Design and Control of a Modified Vinyl Acetate Monomer Process

ARTICLE pubs.acs.org/IECR

Design and Control of a Modified Vinyl Acetate Monomer Process William L. Luyben* Department of Chemical Engineering, Lehigh University, Bethlehem, Pennsylvania 18015, United States ABSTRACT: The vinyl acetate monomer (VAM) process presents several challenging design problems because of the many design optimization variables and several important constraints. The process features a cooled tubular reactor, both gas and liquid recycle streams, two absorbers, and two distillation columns, one involving heterogeneous azeotropic distillation. The major design variables are reactor temperature, reactor size, reactor pressure, ethylene gas recycle and acetic acid liquid recycle. The major constraints are a maximum oxygen flammability limit, a maximum internal reactor temperature limit, and the need to stay above dewpoint temperatures in the reactor. This paper develops an economic optimum flowsheet based on the conditions and parameters provided by M. L. Luyben and B. D. Tyreus over a decade ago as a challenge problem for the academic community for design and control studies (Comput. Chem. Eng. 1998, 22, 867877). The conceptual design developed produces 18% more VAM product for the same oxygen fresh feed flow rate used in the original design. Compared to the original nonoptimized design, the modified design has a reactor that is twice as large with reactor inlet conditions showing lower oxygen concentrations and higher acetic acid concentrations. Operating pressure is lower, liquid acetic acid recycle is larger, and gas ethylene recycle is smaller. A plantwide control scheme that provides effective disturbance rejection and is significantly different than the structure used in the original design is developed.

1. INTRODUCTION The industrial challenge problem presented by Luyben and Tyreus1 has received only limited attention by the academic community despite its rich mixture of important steady-state design and dynamic control features and issues. Three plantwide control studies were reported. All of these control studies are based on the original flowsheet presented in the challenge problem. Luyben et al2 demonstrated that effective plantwide control could be achieved using a conventional PI control structure. Chen et al3 developed a nonlinear dynamic model in Matlab and studied alternative control structures. Olsen et al4 presented other alternative control structures. There was no claim in the original paper that the process design had been optimized from the standpoint of economics. The original flowsheet used a reactor with 622 tubes, 0.037 m in diameter and 10 m in length. The gas recycle was large (939 kmol/h) and the liquid acetic acid recycle was small (84.2 kmol/h). With a fresh oxygen feed of 31.26 kmol/h, the vinyl acetate monomer (VAM) product was 49.56 kmol/h. A conceptual design is developed in this paper that finds the economic optimum equipment sizes and operating conditions, which are significantly different than the original design. The basic unit operations are the same as in the original design, but equipment sizes and operating conditions are quite different. The most significant improvement in performance is an 18% increase in the production of VAM for the identical fresh feed flow rate of oxygen. A bigger reactor is required, but the dominant economic value of the product should easily justify the increase in capital investment. There are three very important and interacting constraints in this process. The vital safety constraint is to keep the oxygen concentration below the flammability limit. The largest oxygen concentration occurs at the reactor inlet. In the original design, this concentration was reported as 7.4 mol % oxygen. In the r 2011 American Chemical Society

proposed flowsheet, the oxygen concentration is only 2.6 mol %, far below the maximum safety limit. The lower oxygen concentrations improve the selectivity since the combustion reaction is lowered. The second constraint is a maximum reactor temperature of 455 K to prevent catalyst degradation. Since the reactor is cooled, the maximum temperature occurs partway down the reactor. With a fixed coolant temperature on the shell side of the reactor tubes and fixed reactor size, this constraint can be maintained by adjusting the flow rate of gas recycle, which is mostly ethylene. The higher the gas recycle, the lower the peak temperature. Of course, higher gas recycle means higher gas compression costs in terms of both capital and energy. The third important constraint is the need to guarantee that there is no liquid in the reactor since the reactions occur in the gas phase. Liquid can form if the temperatures in the reactor are too low, if the pressure is too high, or if the concentration of the highboiling acetic acid is too high. Low reactor temperatures improve selectivity (more VAM and less CO2) because the activation energy of the VAM reaction is lower than that of the combustion reaction, as discuss in the next section. Reactor temperatures are affected by the coolant temperature and the gas recycle. At low coolant temperatures, it is sometimes necessary to increase gas recycle beyond that required to limit the peak reactor temperature in order to keep above dewpoint temperatures. Although selectivity is reduced by running at higher reactor temperatures, there is an important advantage of being able to recover the exothermic heats of reaction through generation of Received: May 25, 2011 Accepted: August 6, 2011 Revised: July 25, 2011 Published: August 07, 2011 10136

dx.doi.org/10.1021/ie201131m | Ind. Eng. Chem. Res. 2011, 50, 10136–10147

Industrial & Engineering Chemistry Research

ARTICLE

Figure 1. Modified flowsheet.

