E- 05 - 35. 5038 OUTLET ----------29A 926. 82 4730. 0000
REACTI ON DATA COMPONENT -----------------2 O2 3 CO 4 CO2 5 H2 6 N2 7 A 8 C1 9 C2 10 C3 11 NC4
- - - - - - - - - RATES, L B- MOL / HR - - - - - - - - - FRACTI ON FEED CHANGE PRODUCT CONVERTED ----------- ----------- ----------- ----------6. 14596E- 09 0. 0000 6. 14596E- 09 1. 04160E- 05 0. 0000 1. 04160E- 05 9. 67314E- 08 0. 0000 9. 67314E- 08 8516. 0375 - 2361. 2530 6154. 7844 0. 2773 2875. 2775 - 787. 0843 2088. 1931 0. 2737 723. 2566 0. 0000 723. 2566 892. 0648 0. 0000 892. 0648 1. 23713E- 03 0. 0000 1. 23713E- 03 1. 21410E- 08 0. 0000 1. 21410E- 08 4. 63078E- 09 0. 0000 4. 63078E- 09
PRO/II CASEBOOK
AMMONIA SYNTHESIS
Page 38
12 NC5 13 NH3 TOTAL BASE COMPONENT -----------------6 N2
3. 02841E- 09 525. 0477
0. 0000 1574. 1687
3. 02841E- 09 2099. 2164
13531. 6852
- 1574. 1687
11957. 5165
REACTI ON ----------1
LB- MOL/ HR CONVERTED ----------787. 0843
FRACTI ON CONVERTED( 1) -----------0. 2737
REACTOR MASS BALANCE Component -------------------2 O2 3 CO 4 CO2 5 H2 6 N2 7 A 8 C1 9 C2 10 C3 11 NC4 12 NC5 13 NH3 Tot al
PRO/II CASEBOOK
- - - - - - - - - - - Rat es, LB/ HR - - - - - - - - - - - Fr act i on Feed Change Pr oduct Conver t ed ----------- ----------- ----------- ----------1. 96663E- 07 0. 0000 1. 96663E- 07 2. 91756E- 04 0. 0000 2. 91756E- 04 4. 25713E- 06 0. 0000 4. 25713E- 06 17167. 3105 - 4760. 0030 12407. 3075 0. 2773 80546. 5254 - 22048. 9708 58497. 5546 0. 2737 28892. 6554 0. 0000 28892. 6554 14311. 1818 0. 0000 14311. 1818 0. 0372 0. 0000 0. 0372 5. 35374E- 07 0. 0000 5. 35374E- 07 2. 69157E- 07 0. 0000 2. 69157E- 07 2. 18500E- 07 0. 0000 2. 18500E- 07 8941. 8560 26808. 9734 35750. 8294 149859. 5666
0. 0000
AMMONIA SYNTHESIS
149859. 5662
Page 39
The following pages show selected parts of the stream information exported to Excel. Component Rates Str eam
Name
NATGAS
Description
NAT GAS STEAM Vapor Vapor
Phase
Temperature Pressure Molecular Weight
F PSIG
60.00 340.0 20.30
Component Molar Rates H2O O2
LB-MOL/HR
5
950.00 334.0 18.02
7 AIR Vapor
PURGE LET DWN GAS NH3 PROD Vapor Vapor Liquid
85.00 4660.0 12.03
79.42 350.0 14.70
79.42 350.0 17.02
211.67
0.00
0.00
0.00
786.70 9.58
0.00 0.00 42.39 14.39 4.97
0.00 0.00 21.45 6.15 4.05
0.00 0.00 1.69 0.44 0.56
20.09
6.11 0.00 0.00 0.00 0.00
