I d . Eng. Chem. Process Des. Dev. 1981, 20, 425-433
425
Process Simulation of Ammonia Plant Chandra P. P. Singh" Gulf Science d Technology Company, Pittsburgh, Pennsylvania 15230
Deokl N. Saraf Depadmnt of Chemical Engineering, Indian InstirUte of Technolcgy, Kanpur 2080 16, India
Models of various important process units comprising an ammonia plant, which were validated against extensive plant data, have been unified in an overall plant simulation. This simulation has been used to check the performance of two plants, each at two levels of operation, one of which is based on naphthas whereas the other one uses natural gas as the raw material. The calculated compositions, temperatures, and flow rates at various points along the process stream, including the recyle around the ammonia synthesis reactor, are in very good agreement with the corresponding measured values. This shows the generality and accuracy of the process unit models as well as that of the overall simulation. Effects of errors in the models on calculated rate of ammonia production, pressure, and concentration of inerts in the ammonia synthesis loop are also illustrated.
Introduction A mathematical description of all the important process units, taking into consideration all the important physical and chemical processes taking place inside them, is necessary for a successful design and operation of a plant. Moreover, when a computer is used for control of a chemical process, a mathematical description of the process units is necessary to perform such tasks as feed-forward and feed-back control, as well as off-line and on-line optimization. This paper presents such a mathematical description of a single-train ammonia plant. Figure 1 shows a simplified flow diagram of the process. The important processes involved in the production of ammonia from a hydrocarbon feedstock, such as natural gas or naphthas, are steam reformation, shift conversion, and ammonia synthesis reaction. These have been dealt with in the following sections in the same order. The processes of removal of COZby absorption and conversion of CO to methane are of relatively lesser importance and hence the performance of these units has been assumed to be ideal. Steam Reformation Primary Reformer (PR). The feed to PR (hydrocarbon-steam mixture) enters the catalyst packed reformer tubes at the top. Endothermic heat of the reaction as well as heat to increase process gas temperature is supplied by the large number of burners distributed uniformly over the walls of the reformer. Rate of heat transfer to the process gas as well as the rates of reactions affects the rate of reformation. The model used in this work (Singh and Saraf, 1979a) considers the problem of heat transfer simultaneous with the reactions taking place in the catalyst tubes. Rate Equations. The hydrocarbons subjected to steam-reforming range from natural gas to naphthas in the boiling range up to 220 "C. Higher hydrocarbons in the feed are assumed to hydrocrack to methane at the inlet to PR. The following pair of reactions describes the subsequent conversion CHI + HzO = CO + 3Hz (1) CO + HzO = COZ + Hz (2) In addition to the above reactions, the ammonia synthesis reaction is the only important reaction in the overall 0196-4305/81/1120-0425$01.25/0
process. The synthesis reaction is expressed stoichiomtrically as 0.5Nz + 1.5& = NHB (3) To evaluate the composition of the process gas a t the PR inlet, hydrocracking of the higher hydrocarbons in the feed to methane is expressed by the relation 3CH, + (2 - y)HZO = (1 + y)CH, + (2 - y)CO (4) Using appropriate values of the preexponential factors and activation energies and considering the effect of pressure simultaneous with the pore diffusion resistance, the rates of reactions in the reformer tubes (reactions 1 and 2) are expressed as
Heat Transfer. Radiation is the prevailing mode for transfer of heat from flue gas and flames to the tube walls. The total heat transfer is the sum of the transfers from these two sources (7) 4 = 4- + 9f=t where qm and qW are the rates of heat transfer to an unit area on the reformer tube from flue gas and flames, respectively
0 1981 American Chemical Society
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Ind. Eng. Chem. Process Des. Dev., Vol. 20, No. 3, 1981
diagram for an ammonia plant,
Secondary Reformer (SR).The feed to the secondary reformer consists of PR exit to which air is mixed. Mixing of air leads to adiabatic combustion of H2, CO, and CHI molecules in the gas. The rate as well as equilibrium for burning (reaction with oxygen) of each of them is different and therefore for evaluating the exact composition of the combustion product, all the combustion reactions should be considered in simultaneous equilibrium (Davis and Lihou, 1971). However, the rate at which hydrogen burns is much faster than the other two. Therefore, in this work, it has been assumed that the extent of combustion of CO and CH4 is negligible compared to that of H2.Close agreement was found between the combustion product temperature calculated on this basis and that calculated by Davis and Lihou (1971). The amount of air required is calculated as follows
Eg = .Ti; Et = .Tt; Ef= UT:; u is the Stefan-Boltzmann constant; T,, Tg,and Toare the adiabatic flame, gas, and is the total exouter tube surface temperatures; (?%,)R change area between flue gas and reformer tubes, and (Gt)R,black is the total exchange area when tube surfaces are assumed to be black; nt and nB are the total number of reformer tubes and burners; tg, tt and tf are emissivities of gas, tubes and flames;At, AR and Af are the surface areas of tubes, refractory walls, and assumed spherical surface of flames;and VF, is the view factor of tubes with respect to the refractory walls. The rate of transfer of heat from a unit area of the inner walls of the reformer tubes to the reacting gas mixture is obtained from the equation qin = hin(Tin - T> (12) where Ti,is the inside tube surface temperature and T is the temperature of the process gas. The inside heat transfer coefficient is 0.4 times the value obtained from Beek's correlation (1962). Assuming that there is no transfer of heat in the axial direction at any point inside the reformer, the conduction of heat through reformer tube is expressed as 2K(To - Tin) (13) q = do 1n do/din and heat transfer to a unit area of the outer surface is related to that from the inner surface to the gas by = dinqin/do (14) where doand dh are outer and inner diameters of reformer tubes and Toand Ti,are the respective surface temperatures. It is assumed that the performance of a single reformer tube is representative of any other tube in the furnace. All the properties are uniform throughout a general cross section of the catalyst bed, and axial diffusions of heat and mass are negligible. With these assumptions the balances of mass and heat over a differential reformer tube section of length dz yields the model differential equations. Pressure drop across the differential section is obtained from Ergun's correlation (1952). Integration of the model differential equations along the length of the reformer yields values of all the variables throughout the reformer. Adiabatic flame temperature, T,, is evaluated from flow rate of fuel and oxygen in the flue gas. At each step of integration, Toand Ti, are evaluated from the iterative application of eq 7 , 12, 13, and 14. Details of the model along with the results showing its validity under diverse conditions of operation are presented elsewhere (Singh and Saraf, 1979a).