steam in the reactor, which is an energy credit. If the cooling medium temperature on the shell side of the cooled reactor can be operated at 433 K, low-pressure steam can be produced (6 bar) with a value of $7.78 per GJ (Turton6). Higher acetic acid concentrations also improve selectivity and can be achieved by increasing acetic acid recycle. However, more gas recycle may be required to prevent the formation of liquid in the reactor. In addition, energy consumption in the acetic acid vaporizer and in the azeotropic distillation column increase. All of these issues must be considered in the conceptual design of this very interesting and complex process. The modified design has 1100 tubes instead of 622. The acetic acid recycle is much larger, which results in a reactor inlet concentration of 30.28 mol % acetic acid compared to the original design of 11 mol %. The operating pressure is reduced from 8.7 to 6.5 atm. The reboiler duty of the azeotropic column is reduced from 4.65 to 3.47 MW. The original design had a high concentration of the chemically inert ethane (21.6 mol % C2H6) in the gas recycle loop, which was assumed to build up. The ethane concentration in the modified design based on steady-state Aspen Plus simulation is found to be very small. There are only 0.05676 kmol/h of ethane entering in the fresh ethylene feed (99.9 mol % C2H4). Most of this leaves as small impurities in the organic product stream (0.02639 kmol/h C2H4) and in the aqueous stream (0.02688 kmol/h C2H4) with trace amounts in the purge and decanter vent streams. In the dynamic simulation discussed later in this paper, a much larger purge flow rate is required to prevent the

buildup of the inert ethane since Aspen Dynamics predicts very small concentrations of ethane in both products.

2. PROCESS DESCRIPTION Figure 1 shows the modified design. In all the cases presented in this paper, the flow rate of fresh oxygen is fixed at 31.26 kmol/ h, which is the same as used in the original design. The conditions shown in Figure 1 are the economic optimum as developed in a later section in this paper. Fresh acetic acid is combined with an acetic acid recycle stream and fed into a vaporizer along with a gas stream, which is the fresh ethylene feed plus a gas recycle stream. The vaporizer operates at 6.6 atm and 423 K, which requires medium-pressure steam (11 bar, 457 K, $8.22 per GJ). Vaporizer energy consumption is 3.88 MW. The vapor stream is combined with fresh oxygen and enters the cooled tubular reactor at 421 K with a concentration of 2.63 mol % oxygen and 30.28 mol % acetic acid, the remainder being mostly ethylene. The reactor contains 1100 tubes, 0.037 m in diameter and 10 m in length. The coolant temperature is 433 K, and the heat removal in the reactor is 2.79 MW. The overall heattransfer coefficient is that used in the original design (0.17 kW m2 K1). Figure 2 gives temperature and composition profiles in the reactor. The maximum temperature is 455 K with the coolant temperature of 433 K and a gas recycle flow rate of 744 kmol/h. Catalyst properties are taken from the Luyben and Tyreus paper (0.8 void volume, 1935 kg/m3 particle density). 10137

dx.doi.org/10.1021/ie201131m |Ind. Eng. Chem. Res. 2011, 50, 10136–10147


ARTICLE

Figure 2. (a) Reactor temperature profile; (b) reactor composition profiles.

The per-pass conversion of ethylene is 7.3%, of acetic acid is 15%,and of oxygen is 97.8%. The selectivity is 23 (ratio of VAM formed in the reactor to CO2 formed in the reactor). It should be noted that the selectivity of the original process is only 11, so a very significant improvement in the yield of VAM has been achieved. The reactor effluent is cooled to 323 K and fed to a flash drum. The liquid stream is fed to the first distillation column. The gas stream is fed to an absorber in which VAM is recovered by an acetic acid wash stream. The gas from the absorber goes to a second absorber that removes carbon dioxide using a monoethanol amine solution. The gas from the top of the CO2 absorber is compressed and recycled to the vaporizer. The liquid streams from separator and from the VAM absorber are fed to a distillation column whose function is to remove the small

amount of ethylene contained in these streams. The de-ethanizer column (D/E) operates at 27.2 atm. The reflux-drum temperature is 257 K, so refrigeration is required in the condenser. The easy separation between ethylene and VAM is achieved using a 21-stage column and a reflux ratio of 1. The bottoms from the D/E column is fed to the second distillation column whose function is to recover the acetic acid for recycle back to the vaporizer. A 21-stage column is assumed, which is the same as that used in the original design. The ternary system of VAM, water, and acetic acid is highly nonlinear with a heterogeneous azeotrope formed. The overhead vapor from the column is cooled to 310 K and forms two liquid phases in a decanter. The aqueous phase is 99.45 mol % water. The organic phase is 94.41 mol % VAM. A small vent stream leaves the top of 10138



