vinyl acetate

Published: August 07, 2011

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

ARTICLE

pubs.acs.org/IECR

Design and Control of a Modified Vinyl Acetate Monomer ProcessWilliam 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 themany designoptimization variables and several important constraints. The process features a cooled tubular reactor, both gas and liquid recyclestreams, two absorbers, and two distillation columns, one involving heterogeneous azeotropic distillation. The major designvariables are reactor temperature, reactor size, reactor pressure, ethylene gas recycle and acetic acid liquid recycle. The majorconstraints are a maximum oxygen flammability limit, a maximum internal reactor temperature limit, and the need to stay abovedewpoint temperatures in the reactor. This paper develops an economic optimum flowsheet based on the conditions and parametersprovided by M. L. Luyben and B. D. Tyreus over a decade ago as a challenge problem for the academic community for design andcontrol studies (Comput. Chem. Eng. 1998, 22, 867�877). The conceptual design developed produces 18% more VAM product forthe same oxygen fresh feed flow rate used in the original design. Compared to the original nonoptimized design, the modified designhas a reactor that is twice as large with reactor inlet conditions showing lower oxygen concentrations and higher acetic acidconcentrations. Operating pressure is lower, liquid acetic acid recycle is larger, and gas ethylene recycle is smaller. A plantwidecontrol scheme that provides effective disturbance rejection and is significantly different than the structure used in the original designis developed.

1. INTRODUCTION

The industrial challenge problem presented by Luyben andTyreus1 has received only limited attention by the academiccommunity despite its rich mixture of important steady-statedesign and dynamic control features and issues. Three plant-wide control studies were reported. All of these control studiesare based on the original flowsheet presented in the challengeproblem. Luyben et al2 demonstrated that effective plantwidecontrol could be achieved using a conventional PI controlstructure. Chen et al3 developed a nonlinear dynamic model inMatlab 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, 0.037m in dia-meter 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 afresh 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 theeconomic optimum equipment sizes and operating conditions,which are significantly different than the original design. Thebasic unit operations are the same as in the original design, butequipment 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 feedflow rate of oxygen. A bigger reactor is required, but thedominant economic value of the product should easily justifythe increase in capital investment.

There are three very important and interacting constraints inthis process. The vital safety constraint is to keep the oxygenconcentration below the flammability limit. The largest oxygenconcentration occurs at the reactor inlet. In the original design,this concentration was reported as 7.4 mol % oxygen. In the

proposed flowsheet, the oxygen concentration is only 2.6 mol %,far below the maximum safety limit. The lower oxygen concen-trations improve the selectivity since the combustion reaction islowered.

The second constraint is a maximum reactor temperature of455 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 reactortubes and fixed reactor size, this constraint can be maintained byadjusting the flow rate of gas recycle, which is mostly ethylene.The higher the gas recycle, the lower the peak temperature. Ofcourse, higher gas recycle means higher gas compression costs interms of both capital and energy.

The third important constraint is the need to guarantee thatthere is no liquid in the reactor since the reactions occur in the gasphase. Liquid can form if the temperatures in the reactor are toolow, if the pressure is too high, or if the concentration of the high-boiling acetic acid is too high.

Low reactor temperatures improve selectivity (more VAMand less CO2) because the activation energy of the VAM reactionis lower than that of the combustion reaction, as discuss in thenext section. Reactor temperatures are affected by the coolanttemperature and the gas recycle. At low coolant temperatures, itis sometimes necessary to increase gas recycle beyond thatrequired to limit the peak reactor temperature in order to keepabove dewpoint temperatures.

Although selectivity is reduced by running at higher reactortemperatures, there is an important advantage of being able torecover the exothermic heats of reaction through generation of

Received: May 25, 2011Accepted: August 6, 2011Revised: July 25, 2011

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10137 dx.doi.org/10.1021/ie201131m |Ind. Eng. Chem. Res. 2011, 50, 10136–10147

Industrial & Engineering Chemistry Research ARTICLE

steam in the reactor, which is an energy credit. If the coolingmedium temperature on the shell side of the cooled reactorcan 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 andcan be achieved by increasing acetic acid recycle. However, moregas recycle may be required to prevent the formation of liquid inthe reactor. In addition, energy consumption in the acetic acidvaporizer and in the azeotropic distillation column increase.