where P,pH,, PCo,and P c h are the inlet rates of flow of air,H2,CO, and CH4,respectively, and xaN2and X*O, are the mole fractions of N2 and O2in air. Amounts of Hz, N2, and Ar in the combustion product are obtained from the following component balances
FH2 = PH2- 2x"o;P
(16)
FN2 = PdN2
(17)
and
F h = P,,h*~"h (18) where FH2, FN2, and FA are the rates of flows of H2, N2, and Ar in the process gas after combustion is complete. Amounts of the other components in the PR exit remain unaffected since they do not participate in the reaction with air. Temperature of the process gas after combustion and before it reaches SR catalyst bed is calculated by equating the s u m of sensible heats in air and PR exit and the heat from combustion to the sensible heat in combustion product, by a trial and error procedure. Rate Equations. As in PR, a nickel catalyst is used in this reactor also with the only difference that the latter is capable of standing relatively higher temperatures. Reactions 1 and 2 describe the conversion. At the SR temperature conditions, the rate of reaction 1 is so fast that it could be assumed that methane molecules hydrocrack instantaneously at the catalyst surface and equilibrium is attained. In such situations the reaction is controlled by the rate at which the reactant diffuses to the catalyst surface from the bulk gas and the rate of reaction is given by R1
= Kmam(Cc, - Cac&)
(19)
where CcH4and C*chare CHI concentrations in the bulk gas and at the surface, respectively, K, is the mass transfer coefficient between bulk gas and solid surface, and a, is the external surface area per unit mass of the catalyst pellets. c",, and the catalyst surface temperature, P,are evaluated from an energy balance on the pellet which can be expressed as
where C is the total concentration of the process gas. For the solution of the above correlation, which involves trial and error calculations, reaction 2 is assumed to be absent.
Ind. Eng. Chem. Process Des. Dev., Vol. 20, No. 3, I981 427
Otherwise, the rate of reaction 2 is described by eq 6. The assumptions and material, energy and momentum balance equations for this reactor are the same as those for PR except that no heat is supplied to the process gas in SR. Integration of the model differential equations, along the length of SR yields temperature, composition, and pressure throughout the catalyst bed, in close agreement with the measured values (Singh, 1978). Shift Conversion The conversion is expressed by eq 2. It is carried out in two consecutive adiabatic reactors which use different catalysts and work under different conditions of temperature. In the first reactor, Le., the high temperature shift reactor (HT), which uses a chromia promoted iron oxide catalyst, reaction is carried out ,between 350 and 500 "C. The low temperature shift reactor (LT) uses a CuO-ZnO catalyst with various promoting components and operates in the range of 180-250 O C . The rate of shift reaction over the shift catalysts (of FPDIL, India) are given by the general expression
where Rf,A , Pf,Sf,and E f f are factors which account for the effects of reduction in specific surface area due to sintering, aging, pressure, H a concentration in process gas, and diffusional resistances, respectively; u2.and e, are preexponential factor and intrinsic energy of activation for reaction 2, and xco and x*co are mole fractions of CO in the reactor and in equilibrium conditions. Due to low operating temperature and insignificant presence of HzS, reduction in rate due to sintering and H2S are negligible for the LT; i.e., Rf and S, are both equal to 1 for the LT. Values of u2 and e, and the various factors in the rate expression 2 are different for HT and LT. Details for evaluating the various factors numerically along with validations of the models against extensive plant data are available in the references (Singh and Saraf, 1977,1980). Ammonia Synthesis Equation 3 describes the reaction, and the following modified Temkin equation (Dyson and Simon, 1968) gives the intrinsic rate of reaction
P u r 9 gas Mp k mole/ hr
B
Recycle gas M'k,mOIelhr
C
Figure 2. Recyle scheme of ammonia synthesis.