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Figure 3. Txy diagram for ethylene/VAM at 27.2 atm.

the decanter. Reflux to the top of the azeotropic column is a portion of the organic phase from the decanter. This reflux is adjusted to maintain the VAM composition of the organic phase in the decanter and varies as acetic acid recycle, reactor size, and system pressure are varied.

3. KINETICS AND PHASE EQUILIBRIUM The kinetic equations given by Luyben and Tyreus are not easily implemented in Aspen simulations because they are not in the convenient power-law form and would require writing and compiling of a user kinetic subroutine. As a practical engineering alternative, the approach taken was to fit the performance results (yield of VAM and CO2) for the reactor given in the original design by selecting pre-exponential factors and activation energies for the two reactions. 1 C2 H4 þ CH3 COOH þ O2 f CHOCOCH3 þ H2 O 2 C2 H4 þ 3O2 f 2CO2 þ 2H2 O ð1Þ The physical dimensions of the reactor and catalyst properties were fixed at those given in the original design (622 tubes, 0.037 m in diameter and 10 m in length filled with 2590 kg of catalyst). The coolant temperature (406 K) and overall heat-transfer coefficient (U = 0.17 kW m2 K1) were those used in the original design. The reactor inlet conditions were set as those given in the original design in terms of inlet temperature, pressure and all molar flow rates. The ratio of the activation energies of the two competing gasphase irreversible reaction was set at 2.5, which is close to the ratio given in the original design (E1 = 8000 kJ/kmol and E2 = 20000 kJ/ kmol). Then the values of the two pre-exponential factors (R1 and R2) were varied by trial and error until the production rates of VAM and CO2 in the reactor were close to those reported in the original design. The kinetic relationships used in this paper are R 1 ¼ R1 eE1 =RT pE pHAc ðpO Þ0:5 ¼ 5:8 1018 e8000=RT pE pHAc ðpO Þ0:5 R 2 ¼ R2 eE2 =RT pE pO ¼ 6:7 1015 e20:000=RT pE pO

ð2Þ

Overall reaction rates R n have units of kmol s1 kg1. Partial pressures of components pj have units of Pascals. Note that the combustion reaction is assumed to be first-order in both oxygen and ethylene partial pressures. These simplified kinetics are compared in the following paragraph with those given in the original paper that show more complex composition dependence. "

pE pHAc pO ð1 þ 1:7pW Þ ½1 þ 0:583pO ð1 þ 1:7pW Þ½1 þ 6:8pHAc " # pO ð1 þ 0:68pW Þ R 2 ¼ 1:9365 105 e10116=T 1 þ 0:76pO ð1 þ 068pW Þ

#

R 1 ¼ 0:1036e3674=T

ð3Þ Overall reaction rates R n have units of mol min1 g1 in eq 3. Partial pressures of components pj have units of psia. Using the reactor inlet compositions, coolant temperature and reactor pressure in the original paper, the two reactions rates can be calculated: R 1 = 3.203 104 and R 2 = 8.83 106 mol min1 g1. These can be compared with those calculated using the reactor conditions in the modified design shown in Figure 1: R 1 = 3.089 104 and R 2 = 0.152 106 mol min1 g1. It is clear that the reactor conditions in the modified design are more favorable for the production of desired VAM. This demonstrates that the simplified kinetics used in this study have been successful in developing a process with a much higher selectivity, as confirmed by the calculations using the “rigorous kinetics”. The phase equilibrium in the VAM process is quite nonlinear with a heterogeneous azeotrope formed in the ternary system of VAM, acetic acid, and water. The Aspen NRTL phase equilibrium parameters are used in the simulations of the vaporizer, reactor, absorber, de-ethanizer, and decanter. The UNIFHOC physical property package is used in the azeotropic column to account for the association of acetic acid. The valid phases in this column are VLL. The separation in the first D/E column between ethylene and VAM is fairly easy, as shown in the Txy diagram given in Figure 3 at the operating pressure of 27.2 atm. A ternary diagram is given in Figure 4 for the principal components in the azeotropic column: VAM, acetic acid, and 10139