7.68 0.00 0.00 0.00 0.00
1.92 0.00 0.00 0.00 0.00
6.27 74.12
33.28 72.61
1535.03 1539.64
0.00 0.00 0.00 0.57 0.19 0.07 0.08 0.00 0.00 0.00 0.00 0.08
0.00 0.00 0.00 0.30 0.08 0.06 0.11 0.00 0.00 0.00 0.00 0.46
0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 1.00
531.97 49.08 21.41 15.22 1.58
NH3
PRO/II CASEBOOK
40
330.00 289.0 28.96
19.43
C1 C2 C3 NC4 NC5
Component Mole Fractions H2O O2 CO CO2 H2 N2 A C1 C2 C3 NC4 NC5 NH3
39
3952.74
CO CO2 H2 N2 A
Total
31P
LB-MOL/HR
658.79
3952.74 1007.95 1.00 0.21
0.03 0.03
0.78 0.01
0.81 0.07 0.03 0.02 0.00
AMMONIA SYNTHESIS
Page 40
Str eam
Name
NATGAS
5
Description
NAT GAS Vapor
Phase
7
31P
39
40
STEAM
AIR
PURGE
Vapor
Vapor
Vapor
Vapor
Liquid
658.79
3952.74
1007.95
74.12
72.61
1539.64
13371.05 607.71
71209.75 1141.79
29194.01 536.33
891.78 37.68
1067.09 36.41
26211.19 682.83
LET DWN GAS NH3 PROD
Total Stream Pr operti es
Rate
LB-MOL/HR
Std. Liquid Rate
LB/HR FT3/HR
Temperature
F
60.00
9 50.00
330.00
85.00
79.42
79.42
Pressure
PSIG
340.00
334.00
289.00
4660.00
350.00
350.00
Molecular Weight Enthalpy
MM BTU/HR
20.30 1.34
18.02 106.60
28.96 1.83
12.03 0.07
14.70 0.34
17.02 1.54
BTU/LB
100.14
1497.06
62.55
76.90
319.43
58.89
N/A 1.36
N/A 1.21
N/A 3.31
N/A 2.92
N/A 1.27
1.00 0.74
0.53
0.11
0.56
10.49
0.38
0.22
0.03 17.48
0.34 8.76
0.03 6.00
-0.09 11.88
0.06 12.29
0.25 12.26
22.00
62.37
54.43
23.67
29.31
38.39
0.35 269.59
1.00 10.00
0.87 30.62
0.38 241.36
0.47 169.61
0.62 98.40
658.79
3952.74
1007.95
74.12
72.61
n/a
LB/HR
13371.05
71209.75
29194.01
891.78
1067.09
n/a
FT3/HR
9627.93
168399.02
28355.77
105.55
1087.83
n/a
Mole Fraction Liquid Reduced Temp. Pres. Acentric Factor Watson K (UOPK) Standard Liquid Density
LB/FT3
Specific Gravity API Gravity Vapor Ph ase Properti es
Rate
Std. Vapor Rate
LB-MOL/HR
FT3/HR
250000.00 1500000.13 382500.00 28127.78
Specific Gravity (Air=1.0) Molecular Weight
27552.61
n/a
0.70 20.30
0.62 18.02
1.00 28.96
0.42 12.03
0.51 14.70
n/a n/a
Enthalpy
BTU/LB
100.14
1497.06
62.55
76.90
319.43
n/a
CP
BTU/LB-F
0.51
0.53
0.25
0.68
0.57
n/a
Density Thermal Conductivity Viscosity
LB/FT3 BTU/HR-FT-F CP
1.39 n/a n/a
0.42 0.04 0.03
1.03 n/a n/a
8.45 n/a n/a
0.98 n/a n/a
n/a n/a n/a
LB-MOL/HR LB/HR
n/a n/a
n/a n/a
n/a n/a
n/a n/a
n/a n/a
1539.64 26211.19
FT3/HR
n/a
n/a