The outlet from one reactor makes the feed to the subsequent one. Air and steam added at SR and H T inlets, respectively, are accounted for by mass balance. Temperature changes associated with heat removal in between any two reactors is not considered; instead the measured values of temperature after cooling are used in simulation calculations. The ammonia synthesis reactor involves recyle of a major part of the product gas. This recyle gas mixed with makeup synthesis gas, i.e., methanator exit, is the synthesis reactor feed. The recyle scheme of the ammonia synthesis reactor is shown schematically in Figure 2. The flow and composition of purge, recyle, and feed is obtained from the following component and overall mass balance equations. The components are numbered for the following equations as: 1, N,; 2, H,; 3, NH3; 4, CH,; and 5, Ar. Mass Balance around Synthesis Loop (Figure 2). Paj" = i@.(xip - ( U ~ ~ . X P-) a3rWr*
j = 1, 2, 4, 5 (24)
Mass Balance around Point A (Figure 2).
P . x j m + iW.x; = Mf-xif j = 1, ..., 5
(25)
Mass Balance around Point D (Figure 2).
Me.x; = (W+ i W ) - x f M e . ~ 3= e (Mp
j = 1, 2, 4,5
+ iW).x3' + MP'*
(26) (27)
Component Balances for Various Streams. 5
Cxi" = 1 j=l
where kgl is the velocity constant of the reverse reaction 3; f N 2 , fH2, fm8are the fugacities of nitrogen, hydrogen, and ammonia, respectively, Kq3 is the equilibrium constant of reaction 3, and b is a constant. The velocity constant is expressed as a function of temperature in Arrhenius equation form as
(28)
5
Ex/= 1
j=l 5
CXj" = 1
(30)
1'1
(23)
where u j and egl are the preexponential factor and the intrinsic energy of activation for reaction 3. The effect of mass transfer resistance on the rate of reaction is accounted for by the effectiveness factor, q, which is evaluated by solving the basic diffusion equation for a catalyst of spherical geometry. Details of the model and its validation against plant data are presented elsewhere (Singh and Saraf, 197910). Overall Plant Simulation There is no recyle of process gas between the point of inlet to the PR and the exit from methanator (Figure 1).
5
Cxi' = 1 ]=1
and xjP = x; j = 1, ..., 5
(32)
where M is molar flow rate, x j is mole fraction of component j , and superscripts e, f, m, p, prod, and r represent the synthesis reactor exit, feed, makeup, purge, product, and recyle streams, respectively. There are 25 unknowns and 22 independent equations (24 through 32). Therefore values of three variables are to be known to obtain the flow rates and compositions of all the streams in the synthesis
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Ind. Eng. Chem. Process Des. Dev., Vol. 20, No. 3, 1981
Table I. Measured and Calculated Data for Naphtha Based Ammonia Plant number of PR tubes diameter of PR tubes heated length of PR tubes SR catalyst HT catalyst LT catalyst synthesis catalyst
160 115.3 mm i.dJ147.7 mm 0.d. 11.36 m 25 tonsl19.4 m3 75 tonsl58.2 m3 73 tons/56 m3 57.2 tons/20.9 m 3
process inlet PR
case I
case I1
process naphtha flow, kg/h steam flow, kg/h
1 4 100 67200
11900 56000
location
CO Case I.
exit PR exit SR exit H T exit LT makeup syn. gas inlet NH, reactor exit NH, reactor recycle gas purge
exptl calcd exptl calcd exptl calcd exptl calcd exptl calcd exptl calcd exptl calcd calcd calcd
CO,
H,
CH,
process inlet PR pressure, kg/cm temperature, "C SR inlet air temperature, "C H T inlet temperature, "C L T inlet temperature, "C synthesis reactor inlet temperature, "C 1st bed 2nd bed 3rd bed pressure at synthesis reactor inlet, kg/cm2
N,
Ar
NH,
case I
case I1
24.8 443 231 377 202
24.3 440 231 364 200
395 412 406 228
389 428 404 214
Press., t e y ~ , SIG kg/cm2 C ratio
flow,N m3/h
Analysis of Process Gas at Different Points in Ammonia Plant
9.93 11.21 13.36 13.18 3.21 3.08 0.31 0.33
17.43 16.97 13.06 12.74 21.42 20.61 23.29 22.71
64.16 64.44 52.26 52.14 55.86 56.38 57.03 57.60 74.59 74.19 65.90 66.52 53.46 53.87 63.40 63.40
8.13 7.38 0.36 0.34 0.32 0.31 9.32 0.30 0.80 0.83 6.54 6.50 7.32 7.49 8.81 8.81
20.63 21.34 18.95 19.38 18.81 18.83 24.32 24.68 22.03 21.78 18.12 17.50 20.60 20.60
0.23 0.26 0.24 0.24 0.24 0.23 0.29 0.30 2.51 2.35 3.10 2.71 3.19 3.19
21.90 21.75 21.10 21.10 20.50 20.00
3.02 2.84 18.00 18.43 4.00 4.00
761 770 905 902 445 443 220 221
0.927 0.864 0.703 0.693 0.530 0.527 0.480 0.487
118 100 118 839 143 100 145 980 143 100 145 980 143 100 145 980 72 900 74 898 248 000 258 805 216 000 224 741 183 907 7 052
0.821 0.892 0.674 0.694 0.550 0.546 0.510 0.508
98 800 98 700 123 900 122 037 123 900 122 037 123 900 122 037 63 700 62 836 226 000 231 393 197 000 202 353 168 557 4 979
228 222
total amount of ammonia produced, tons/day: plant, 619; calcd, 616 exit PR exit SR exit HT exit LT makeup syn. gas inlet NH, reactor exit NH, reactor recycle gas purge gas
Case 11. Analysis of exptl 12.60 17.20 calcd 11.32 17.01 exptl 14.10 12.80 calcd 13.36 12.42 exptl 3.10 21.20 calcd 3.11 20.36 exptl 0.30 22.84 calcd 0.32 22.51 exptl calcd exptl calcd exptl calcd calcd calcd