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Figure 4. Ternary diagram for VAM/water/acetic acid at 0.33 atm.

water. The pressure is 0.33 atm, which gives a heterogeneous azeotropic temperature of 312 K that is close to the decanter temperature. The bottoms from the column contains 5 mol % water. The composition of the overhead vapor from the column is close to the azeotropic composition. The column feed is 535.3 kmol/h and the organic reflux is 390 kmol/h, giving the mix point shown in Figure 4 between the feed composition and the organic composition. The component balance line runs through the mix point and connects the bottoms point and the overhead point. The AMINES physical property package is used in the CO2 absorber. The MEA/water solvent fed to the top of the absorber would come from the base of a stripping column, which is not included in the simulation because the stripping column has little economic effect on the major units of the process.

4. FLOWSHEET CONVERGENCE It might be useful to say a few words about the issue of flowsheet convergence for this complex multiunit process. The existence of two recycle streams can present significant problems in these types of processes. However, the two recycles were successfully closed by employing the technique of fixing the total flows (fresh plus recycle) and permitting the makeup streams to vary. For example, the total flow rate of gas to the vaporizer (fresh ethylene feed plus recycle gas) is held constant using an Aspen Plus Flowsheet Design Specification that varied the ethylene fresh feed. A second Aspen Plus Flowsheet Design Specification varied the acetic acid fresh feed to maintain a specified total acetic acid flow rate (fresh plus recycle) to the vaporizer. This convergence method worked quite well and was successful in handling a wide range of values of design variables. 5. EFFECT OF DESIGN VARIABLES The four major design variables in the vinyl acetate monomer process are reactor temperatures (set by the temperature of the

coolant and the flow rate of gas recycle that maintains a 455 K peak temperature), the reactor size (set by the number of tubes), the pressure in the gas loop (vaporizer, reactor, separator, and absorbers), and the flow rate of the recycle acetic acid. An iterative exhaustive enumeration of these four variables was used to find the optimum values of the design variables. The economic objective function was incremental return on incremental investment. For example, as more tubes are used in the reactor, capital investment increases due to the cost of the vessel (more heattransfer area) and the cost of the catalyst. Production of VAM product increases as more tubes are added, but the increase is at an increasingly smaller rate. The point that gives an incremental ROI of at least 30% was selected. The three constraints that must be satisfied are a maximum peak temperature in the reactor, an upper limit on the oxygen concentration of the feed to the reactor (assumed to be 7 mol % oxygen), and the requirement that the material in the reactor is all gas phase (temperature must be above the dewpoint temperature of the mixture at all points in the reactor). Keep in mind that the flow rate of fresh oxygen is fixed for all cases. 5.1. Effect of Reactor Size. The catalyst in the VAM process is quite expensive because it contains palladium. A catalyst cost of $800 per kg is assumed (2.27 wt % palladium at $800 oz), which favors small reactors to reduce capital investment. A reactor coolant temperature of 433 K is initially selected so that the steam generated in the reactor can be used elsewhere in the plant as low-pressure steam. Operating with lower coolant temperature is explored in a later section. The peak reactor temperature is maintained at 455 K by varying gas recycle flow rate using a Flowsheet Design Specification. The acetic acid recycle is initially set at 375 kmol/h. Reactor inlet pressure is initially set at 6.5 atm. With these design variables fixed, the effect of changing the number of reactor tubes is shown in Figure 5. The production of VAM and the consumption of reactants increase as more tubes are used in the reactor, but these flow rates begin to level out for 10140



ARTICLE

Figure 5. Effect of reactor size.

more than about 1200 tubes. Since the product is worth more than the reactants, profit increases. But capital cost of the reactor vessel and the catalyst increase linearly as more tubes are used. The cost of the reactor vessel is based on the heat-transfer area. The economic parameters for chemical costs given by Tyreus and Luyben are used to assess the economics: VAM = $0.971 per kmol, ethylene = $0.442 per kmol and acetic acid = $0.596 per kmol. The cost of oxygen is not considered because the fresh oxygen flow rate is the same in all cases. A “pseudo-profit” parameter is defined as the value of the VAM produced minus the cost of the ethylene and acetic acid reactants minus the energy costs of the compressor, vaporizer, and azeotropic column reboiler plus the value of the reactor steam:

Table 1. Effect of Reactor Sizea Ntubes = 1000 Ntubes = 1100 Ntubes = 1200

profitð$=hÞ ¼ ðproduce kmol=hÞð$0:971Þ ðfresh ethylene kmol=hÞð$0:442Þ ðfresh HAc kmol=hÞð$0:596Þ 3600 s=h ð$16:8=GJÞ ðcompressor work kWÞ 106 kW=GW 3600 s=h ð$8:22=GJÞ ðvaporizer energy MWÞ 103 MW=GW 3600 s=h ð$7:78=GJÞ ðazeo column reboiler MWÞ 103 MW=GW 3600 s=h ð$7:88=GJÞ ð4Þ þ ðreactor steam MWÞ 103 MW=GW The energy costs in the de-ethanizer (reboiler heat input and condenser refrigeration) and in the absorbers remain essentially the same as conditions in the reactor are changed, so they are not considered. Capital investment is the sum of the installed costs of the reactor vessel, catalyst, vaporizer, and compressor. The capital costs of all the other units remain essentially the same as conditions in the reactor change. They are not considered in the economics since an incremental return on investment is used

product gastot

kmol/h kmol/h

57.538 821.6

57.983 796.9

58.202 786.3

ethylene

kmol/h

56.27

56.61

56.79

acetic acid

kmol/h

55.15

55.57

55.77

organic reflux kmol/h

369

371

372

Qvap

MW

3.833

3.807

3.801

Tvap

K

421.6

422.2

423

QRX

MW

2.724

2.784

2.824

QR2 compressor

MW kW

2.638 266.3

2.658 257.2

2.668 253.3

yRXin

mol % O2

2.71

2.66

2.64

mol % C2H4

64.35

63.81

63.57

mol % HAc

29.03

29.66

29.94

xB2

mol % H2O

6.19

6.03

5.97

profit

106 $/y

14.606

14.761

14.841

capital

106 $

5.6159

6.0577

6.5096

a

Conditions: 6.5 atm; HAc total = 375 kmol/h; Tcool = 455 K; Tmax = 455 K.

to select the optimum values of design variables. Cost equations from Douglas5 and Turton et al6 are used to calculate these capital costs. Figure 5 and Table 1 show that profit is increasing but capital investment is also increasing as more reactor tubes are used. The incremental return on investment decreases. Moving from 1000 to 1100 tubes gives an incremental return on investment of 35%. Adding more tubes yields a lower incremental return on investment. Note that the required total gas recycle is set to achieve a 455 K peak reactor temperature. Under these conditions, with a 433 K reactor coolant temperature, the dewpoint limitation is avoided (the process stream leaving the reactor has a vapor fraction >1). Reactor inlet oxygen concentration (2.66 mol %) is 10141



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Figure 6. Effect of HAc recycle.

Table 2. Effect of Acetic Acid Recyclea HAc total = 375 kmol/h

HAc total = 385 kmol/h

HAc total = 395 kmol/h

product gastot

kmol/h kmol/h

57.983 796.9

58.188 800.6

58.388 804.3

ethylene

kmol/h

56.61

56.74

56.88

acetic acid

kmol/h

55.57

55.76

55.96


371

390

410

Qvap

MW

3.807

3.881

3.968

Tvap

K

422.2

423

423

QRX

MW

2.784

2.789

2.795

QR2 compressor

MW kW

2.658 257.2

2.806 258.2

2.986 259.1

yRXin

mol % O2

2.66

2.63

2.59

mol % C2H4

63.81

63.41

63.02

mol % HAc

29.66

30.26

30.89

xB2

mol % H2O

6.03

5.29

4.69

profit

106 $/y

14.761

14.776

14.772

capital

106 $

6.0577

6.0601

6.0622

a

Conditions: 6.5 atm; Ntubes = 1100; Tcool = 455 K; Tmax = 455 K.