n/a
n/a
n/a
701.00
FT3/HR
n/a
n/a
n/a
n/a
n/a
682.83
n/a n/a
n/a n/a
n/a n/a
n/a n/a
n/a n/a
0.62 17.02
Li quid Phase Properti es
Rate
Std. Liquid Rate Specific Gravity (H2O @ 60 F) Molecular Weight Enthalpy
BTU/LB
n/a
n/a
n/a
n/a
n/a
58.89
CP Density
BTU/LB-F LB/FT3
n/a n/a
n/a n/a
n/a n/a
n/a n/a
n/a n/a
1.25 37.39
Surface Tension
DYNE/CM
n/a
n/a
n/a
n/a
n/a
n/a
Thermal Conductivity Viscosity
BTU/HR-FT-F CP
n/a n/a
n/a n/a
n/a n/a
n/a n/a
n/a n/a
n/a n/a
PRO/II CASEBOOK
AMMONIA SYNTHESIS
Page 41
APPENDIX C - PRO/II INPUT FILE TI TLE DATE=NH3 SYN DESC Thi s i s t he PRO/ I I si mul ati on f i l e cor r espondi ng t o t he DESC PRO/ I I Casebook on Ammoni a Synt hesi s. DESC For mor e i nf or mat i on on t hi s model DESC pl ease contact Techni cal Suppor t DESC PRI NT I NPUT=ALL, STREAM=COMPONENT, RATE=M, MBALANCE, I ON=NONE TOLERANCE STREAM=0. 0005, - 0. 1, 0. 0001, 0. 001 DI MENSI ON ENGLI SH, PRES=PSI G SEQUENCE DEFI NED=GASHEATER, X- 1A, RX- 1, S1, RX- 2, RX- 3, H- 1, RX- 4, RX- 5, & REB1, CW- 1, D- 1, T- 1, SAT, RX- 6, CT1, X- 3, X- 2, WHB2, WHB1, X1- B, CW2, & D- 2, CP- 1, CP- 2, CP- 3, SEP1, REFC, CP- 4, X- 4, M1, D- 7, DM2, DM1, RX- 7, & FDEF, D- 6, DM3, DM4, BD1, SP1, DUM1, CAL1, DUM2, CTL1, D- 8 CALCULATI ON TRI ALS=150, RECYCLE=ALL, COMPCHECK=CALC, DVARI ABLE=ON, & FL ASH=DEFAULT, MAXOPS=1000000, CDATA=FI X COMPONENT DATA LI BI D 1, H2O/ 2, O2/ 3, CO/ 4, CO2/ 5, H2/ 6, N2/ 7, A/ 8, C1/ 9, C2/ 10, C3/ 11, NC4/ & 12, NC5/ 13, NH3, BANK=SI MSCI , PROCESS THERMODYNAMI C DATA METHOD SYSTEM=SRK, TRANSPORT=NONE, DENSI TY( L) =RCK1, SET=1, DEFAULT METHOD SYSTEM=SRK, DENSI TY( L) =RCK1, SET=2 KVAL( VLE) SRK( R) 5, 6, 0. 085, 0, 0 SRK( R) 5, 7, 0. 0004, 0, 0 SRK( R) 5, 8, - 0. 2079, 0, 0 SRK( R) 7, 8, 0. 0204, 0, 0 SRK( R) 5, 13, 0. 276, 0, 0 SRK( R) 6, 13, 0. 31, 0, 0 SRK( R) 7, 13, 0. 3383, 0, 0 SRK( R) 8, 13, 0. 18, 0, 0 STREAM DATA PROPERTY STREAM=1, TEMPERATURE=60, PRESSURE=340, PHASE=M, & RATE( GV) =250000, COMPOSI TI ON( M) =4, 2. 95/ 6, 3. 05/ 8, 80. 75/ & 9, 7. 45/ 10, 3. 25/ 11, 2. 31/ 12, 0. 24 PROPERTY STREAM=5, TEMPERATURE=950, PRESSURE=334, PHASE=M, & COMPOSI TI ON( M, LBM/ H) =1, 1000 PROPERTY STREAM=7, TEMPERATURE=330, PRESSURE=289, PHASE=M, & RATE( GV) =375000, COMPOSI TI ON( M) =2, 21/ 6, 78. 05/ 7, 0. 