Process Gas at Different Points in Ammonia Plant 63.30 6.90 21.60 775 64.38 7.29 21.65 772 50.78 0.10 21.90 0.20 21.00 925 52.26 0.29 21.42 0.25 20.92 905 55.60 0.10 19.30 0.20 19.90 436 56.58 0.25 19.47 0.23 431 57.30 0.28 19.03 0.25 19.30 220 57.75 0.24 18.95 0.23 219 73.99 0.67 25.01 0.33 74.22 0.73 24.75 0.30 65.14 6.35 22.32 2.65 3.54 214 65.16 6.91 21.81 2.84 3.28 53.62 6.97 18.33 2.76 18.32 207 52.98 7.90 17.77 3.25 18.10 61.78 9.21 20.72 3.79 4.5 61.78 9.21 20.72 3.79 4.5
total amount of ammonia produced, tonslday: plant, 543; calcd, 525
loop. The three variables selected for the purpose are concentration of inerts in the recyle stream, i.e., ( x l + x;), concentration of ammonia in the recyle stream, i.e., x i , and the concentration of ammonia at the reactor exit. The values of the first two variables are taken from the plant data whereas that of ammonia concentration in the exit stream is assumed and the flow rate and composition of the feed stream are calculated. The synthesis reactor model evaluates the ammonia exit concentration. The assumed and calculated values are matched by an iterative procedure to arrive at the final results.
Results and Discussion Validation of the Simulation. The calculated and measured values of process and operating variables at various points in two ammonia plants, each at two levels of operation, are presented in Tables I and 11. It should be noted that the measured rates of ammonia production,
in the various cases (bottom lines of the tables), are based on weight measurements and are independent of the concentration measurements of ammonia in various streams of the synthesis loop (Figure 2). Therefore the rate of ammonia production based on the measured concentration of ammonia in the process streams may not necessarily equal the measured rate of ammonia production reported in the bottom lines of Tables I through 111. Measured flow, temperature, and composition of the process gas at the exit from PR, SR, HT, LT, and the synthesis reactor, for all the cases, compare very well with the corresponding calculated values. For all the four cases, measured and calculated pressure drops across the PR and SR are also in very good agreement. Similar agreements exist for the flow and composition of various streams around the synthesis loop. The maximum difference between the calculated and measured rates of ammonia production are within 4 % . In addition to being small, the
Ind. Eng. Chem. Process Des. Dev., Vol. 20, No. 3, 1981 429 Table 11. Measured and Calculated Data for Natural Gas Based Ammonia Plant number of PR tubes diameter of PR tubes heated length of PR tubes SR catalyst HT catalyst LT catalyst synthesis catalyst
200 96 mm i.dJ133.8 mm 0.d. 12.2 m 14.83 tonsl12.4 m3 53.5 tonsl44.23 m3 70.3 tonsI50.2 m3 52.8 tonsl20.26 m3
process inlet PR
case I
case I1
process natural gas inlet, N m3/h steam flow, kg/h pressure, kg/cmz
13 920 61000 35.0
12 014 55000 32.3
location
CO
CO,
H,
Case I. Analysis of exptl 10.17 12.27 calcd 9.56 12.56 exptl 12.20 9.40 calcd 11.96 9.44 exptl 2.16 17.62 calcd 1.83 17.52 exptl 0.18 18.93 calcd 0.16 18.87 exptl calcd exptl calcd exptl calcd calcd calcd
exit PR exit SR exit HT exit LT makeup syn. gas inlet NH, reactor exit NH, reactor recycle gas purge gas
CH,
process inlet PR
case I
case I1
temperature, "C air temperature at SR inlet, "C steam added at HT inlet, kglh HT inlet temperature, "C LT inlet temperature, "C synthesis reactor inlet temperature, "C 1st bed 2nd bed 3rd bed pressure at synthesis reactor inlet, kglcm
456.4 460.0
447.4 460.0
N,
Ar
NH,
7200
6000
361 198
351 200
385 433 415 235
406 430 401 210
S/G ratio
flow N m3/h
713 196 962 957 425 424 213 208
0.878 0.898 0.694 0.111 0.649 0.653 0.630 0.626
117 100 116 423 144 300 144 529 151 600 153 489 151 600 153 489 78 100 76 119 242 160 244 898 216 000 209 509 168 780 5 518
193 184 932 921 415 414 209 207
0.910 0.931 0.135 0.148 0.670 0.667 0.650 0.649
103 500 101 098 125 800 121 553 134 000 135 019 134 000 135 019
press., t e p p , kg/cmz C
Process Gas at Different Points in Ammonia Plant 67.22 69.61 55.58 55.73 59.40 59.83 60.29 60.48 74.55 14.39 67.03 67.25 53.90 53.38 64.31 64.31
9.83 8.27 0.32 0.34 0.23 0.31 0.34 0.31 0.75 0.57 5.46 5.65 6.30 6.59 7.90 7.90
22.11 22.26 20.12 20.31 20.01 20.03 24.37 24.14 22.19 21.80 17.80 17.08 20.44 20.44
0.38 0.21 0.23 0.24 0.25 0.23 0.32 0.30 2.56 2.95 2.90 3.44 4.10 4.10
31.30 31.56 30.20 30.40 29.50 28.10
2.76 2.35 19.10 19.51 3.25 3.25
235 228
total amount of ammonia produced, tonslday: plant, 618; calcd, 641 Case 11. Analysis of Process Gas at exptl 9.86 11.79 69.03 9.32 calcd 9.90 12.01 69.59 8.90 exptl 12.40 9.60 55.37 0.33 calcd 12.52 9.26 55.43 0.33 exptl 1.18 18.05 60.02 0.30 calcd 1.23 18.13 59.80 0.30 exptl 0.13 18.00 61.60 0.25 calcd 0.13 19.02 60.24 0.29 exptl 74.37 0.49 calcd 14.32 0.52 exptl 64.85 5.48 calcd 65.54 5.58 exptl 52.61 6.23 calcd 53.00 6.40 calcd 62.18 7.51 calcd 62.18 7.51
exit PR exit SR exit HT exit LT makeup syn. gas inlet NH, reactor exit NH, reactor recycle gas purge gas
Different Points in Ammonia Plant 22.10 22.25 20.09 20.22 19.63 19.99 24.72 24.18 21.54 21.91 18.21 17.76 20.82 20.82