well below the flammability limit of 7 mol %. However, the effects of other important design optimization variables must be considered. One of the most important is acetic acid recycle. 5.2. Effect of Acetic Acid Recycle. The kinetic relationships indicate that higher acetic acid concentrations in the reactor feed should promote the desired VAM reaction. However, more acetic acid recycle means higher energy consumption in the vaporizer. The azeotropic column energy consumption is also higher because more organic reflux is required to maintain the VAM purity of the organic phase in the decanter. In addition, since acetic acid is the heaviest component in terms of volatility, higher acetic acid concentration in the reactor raise the dewpoint

temperature, which could require either more gas recycle or higher reactor coolant temperatures. Figure 6 and Table 2 give results for a range of acetic acid recycle flow rates. The abscissa in Figure 6 is the total acetic acid (fresh feed plus recycle). The production of product VAM increases, as does the consumption of reactants (ethylene and acetic acid). Gas recycle flow rate increases to keep the 455 K peak temperature, which increases compressor work. Naturally vaporizer energy consumption increases, as does the required organic reflux (R2) to the azeotropic column. Profit initially increases, but hits a maximum at a total acetic acid flow rate of 385 kmol/h. Capital investment increases steadily but quite slowly as acetic acid recycle is increased. Table 2 shows a large incremental return on an incremental investment in going from 375 to 385 kmol/h. Profit increases by $15,000 per year for a small $24,000 increase in capital investment. It should be noted that the reactor inlet acetic acid composition in the modified design (30.26 mol % HAc) is very significantly different than that in the original (11 mol % HAc) because of the much larger acetic acid recycle flow rate (329.2 compared to 84.2 kmol/h). The higher acetic acid concentrations in the reactor promote the desired VAM reaction. In addition, the reactor inlet oxygen composition in the modified design (2.63 mol % O2) is significantly different than that in the original (7.4 mol % O2). This inhibits the undesired combustion reaction. Finally, the reactor size is significantly different (1100 versus 622 tubes), which means higher capital investment in the reactor vessel and catalyst. However, the valuable VAM produced is much larger (58.19 versus 49.56 kmol/h), which can justify a significant increase in capital investment and energy consumption. Note also that the vaporizer energy consumption is higher (3.88 MW in the modified design versus 1.5 MW). However, the azeotropic column reboiler duty is lower (2.806 MW in the modified versus 4.65 MW). Since the recycle gas flow rate is smaller in the modified design, the compressor work is smaller (258 versus 350 kW). 10142



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Figure 7. Effect of pressure.

Table 3. Effect of Pressurea P = 6.0 atm

P = 6.5 atm

P = 7.0 atm

product

kmol/h

57.368

57.538

57.629

gastot ethylene

kmol/h kmol/h

676.0 56.03

821.6 56.27

966.1 56.37

acetic acid

kmol/h

55.00

55.15

55.24

organic reflux

kmol/h

387

369

357

Qvap

MW

3.650

3.882

4.000

Tvap

K

423

422

420

QRX

MW

2.784

2.724

2.670

QR2

MW

2.804

2.638

2.529

compressor yRXin

kW mol % O2

233.2 3.21

266.3 2.71

294.3 2.36

mol % C2H4

63.81

64.35

65.02

mol % HAc

33.21

29.03

25.77

xB2

mol % H2O

6.03

6.19

7.09

profit

106 $/y

14.598

14.606

14.584

capital

106 $

5.5409

5.6159

5.6781

a Conditions: HAc total = 375 kmol/h; Ntubes = 1000; Tcool = 455 K; Tmax = 455 K.

5.3. Effect of Reactor Pressure. The reaction rates of both reactions depend directly on pressure, so higher pressures increase both reaction rates, as shown in the kinetics given in eq 2, which should decrease reactor size for a given conversion. The undesirable combustion reaction is assumed to have first-order dependence on oxygen while the desirable VAM reaction is assumed to have a squareroot dependence. Therefore, lower pressure should favor the desired reaction, and selectivity should improve lower pressures. Both fresh gas streams are available at high pressure, so feed gas compression is not required over the range of system pressures considered.

The results of running several pressures are shown in Figure 7 and Table 3. The other three design variables are fixed at Tcool = 433 K, Ntubes = 1000 and HAc total = 375 kmol/h. The most important effect of changing pressure is the flow rate of gas recycle, which is set to maintain a peak reactor temperature of 455 K. The middle left graph in Figure 7 shows a very large change in the Gastot required. As the pressure at the inlet of the reactor is varied from 5.5 to 7 atm, the total gas increases by a factor of 2. The reactor heat removal changes only slightly, and the reactor temperature profiles are very similar at all pressures. The higher reaction rates at higher pressure require more gas recycle so that the maximum peak temperature in the reactor is not exceeded. Naturally, the larger gas recycle increases compressor costs. Table 3 shows that vaporizer energy consumption also increases. Since the gas recycle goes through the vaporizer in order to lower vaporizer temperatures, more gas must be heated as gas recycle increases. More product is produced as pressure is increased, but profit hits a peak at 6.5 atm because of the increases in energy costs in the compressor and the vaporizer. 5.4. Effect of Reactor Coolant Temperature. The kinetic relationships indicate that lower reactor temperatures should improve selectivity because the activation energy of the desired VAM reaction is smaller than the activation energy of the undesired combustion CO2 reaction. However, lower reactor temperatures decrease conversion so bigger reactors are required. Lower reactor temperatures also require higher gas recycle flow rate to keep from forming liquid in the reactor. The optimum design involves an economic trade-off among these various effects. Running the reactor hot enough to generate steam would appear to be the best design, but several other coolant temperatures were explored. Figure 8 and Table 4 give results over a range of reactor coolant temperatures. The other variables fixed in these runs are 1100 tubes, 6.5 atm, and 385 kmol/h HAc total. In most cases the gas 10143