95 PROPERTY STREAM=WAT, TEMPERATURE=100, PRESSURE=271, PHASE=M, & COMPOSI TI ON( M, LBM/ H) =1, 1000 PROPERTY STREAM=31R, TEMPERATURE=85, PRESSURE=4660, PHASE=M, & COMPOSI TI ON( M, LBM/ H) =5, 6000/ 6, 2000/ 7, 300/ 8, 1000/ 13, 700 PROPERTY STREAM=31X, TEMPERATURE=85, PRESSURE=4660, PHASE=M, & COMPOSI TI ON( M, LBM/ H) =7, 7 PROPERTY STREAM=36A, TEMPERATURE=40, PRESSURE=4840, PHASE=M, & COMPOSI TI ON( M, LBM/ H) =5, 8514. 13/ 7, 724. 309/ 2, 6. 16216E- 9/ & 3, 1. 04288E- 5/ 4, 9. 6746E- 8/ 6, 2874. 7/ 8, 893. 043/ 9, 0. 00123812/ & 10, 1. 21531E- 8/ 11, 4. 64927E- 9/ 12, 3. 04016E- 9/ 13, 525. 107 PROPERTY STREAM=S1, TEMPERATURE=93. 719, PRESSURE=4950, PHASE=M, & COMPOSI TI ON( M, LBM/ H) =5, 8517. 75/ 7, 724. 917/ 2, 6. 18542E- 9/ & 3, 1. 04418E- 5/ 4, 1. 01574E- 7/ 6, 2875. 69/ 8, 894. 384/ 9, 0. 00124165/ & 10, 1. 21867E- 8/ 11, 4. 65563E- 9/ 12, 3. 04603E- 9/ 13, 902. 201 PROPERTY STREAM=13, TEMPERATURE=400, PRESSURE=274, REFSTREAM=11 PROPERTY STREAM=21, TEMPERATURE=675, PRESSURE=254, REFSTREAM=19 PROPERTY STREAM=38B, TEMPERATURE=571. 14, PRESSURE=4760, & REFSTREAM=38 PROPERTY STREAM=8_ R1, REFSTREAM=8 NAME 1, NAT GAS/ 5, STEAM/ 7, AI R/ 31R, RECYCLE/ 11, HTS OUT/ 19, TRTD GAS/ & 38, RX FEED/ 8, SEC REF OUT/ 6, PR REF OUT/ 14, LTS OUT NAME 17, MEA FEED/ 22, METH PROD/ 25, SYN GAS/ 31P, PURGE/ 31, PRI SEP/ & 36, SEC SEP/ 39, LETDWN GAS/ 40, NH3 PROD RXDATA RXSET I D=1 REACTI ON I D=1 STOI CHI OMETRY 5, - 3/ 6, - 1/ 13, 2 HORX HEAT=- 45. 108, REFCOMP=6, REFTEMP=800 EQUI LI BRI UM A=- 32. 975, B=22930. 4 UNI T OPERATI ONS
PRO/II CASEBOOK
AMMONIA SYNTHESIS
Page 42
HX
UI D=GASHEATER COLD FEED=1, M=2 OPER DUTY=1. 6 HX UI D=X- 1A COLD FEED=2, V=3, DP=2 OPER CTEMP=750 FL ASH UI D=RX- 1, NAME=DESULFURI ZER FEED 3 PRODUCT V=4 I SO TEMPERATURE=740, DP=2 CALCULATOR UI D=S1, NAME=STM- GAS FL OW SEQUENCE STREAM=1, 5 PROCEDURE R( 1) = 6. 0 * SMR( 1) $ CALCULATE STEAM RATE CALL SRXSTR( SMR, R( 1) , 5) $ SET STEAM RATE TO CALCULATED RATE RETURN GI BBS UI D=RX- 2, NAME=PRI REFORMER FEED 4, 5 PRODUCT V=6 OPERATI ON DP=45, TEMPERATURE=1360, I SOTHERMAL ELEMENTS REACTANTS= 1/ 3/ 4/ 5/ 8/ 9/ 10/ 11/ 12 CONVERSI ON APPROACH=- 35 GI BBS UI D=RX- 3, NAME=SEC REFORMER FEED 6, 7 PRODUCT V=8 OPERATI ON DP=5, ADI ABATI C ELEMENTS REACTANTS= 1/ 2/ 3/ 4/ 5/ 8/ 9/ 10/ 11/ 