0.40 0.36 0.36 0.32 0.39 0.31 0.42 0.38 5.16 4.01 4.63 4.68 5.49 5.49
28.80 29.12 28.00 28.30 27.50 27.00
2.91 2.89 18.32 18.16 4.0 4.0
216 210
65 383 228 300 236 198 204 000 205 619 170 815 4 521
total amount of ammonia produced, tons/day: plant, 548; calcd, 553 Table 111. Compositions and Flow Rates to the Ammonia Synthesis Reactor Based on Measured Flow and Composition of Makeup Gas process gas stream
H, makeup synthesis gas inlet NH, reactor outlet NH, reactor
CH,
N,
Ar
exptl 74.59 0.80 24.32 0.29 exptl calcd exptl calcd
65.90 69.48 53.46 57.64
6.54 6.59 7.32 7.52
22.03 18.71 18.12 14.11
NH,
flow N m3/h 72 900
2.51 3.02 248 000 2.38 2.87 259603 3.10 18.00 216 000 2.73 18.00 226332
total ammonia produced tonslday: plant, 619; calculated, 601
differences between measured and calculated values are randomly distributed. For example, methane concentra-
tion in Table I1 Case I, shows a higher calculated value at the HT exit, a lower value at the PR exit and in the make-up gas and almost equals the measured values at all other points. As another example, the calculated methane concentration at the PR exit is observed to be less than the corresponding measured values in three cases (Table I, Case I and Table 11, Cases I and 111),whereas in the remaining case (Table I, Case 11) the reverse is true. Similar observations can be made for concentrations of other components and temperatures as well. Thus it can be claimed that the overall simulation represents the plant very well. The above discussed results may be expected in view of the fact that models for each of the process units, namely, PR, SR, HT, LT, and ammonia synthesis reactor, have independently been validated against extensive plant data. The ranges for the variables used in the present study are
430
Ind. Eng. Chem. Process Des. Dev., Vol. 20,No. 3, 1981
rather limited, compared to those used in the validations of unit models (Singh and Saraf, l977,1979a,b, 1980). (It is worthwhile to mention here that this design, i.e., single train, of an ammonia plant is self-sufficient in steam and operating the plant a t relatively lower capacities are impractical due to additional steam requirements). However, there are some limitations to the applicability of the unit models, based on once-through data, to the recyle systems. For example, Table 111presents the calculated composition and flow rates of process gas at the inlet and outlet of the synthesis reactor based on the measured composition and flow rate of makeup synthesis gas in Table I, Case I. A comparison between measured and calculated composition in Table I11 and that in Table I, Case I clearly shows that the difference between measured and calculated values based on the measured makeup gas composition (Table 111)are much higher than the corresponding ones for the overall simulation (Table I, Case I. In the latter case (Table I, Case I) the differences are much smaller at various points in the recyle loop (Figure 2), whereas in the former case (Table 111)the difference between measured and calculated nitrogen concentration, at the inlet to and exit from the synthesis reactor, is around 20%. The differences between measured and calculated hydrogen concentrations are also much higher in this case (Table 111). It is important to note here that the differences between the measured and calculated values of makeup gas rates and compositions are very small (Table I, Case I) whereas the corresponding calculated values for the recyle system show large differences (Table 111). This shows how the magnitude of an error in the once-through stream is amplified in the recyle system which represents the overall performance of the plant. The above example also shows the advantage of using the overall process simulation starting with the input to the plant. Sensitivity Analysis. In a single train ammonia plant, the effect of the performance of any of the process units on the plant performance is not limited to that of the composition, temperature, and pressure of the exit stream. For example, the PR can be operated at different throughputs, and the exit composition, temperature, and pressure can be maintained constant. However, at a lower throughput the amount of flue gas generated in the PR may not be sufficient to obtain enough high-pressuresteam to run the compressors at the required capacity. This would make the operation of the plant impractical. Similarly, other changes affect the plant performance in various complex ways. To account for all the effects of change in the values of a particular variable is outside the scope of this work, since the models for heat recovery, steam generation, compressor performance, etc. are not included. However, this process simulation can be used to illustrate the effects of poor performance of the process units or the effects of poor representation of the process units by the corresponding models. The effect of poor performance of the reactors upstream would be reflected in the makeup synthesis gas as reduction in the rate of generation, higher inerts (CHI + Ar) concentration and/or H2/N2 ratio different from 3. Conversely, a change in composition or flow rate of the synthesis gas reflects on the performance of the process units upstream. Therefore, the composition and flow rate of makeup synthesis gas on the performance of the plant has been selected for study. Makeup Gas Rate. Figure 3 shows the rate of production of ammonia and the flow rate of process gas at the inlet to the ammonia synthesis reactor, both as functions of rate of makeup gas flow. The levels of purge rate (as