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Figure 8. Effect of reactant coolant temperature.

Table 4. Effect of Reactor Coolant Temperaturea Tcool = 430 K Tcool = 433 K Tcool = 435 K product gastot

kmol/h kmol/h

58.801 665.8

58.188 800.6

57.564 919.3

ethylene

kmol/h

57.19

56.74

56.31

acetic acid

kmol/h

56.33

55.76

55.19


419

390

375

Qvap

MW

3.797

3.881

3.964

Tvap

K

427

423

420

QRX

MW

2.984

2.789

2.619

QR2 compressor

MW kW

3.090 209.2

2.806 258.2

2.673 301.2

mol % O2

2.90

2.63

2.46

mol % C2H4

59.40

63.41

66.23

yRXin

a

mol % HAc

34.33

30.26

27.43

xB2

mol % H2O

3.90

5.39

6.35

profit

106 $/y

14.259

14.776

14.521

capital

106 $

5.948

6.0601

6.1553

Conditions: HAc total = 385 kmol/h; Ntubes = 1100; Tmax = 455 K; 6.5 atm.

recycle is set to maintain a 455 K peak reactor temperature, but at low coolant temperatures the gas recycle must be increased to keep the temperatures in the reactor above the dewpoint temperature. The profit points shown in the left bottom graph in Figure 8 do not include a credit for steam generated in the reactor for coolant temperatures below 433 K. Gas recycle is set by the reactor peak temperature limitation for coolant temperature down to 427 K. Below this coolant temperature, the dewpoint limitation comes into effect, which requires a higher gas recycle flow rate (second left graph in Figure 8) and more compressor work (third right graph in Figure 8).

Capital investment increases with reactor coolant temperature, but profit takes a jump up at 433 K because of the reactor steam credit. The increase in profit in going from 430 to 433 K is $517,000 per year. The corresponding reactor steam credit is $684,000 per year. The incremental capital cost is $112,100, which is easily justified.

6. PLANTWIDE CONTROL Now that the modified flowsheet has been determined, its dynamic controllability is explored. Figure 9 gives the plantwide control structure developed. It is significantly different than those proposed for the original design. Since the oxygen concentrations are well below the flammability limit, the flow rate of fresh oxygen can be used as the throughput handle (sets the production rate of product). The control structure has the following loops: 1. Fresh oxygen feed is flow controlled. 2. Total acetic acid is ratioed to the flow rate of fresh oxygen. The flow rate of fresh acetic acid is manipulated to control the total of the fresh and recycle acetic acid. This configuration prevents snowballing problems in the liquid recycle loop. 3. Vaporizer level is controlled by manipulating heat input. 4. Pressure in the vaporizer (and thus the gas loop) is controlled by manipulating the flow rate of fresh feed of ethylene. 5. Reactor exit temperature is controlled by manipulating the cooling medium temperature. 6. Separator temperature is controlled by manipulating heat removal in the partial condenser. 7. Liquid level in the separator is controlled by manipulating the exiting liquid stream. 8. The acetic acid wash stream to the top of the absorber is flow controlled and its temperature is controlled by manipulating heat removal in the cooler. 10144



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Figure 9. Plantwide control structure.

9. The liquid level in the base of the absorber is controlled by manipulating the exiting liquid stream. 10. The MEA solvent to the top of the CO2 absorber is flow controlled. 11. The liquid level in the base of the CO2 absorber is controlled by manipulating the exiting liquid stream. 12. The ethane concentration in the purge stream is controlled by manipulating the flow rate of the purge. 13. The work of the recycle compressor is held constant. 14. The pressure in the D/E column is controlled by manipulating the vapor distillate product from the reflux drum. 15. The reflux flow rate is flow controlled. 16. Reflux drum level is controlled by manipulating condenser heat removal. 17. Stage-8 temperature is controlled by manipulating reboiler heat addition. 18. Base level is controlled by manipulating the flow rate of the bottoms. 19. The pressure in the azeotropic column is controlled by the valve in the overhead line. 20. Organic reflux is ratioed to column feed in the azeotropic column.