12 CONVERSI ON APPROACH=- 35 HX UI D=H- 1, NAME=COOL REFGAS HOT FEED=8, M=10, DP=4 OPER HTEMP=675 EQUREACTOR UI D=RX- 4, NAME=H T SHI FT FEED 10 PRODUCT V=11 OPERATI ON ADI ABATI C, DP=2 RXCALCULATI ON MODEL=SHI FT REACTI ON SHI FT EQUREACTOR UI D=RX- 5, NAME=L T SHI FT FEED 13 PRODUCT V=14 OPERATI ON ADI ABATI C, DP=2 RXCALCULATI ON MODEL=SHI FT REACTI ON SHI FT FL ASH UI D=REB1, NAME=REBOI LER FEED 14 PRODUCT V=15 ADI ABATI C DP=4, DUTY=- 45 HX UI D=CW- 1 HOT FEED=15, V=16, DP=2 OPER HTEMP=100 FLASH UI D=D- 1, NAME=COND SEP FEED 16 PRODUCT V=17, W=18 ADI ABATI C STCALCULATOR UI D=T- 1, NAME=MEA COLUMN FEED 17, 1 OVHD V=CO2, DTAD=0 BTMS L=19A, DP=12, TEMPERATURE=100 FOVHD( M) 1, 3, 0 FOVHD( M) 4, 4, 0. 9992 FOVHD( M) 5, 13, 0 OPERATI ON STOP=ZERO FLASH UI D=SAT, NAME=H2O SAT FEED 19A, WAT PRODUCT V=19, W=XS I SO TEMPERATURE=100 EQUREACTOR UI D=RX- 6, NAME=METHANATOR FEED 21 PRODUCT V=22 OPERATI ON ADI ABATI C, DP=2 RXCALCULATI ON MODEL =METHANATI ON
PRO/II CASEBOOK
AMMONIA SYNTHESIS
Page 43
REACTI ON METHANATI ON REACTI ON SHI FT CONTROLLER UI D=CT1 SPEC STREAM=22, RATE( LBM/ H) , COMP=5, WET, DI VI DE, STREAM=22, & RATE( LBM/ H) , COMP=6, WET, VALUE=3 VARY STREAM=7, RATE( LBM/ H) CPARAMETER I PRI NT, NOSTOP HX UI D=X- 3 HOT FEED=22, M=23, DP=2 COLD FEED=19, V=20, DP=2 CONFI GURE COUNTER, U=90, AREA=1025 HX UI D=X- 2 HOT FEED=11, V=12, DP=2 COLD FEED=20, V=21X, DP=2 CONFI GURE COUNTER OPER CTEMP=675 HX UI D=WHB2 HOT FEED=12, V=13X, DP=2 OPER HTEMP=400 HX UI D=WHB1 HOT FEED=8_ R1, V=9, DP=2 DEFI NE DUTY( BTU/ HR) AS HX=H- 1, DUTY( BTU/ HR) , MI NUS, HX=X- 1A, & DUTY( BTU/ HR) HX UI D=X1- B HOT FEED=9, V=10X, DP=1 DEFI NE DUTY( BTU/ HR) AS HX=X- 1A, DUTY( BTU/ HR) , TI MES, - 1 FLASH UI D=CW2 FEED 23 PRODUCT M=24 I SO TEMPERATURE=100 FLASH UI D=D- 2, NAME=COND SEP FEED 24 PRODUCT V=25, W=24W ADI ABATI C COMPRESSOR UI D=CP- 1, NAME=1ST STAGE FEED 25 PRODUCT V=26, W=WA1 OPERATI ON CALCULATI ON=GPSA, PRES=700, EFF =95 COOLER ACDP=5, ACTEMP=95 COMPRESSOR UI D=CP- 2, NAME=2ND STAGE FEED 26 PRODUCT V=27, W=WA2 OPERATI ON CALCULATI ON=GPSA, PRES=1950, EFF=95 COOLER ACDP=5, ACTEMP=95 COMPRESSOR UI D=CP- 3, NAME=3RD STAGE FEED 27 PRODUCT V=28A OPERATI ON CALCULATI ON=GPSA, PRES=4960, EFF=95 COOLER ACDP=5, ACTEMP=95 STCALCULATOR UI D=SEP1, NAME=WATER