EE
-
UI
-1.8
m L 0 s
a
-
3001
1
1
I
1
1
0.
1.6
A
a fraction of makeup gas), concentration of ammonia in the purge or recyle stream, and average pressure in the system are maintained constant. The H2/N2ratio is 3 and inlet temperatures to the catalyst beds are the same as that in Table I, Case I. As shown, the rate of ammonia production decreases linearly with the decrease in the rate of production of the makeup synthesis gas. The rate of flow of the process gas to the reactor increases with increase in the makeup gas rate. The rate of increase in process gas flow, with increase in makeup gas flow, is higher at higher values of the latter due to decrease in the concentration of ammonia at the synthesis reactor exit. The rate of production of ammonia increases since the increase in throughput more than compensates for the decrease in ammonia concentration in the exit stream. Inerts (CHI + Ar) in the Makeup Gas. The change in the rate of ammonia production subject to change in the concentration of inerts (CHI + Ar) in the makeup gas, and the corresponding variation in the process gas flow rate, is shown in Figure 4. The level of inerts in the recyle stream and pressure is held constant. To maintain constant level of inerts in the recyle gas, the purge rate increases with increase in inerts concentration of the makeup gas and the production of ammonia decreases. The decrease in the rate of ammonia production is linear since the amount of gas lost in the purge is directly proportional to the inerts concentration in the makeup gas. The increase in the rate of process gas inlet to the synthesis reactor is slightly higher at lower inerts Concentration in the makeup gas. The increase in rate of process gas feed at lower inerts concentration in makeup gas results from the counteracting effects of decrease in the purge rate and increase in the equilibrium concentration of ammonia. The results in Figure 4 show that the former has a more significant effect on the flow rate. Figure 5 shows the effect of change in the concentration of inerta in the makeup synthesis gas when the recyle and purge rates are held constant. For a constant purge rate, the rate of ammonia production remains constant. However, the level of inerts in the recyle stream and the average pressure in the reactor increase. The former increases
Ind. Eng. Chem. Process Des. Dev., Vol. 20, No. 3, 1981 431 I
1
0.5
09
1.3
lnerts in makeup gas, %
1.r
2.1
Figure 4. Effect of inerts in makeup gas on ammonia production and synthesis reactor throughput.
linearly with the inerts concentration in makeup gas whereas the rate of increase in pressure is higher at higher inerts concentration. H2/N2Ratio in the Makeup Gas. The effect of Hz/Nz ratio in the makeup gas on average pressure in the synthesis reactor is shown in Figure 6. For a constant rate of makeup gas production, and purge rate, the operating pressure required to achieve the conversion shows a minimum when the H2/N2ratio equals 3. Any change in the Hz/Nz ratio from 3 increases the average pressure in the reactor and the rate of increase in pressure also increases with increase in the difference of H2/N2ratio from 3. The corresponding change in the concentrations of H2 and Nz and H2/Nz ratio in the ammonia reactor exit is shown in Figure 7. As the H2/N2ratio in the makeup synthesis gas changes from 2.8 to 3.2, the ratio of Hz/Nz in the reactor exit changes from 1.05 to 20.2. The corresponding change in mole fractions of hydrogen and nitrogen are from 37.35 to 3.36. It is important to note from the results in Figures 6 and 7 that at H2/N2ratio lower than 3 the decrease in hydrogen mole fraction has a more significant effect on the rate of synthesis reaction compared to the effect of increase in the mole fraction of nitrogen on the same, whereas for H2/N2 ratio greater than 3 the effect of decrease in nitrogen concentration is more important. Thus, the rate of reaction at a given pressure decreases and the pressure required to achieve the same extent of conversion increases with any change in H2/N2 ratio from 3. Also, a small change in the H2/N2ratio in the makeup gas results in a many times greater change in the H2/N2ratio in the reactor and all the streams in the recyle loop. It is clear from the results in Figures 3 through 7 that the performance of an ammonia plant is very sensitive to
1801 0.1
1
0.5 Inerts in
1 1 0.9 1.3 makeup g a s , %
I 1.7
I
2.1
Figure 5. Effect of inerts in makeup gas on synthesis reactor pressure and inerts in recyle gas.