All of these issues must be considered in the conceptual de-sign of this very interesting and complex process. The modifieddesign has 1100 tubes instead of 622. The acetic acid recycle ismuch larger, which results in a reactor inlet concentration of30.28mol % acetic acid compared to the original design of 11mol%. The operating pressure is reduced from 8.7 to 6.5 atm. Thereboiler duty of the azeotropic column is reduced from 4.65 to3.47 MW.

The original design had a high concentration of the chemicallyinert ethane (21.6 mol % C2H6) in the gas recycle loop, whichwas assumed to build up. The ethane concentration in themodified design based on steady-state Aspen Plus simulation isfound to be very small. There are only 0.05676 kmol/h of ethaneentering in the fresh ethylene feed (99.9 mol % C2H4). Most ofthis leaves as small impurities in the organic product stream(0.02639 kmol/h C2H4) and in the aqueous stream (0.02688kmol/h C2H4) with trace amounts in the purge and decantervent streams. In the dynamic simulation discussed later in thispaper, a much larger purge flow rate is required to prevent the

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

2. PROCESS DESCRIPTION

Figure 1 shows the modified design. In all the cases presentedin 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 conditionsshown in Figure 1 are the economic optimum as developed in alater section in this paper.

Fresh acetic acid is combined with an acetic acid recyclestream and fed into a vaporizer along with a gas stream, which isthe fresh ethylene feed plus a gas recycle stream. The vaporizeroperates at 6.6 atm and 423 K, which requires medium-pressuresteam (11 bar, 457 K, $8.22 per GJ). Vaporizer energy consump-tion is 3.88 MW.

The vapor stream is combined with fresh oxygen and entersthe cooled tubular reactor at 421 K with a concentration of 2.63mol % oxygen and 30.28 mol % acetic acid, the remainder beingmostly ethylene. The reactor contains 1100 tubes, 0.037 m indiameter and 10 m in length. The coolant temperature is 433 K,and the heat removal in the reactor is 2.79 MW. The overall heat-transfer coefficient is that used in the original design (0.17 kWm�2 K�1). Figure 2 gives temperature and composition profilesin the reactor. The maximum temperature is 455 K with thecoolant temperature of 433 K and a gas recycle flow rate of 744kmol/h. Catalyst properties are taken from the Luyben andTyreus paper (0.8 void volume, 1935 kg/m3 particle density).

Figure 1. Modified flowsheet.

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The per-pass conversion of ethylene is 7.3%, of acetic acidis 15%,and of oxygen is 97.8%. The selectivity is 23 (ratio of VAMformed in the reactor to CO2 formed in the reactor). It should benoted that the selectivity of the original process is only 11, so a verysignificant 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 gasstream is fed to an absorber in which VAM is recovered by anacetic acid wash stream. The gas from the absorber goes to asecond absorber that removes carbon dioxide using a mono-ethanol 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 columnwhose function is to remove the small

amount of ethylene contained in these streams. The de-ethanizercolumn (D/E) operates at 27.2 atm. The reflux-drum temperatureis 257 K, so refrigeration is required in the condenser. The easyseparation between ethylene and VAM is achieved using a 21-stagecolumn and a reflux ratio of 1.

The bottoms from the D/E column is fed to the seconddistillation columnwhose function is to recover the acetic acid forrecycle back to the vaporizer. A 21-stage column is assumed,which is the same as that used in the original design. The ternarysystem of VAM, water, and acetic acid is highly nonlinear with aheterogeneous azeotrope formed. The overhead vapor from thecolumn is cooled to 310 K and forms two liquid phases in adecanter. The aqueous phase is 99.45 mol % water. The organicphase is 94.41 mol % VAM. A small vent stream leaves the top of

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

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the decanter. Reflux to the top of the azeotropic column is aportion of the organic phase from the decanter. This reflux isadjusted to maintain the VAM composition of the organic phasein the decanter and varies as acetic acid recycle, reactor size, andsystem pressure are varied.