21. Base level is controlled by manipulating bottoms flow rate. 22. Stage 7 temperature is controlled by manipulating reboiler duty. 23. Decanter temperature is controlled by manipulating heat removal in the condenser. 24. Decanter pressure is controlled by manipulating the flow rate of the vent stream. 25. Interface level is controlled by manipulating the flow rate of the aqueous product stream. 26. Organic level is controlled by manipulating the flow rate of the VAM product stream. The dynamic flowsheet was converged to a steady state in Aspen Dynamics and gave results quite similar to the steady-state Aspen Plus results, with one exception. The major difference is the flow rate of the purge stream required to prevent the buildup of inert ethane. The purge flow rate is increased from 0.143 to 1.055 kmol/h to maintain an ethane concentration in the recycle gas of 3.89 mol %. The reason for this difference is thought to be the result of the convergence inaccuracies in steady-state simulations of individual units where the very minute concentrations of ethane can lead to incorrect component balances. The higher purge flow rate gives a greater loss of ethylene, so the VAM produced is reduced from 58.19 10145



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Figure 10. Throughput changes.

to 57.73 kmol/h, which is still significantly greater than the original 49.56 kmol/h. The performance of the control structure was tested for throughput and feed composition disturbances. Figure 10 gives results for 15% step disturbances in the set point of the oxygen flow controller. Solid lines are for an increase and dashed lines are for a decrease. The disturbance is handled well with VAM product purity (third right graph in Figure 10A) held very close to the desired purity. The other two fresh feed streams eventually change in the appropriate direction (ethylene brought in to hold system pressure and acetic acid

brought in to hold the total acetic acid flow to the reactor, which is ratioed to the oxygen flow rate). Figure 10B shows that there are some dynamic changes in the purge flow rate as the composition controller attempts to maintain the ethane composition in the recycle gas, but the final steady-state purge flow rate changes only slightly from the initial steady-state value. Figure 11 gives result when the ethane impurity in the fresh ethylene feed undergoes a drastic change from 0.1 to 1 mol %. The purge flow rate must increase to remove this 10-fold increase 10146



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Figure 11. Ethane disturbances.

in inert entering the system. A proportional-only controller (KC = 5) is used for this composition, so the ethane composition increases to about 6.5 mol % and the purge flow rate increases to almost 10 kmol/h. The increased loss of ethylene results in a reduction the product formed and the fresh acetic acid fed. The plantwide control structure provides effective regulatory control of this complex process.

(5) Douglas, J. M. Conceptual Design of Chemical Processes; McGraw-Hill: New York, 1988. (6) Turton, R., Bailie, R. C., Whiting, W. B., Shaelwitz, J. A. Analysis, Synthesis and Design of Chemical Processes, 2nd ed.; Prentice Hall: Upper Saddle River, NJ, 2003.

7. CONCLUSION A modified vinyl acetate flowsheet has been developed that looks very favorable from a steady-state economic point of view. Selectivity is significantly improved by lowering oxygen concentrations and increasing acetic acid concentrations in the reactor. An effective plantwide control structure is developed that is significantly different than that proposed for the original process and provides stable regulatory control in the face of large disturbances. ’ AUTHOR INFORMATION Corresponding Author

*E-mail: [email protected]. Tel.: 610-758-4256. Fax: 610-758-5057.

’ REFERENCES (1) Luyben, M. L.; Tyreus, B. D. An industrial design/control study for the vinyl acetate monomer process. Comput. Chem. Eng. 1998, 22, 867–877. (2) Luyben, W. L., Tyreus, B. D., Luyben, M. L. Plantwide Process Control; McGraw-Hill: New York, 1999. (3) Chen, R.; Daves, K; McAvoy, T. J. A nonlinear dynamic model of a vinyl acetate process. Ind. Eng. Chem. Res. 2003, 42, 4478–4487. (4) Olsen, D, G; Svrcek, W. Y.; Young, B. R. Plantwide control study of a vinyl acetate monomer process design. Chem. Eng. Commun. 2005, 192, 1243–1257. 10147


Design and Control of a Modified Vinyl Acetate Monomer Process

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