I
2.8
I 2.9
I
3.0
H 2 / N 2 in makeup gas
1 3 .l
2 __c
Figure 6. Effect of Hz/N2 ratio on synthesis reactor pressure.
the rate of makeup gas production, and the concentration of inerts and H2/N2 ratio in the makeup synthesis gas. Low rate of makeup synthesis gas production, higher inerts
432
Ind. Eng. Chem. Process Des. Dev., Vol. 20, No. 3, 1981
subject to changes in any of the input conditions, and meaningful optimization of the plant performance can be obtained. Nomenclature A = area, m2 A , = aging factor a = preexponential factor a,,, = external surface area per unit mass of catalyst, m2/kg
of catalyst
b = a constant C = concentration, kg-mol/N m3
C,
01 2.0
I 2.9 H2/N2
in
I
I
3.0
3.1
makeup 98s
lo
3.2 d
Figure 7. Effect of H2/N2ratio in makeup gas on the H2/N2 ratio in the ammonia synthesis loop.
in the makeup gas, or improper H2/N2ratio decrease the rate of ammonia production. The decrease ie the rate of ammonia production can be reduced and in many cases the negative effects can be compensated by an increase in the operating pressure. However, the required increase in pressure is limited by the capacity of the compressors and availability of steam for running the compressors. In general, low makeup gas rate, higher inerts concentration, or improper H2/N2 ratio would result in reduced production of ammonia. In view of the sensitivity of the rate of ammonia production to the changes in the makeup synthesis gas production, concentration of inerts (CHI + Ar) and H2/N2 ratio in the makeup gas and the fact that each of the process units contribute to the rate of generation and composition of the makeup gas shows that the unit models are very accurate in predicting the performance of the process units, and the unified simulation is capable of predicting the plant performance. However, to have a complete picture of the complex interactions between the various units and their effects on the overall performance of an ammonia plant, any futqre work should include modules for the recovery of heat at various locations and generation of steam from various hot streams and compressors. Conclusions A process simulation of hydrocarbon based single train ammonia plant has been obtained using mathematical models for the important process units and considering the performance of other units to be ideal. The close agreement between measured and calculated values confirms the accuracy of the process unit models and the overall simulation in predicting the plant performance. Analysis based on the simulation shows that the latter is very sensitive to the flow rate and composition of the makeup synthesis gas. Future work on the ammonia plant simulation should include models for all the process units as well as those for heat exchangers and compressors. Such a simulation can be used to study the overall response of the plant,
= average specific heat capacity, kcal/kg K C,, = specific heat capacity of flue gas, kcal/k%K C, = concentration of methane, kg-mol/N m d = diameter, m E = emissive power, kcal/m3 e = energy of activation, kcal/kg-mol K e, = apparent energy of activation, kcal/kg-mol K F = volume flow rate, N m3/h f = fugacity of component j ht= heat of ith reaction, kcal/kg-mol h = heat transfer coefficient, kcal/m2 K h K = conductivity, kcal/m2 K h Kq = equilibrium constant k = rate constant It, = mass transfer coefficient, m/h M = molar flow rate, kg-mol/h nB = number of burners n, = number of tubes P = pressure, atm Pf = pressure factor q = rate of heat transfer per unit area on the outer surface of reformer tubes, kcal/m2 h qftt, q+ = rates of heat transfer to a unit area on reformer tube from flue gas and flames, kcal/m2 h R = rate of reaction, kg-mol/kg of cat. h Rf = reduction factor R, = universal gas constant Sf = sulfur factor T = temperature, K VF = view factor 3c = mole fraction a = stoichiometric coefficient e = emissivity
Subscripts f = flames g = gas
i = reaction i in = inside wall of reformer tubes j = component j o = outside wall of reformer tubes R = refractory wall t = reformer tube Superscripts
a = air e = exit synthesis reactor f = feed to the synthesis reactor m = makeup synthesis gas o = inlet condition p = purge gas prod = product stream r = recyle gas in the ammonia synthesis loop s = catalyst surface ' = reverse reaction * = equilibrium condition Literature Cited Beak, J. Adv. Chem. Eng. 1862. 3, 234. Davis, J.; Lihou, D. A. Chem. Process Eng. 1971. 52, 71. Dyson, D. C.; Simon, J. M. I d . Eng. Chem. Fundam. 1968, 7, 605. Ergun, S. Chem. Eng. Prog. 1852. 48, 89. Singh, C. P. P. Ph.D. Thesis, I..T, Kanpur, India, 1978.