REMOVL FEED 28A, 1 OVHD L=WA3, TEMPERATURE=95 BTMS V=28, TEMPERATURE=95 FOVHD( M) 1, 1, 1 FOVHD( M) 2, 13, 0 OPERATI ON STOP=ZERO CALCULATOR UI D=REFC, NAME=REF_RATE SEQUENCE STREAM=31R DEFI NE P( 1) AS STREAM=31R, RATE( LBM/ H) , TOTAL, WET PROCEDURE I F ( R( 1) . GT. 0) GOTO 100 R( 1) = P( 1) 100 CALL SRXSTR( SMR, R( 1) , 31R) RETURN COMPRESSOR UI D=CP- 4, NAME=RECYCLE COMP FEED 31R PRODUCT V=33 OPERATI ON CALCULATI ON=GPSA, PRES=4950, EFF=95, WTOL=1E- 5 HX UI D=X- 4 HOT FEED=S1, M=34, DP=50 COLD FEED=36A, V=38, DP=50
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AMMONIA SYNTHESIS
Page 44
CONFI GURE COUNTER OPER CTEMP=84. 999 MI XER UI D=M1 FEED 33, 28 PRODUCT M=S1 FL ASH UI D=D- 7, NAME=SEC SEP FEED 34 PRODUCT V=36, L=37 I SO TEMPERATURE=39. 999, PRESSURE=4840 METHOD SET=2 FLASH UI D=DM2 FEED 36 PRODUCT V=36A I SO TEMPERATURE=39. 999 FLASH UI D=DM1 FEED 37 PRODUCT L=37A I SO TEMPERATURE=39. 999 EQUREACTOR UI D=RX- 7, NAME=CONVERTER FEED 38B PRODUCT V=29A OPERATI ON ADI ABATI C, PHASE=V, DP=30, TEMP=900 RXCALCULATI ON MODEL=STOI C RXSTOI C RXSET=1 REACTI ON 1 BASE COMPONENT=6 APPROACH DT=20 HX UI D=FDEF, NAME=RX EFFL EXCH HOT FEED=29A, V=29, DP=30 COLD FEED=38, V=38B, DP=30 CONFI GURE COUNTER OPER HTEMP=430 FL ASH UI D=D- 6, NAME=PRI SEP FEED 29 PRODUCT V=31, L=32 I SO TEMPERATURE=84. 999, PRESSURE=4660 METHOD SET=2 FLASH UI D=DM3 FEED 31 PRODUCT V=31A I SO TEMPERATURE=84. 999 FLASH UI D=DM4 FEED 32 PRODUCT L=32A I SO TEMPERATURE=84. 999 CALCULATOR UI D=BD1, NAME=PURGE RATE SEQUENCE STREAM=31X DEFI NE P( 1) AS STREAM=28, RATE( LBM/ H) , COMP=7, WET DEFI NE P( 2) AS STREAM=32A, RATE( LBM/ H) , COMP=7, WET DEFI NE P( 3) AS STREAM=37A, RATE( LBM/ H) , COMP=7, WET PROCEDURE V( 1 ) = P( 1 ) - P( 2 ) - P( 3) I F ( V( 1) . L E. 0 . 001) V( 1) = 0. 001 CALL SRXSTR( SMR, V( 1) , 31X) RETURN SPLI TTER UI D=SP1, NAME=PURGE FEED 31A PRODUCT M=31P, M=31RA OPERATI ON OPTI ON=FI LL SPEC STREAM=31P, RATE( LBM/ H) , COMP=7, WET, DI VI DE, STREAM=31X, & RATE( LBM/ H) , COMP=7, WET, VALUE=1, RTOLER=1E- 5 FLASH UI D=DUM1 FEED 31RA PRODUCT M=31RB ADI ABATI C CALCULATOR UI D=CAL1, NAME=SET_RATE SEQUENCE STREAM=31RB DEFI NE P( 1) AS CALCULATOR=REFC, R( 1) PROCEDURE CALL SRXSTR( SMR, P( 1) , 31RB) RETURN
PRO/II CASEBOOK
AMMONIA SYNTHESIS
Page 45