Ind. Eng. Chem. Process Des. Dev. 1981, 20,433-435
Singh, C. P. P.; Saraf, D. N. Ind. Eng. Chem. Process Des. Dev. 1977, 16, 313.
433
Singh, C. P. P.; Saraf, D. N. Ind. Eng. Chem. Process Des. Dev. 1980, 10, 393.
Singh, C. P. P.; Saraf, D. N. I d . Eng. Chem. Process Des. Dev. 1979a, 18, 1.
Received f o r review October 12, 1979 Accepted March 16,1981
Singh, C. P. P.; Waf, D. N. Ind. Eng. Chem. Process Des. Dev. 1979b, 18, 364.
Extraction of Acetic Acid from Dilute Aqueous Solutions with Trioctylphosphine Oxide Janvlt Golob," Vlktor Grllc, and Bojan Zadnlk &Pam"
of Chemistry and Chemical Technology, Edvard Kardeu Unlversky, 61000 Ljubljana, Yugoslavia
Extraction characteristics of trioctylphosphine oxide (TOPO) dissolved in commercial kerosene for recovery of acetic acid from dilute water solutions were studied. Some basic thermodynamic properties such as viscosity, density, interfacial tension, and primary settling time of the extraction system were measured. Extraction isotherms at 20 OC were obtained for 3, 5 , and 18 wt % of TOPO in kerbsene. Finally, a simulation of countercurrent extraction of water solution, containing 5 wt % acetic acid, and regeneration of the solvent was performed on laboratory scale equipment.
Introduction The present work deals with the recovery of acetic acid from dilute aqueous solutions. Removal of acetic acid from waste waters, concentration ranging from 0.5 to 5 wt 90, is of increasingly interest due to economic and environmental reasons (Helsel, 1977). For extraction of acetic acid on an industrial scale, solvents such as ethyl acetate or diethyl ketone have been used (Brown, 1963; Ricker et al., 1980). Recovery of acetic acid with such a conventional solvent is not economically attractive as long as the feed contains less than 3 wt ?& of acetic acid. Therefore, investigations aiming to find solvents that would provide higher values of distribution coefficients were initiated (Helsel, 1977; Wardell and King, 1978; Ricker et al., 1979). Application of trioctylphosphine oxide (TOPO) as an effective solvent for extraction of acetic acid from dilute solutions is discussed in a limited number of papers (Helsel, 1977; Kohn, 1978; Wardell and King, 1978; Ricker et al., 1979). The distribution coefficient varied between 3 and 5. In recent years the use of other organophosphorus compounds and high molecular weight alkylamines dissolved in various diluents (not kerosene) was explored for the acetic acid (Ricker et al., 1979). Ricker et al. (1979) put forward a comparison between TOPO and amine extractants. Due to the ten times higher price of TOPO compared with conventional amine solvents (Alamine 336, Adogen 364), it was concluded that any significant loss of TOPO would have a pronounced negative effect on the economics of an acetic acid recovery process. Another disadvantage of TOPO is its very high boiling temperature which involves difficult recovery from nonvolatile substances (tar) by distillation (Ricker et al., 1980). An advantage introduced by the TOPO/diluent system is that there is no need for additional thermal or chemical treatment of the water phase after extraction and coalescence, providing that insoluble diluent is used. Ricker could not detect traces of TOPO in the water phase after containing 30 wt ?& TOPO in Chevron Solvent No. 25 (Ricker et al., 1979). Some authors also considered co-
Table I. Properties of TOPO-Kerosene Solutions at 20 "C
TOPO concn in kerosene, wt 5% properties
0
kg/m3 786.0 1.18 qo, mPas 33.79 Y, mN/m 15 t, s p0,
3 787.5 1.27 24.76 25
5 790.0 1.34 12.15 30
18 804.5 1.94 10.13 40
extraction of other components such as formic acid, formaldehyde, methanol, etc., present in acetic acid waste streams. Their presence demands additional separation steps for their isolation (Helsel, 1977; Wardell and King, 1978; Ricker et al., 1980). The aim of the present work is to examine the usefulness of TOPO/kerosene solvent. Thermodynamic data of the extraction system, such as the solubility of TOPO, distribution coefficients, viscosity, density, interfacial tension, and primary settling time are needed. Simulation of the extraction/regeneration process on a laboratory scale would also provide useful insight into operation requirements. Extraction Characteristics bf TOPO-in-Kerosene and Acetic Acid-in-Water Solutions The thermodynamic data for extraction of acetic acid from dilute aqueous solutions are available for various TOPO diluents excluding kerosene and do not cover all relevant property data important for the design of the apparatus for extraction and phase separation. Therefore, it was one of the aims of this work to determine experimentally some most important properties of TOPO-inkerosene solutions. The TOPO used in this study was of technical grade (95% purity), manufactured by Fluka AG, Switzerland. Kerosene was commercial grade jet fuel supplied by Petrol Oil Co., with composition as follows: 16.7 wt % aromatica, 40.8 wt % C8-CI4 n-parafins, and 42.5 wt % naphthenes. the solubility tests yielded 21 wt % saturated solution of TOPO in kerosene at 20 "C. Other physical properties
0196-4305/81/1120-0433$0l.25/0 0 1981 American Chemical Society