3. KINETICS AND PHASE EQUILIBRIUM

The kinetic equations given by Luyben and Tyreus are noteasily implemented in Aspen simulations because they are not inthe convenient power-law form and would require writing andcompiling of a user kinetic subroutine. As a practical engineeringalternative, the approach taken was to fit the performance results(yield of VAM and CO2) for the reactor given in the originaldesign by selecting pre-exponential factors and activation en-ergies for the two reactions.

C2H4 þ CH3COOH þ 12O2 f CHOCOCH3 þ H2O

C2H4 þ 3O2 f 2CO2 þ 2H2O

ð1ÞThe physical dimensions of the reactor and catalyst properties

were fixed at those given in the original design (622 tubes, 0.037 min diameter and 10 m in length filled with 2590 kg of catalyst). Thecoolant temperature (406 K) and overall heat-transfer coefficient(U = 0.17 kW m�2 K�1) were those used in the original design.

The reactor inlet conditions were set as those given in the originaldesign in terms of inlet temperature, pressure and all molar flowrates. The ratio of the activation energies of the two competing gas-phase irreversible reaction was set at 2.5, which is close to the ratiogiven in the original design (E1 = 8000 kJ/kmol and E2 = 20000 kJ/kmol). Then the values of the two pre-exponential factors (R1 andR2) were varied by trial and error until the production rates of VAMand CO2 in the reactor were close to those reported in the originaldesign. The kinetic relationships used in this paper are

R 1 ¼ R1e�E1=RTpEpHAcðpOÞ0:5 ¼ 5:8� 10�18e�8000=RTpEpHAcðpOÞ0:5R 2 ¼ R2e�E2=RTpEpO ¼ 6:7� 10�15e�20:000=RTpEpO

ð2Þ

Overall reaction rates R n have units of kmol s�1 kg�1. Partialpressures of components pj have units of Pascals. Note that thecombustion reaction is assumed to be first-order in both oxygen andethylene partial pressures.

These simplified kinetics are compared in the followingparagraph with those given in the original paper that show morecomplex composition dependence.

R 1 ¼ 0:1036e�3674=T pEpHAcpOð1 þ 1:7pWÞ½1 þ 0:583pOð1 þ 1:7pWÞ�½1 þ 6:8pHAc�

" #

R 2 ¼ 1:9365� 105e�10116=T pOð1 þ 0:68pWÞ1 þ 0:76pOð1 þ 068pWÞ

" #

ð3ÞOverall reaction rates R n have units of mol min�1 g�1 in eq 3.Partial pressures of components pj have units of psia. Usingthe reactor inlet compositions, coolant temperature and re-actor pressure in the original paper, the two reactions rates canbe calculated: R 1 = 3.203 � 10�4 and R 2 = 8.83 � 10�6 molmin�1 g�1. These can be compared with those calculated usingthe reactor conditions in the modified design shown in Figure 1:R 1 = 3.089� 10�4 andR 2 = 0.152� 10�6 mol min�1 g�1. It isclear that the reactor conditions in the modified design are morefavorable for the production of desired VAM. This demonstratesthat the simplified kinetics used in this study have been successfulin developing a process with a much higher selectivity, asconfirmed by the calculations using the “rigorous kinetics”.

The phase equilibrium in the VAM process is quite nonlinearwith a heterogeneous azeotrope formed in the ternary system ofVAM, acetic acid, and water. The Aspen NRTL phase equilib-rium parameters are used in the simulations of the vaporizer,reactor, absorber, de-ethanizer, and decanter. The UNIF-HOC physical property package is used in the azeotropic columnto account for the association of acetic acid. The valid phases inthis column are VLL. The separation in the first D/E columnbetween ethylene and VAM is fairly easy, as shown in the T�xydiagram given in Figure 3 at the operating pressure of 27.2 atm.

A ternary diagram is given in Figure 4 for the principalcomponents in the azeotropic column: VAM, acetic acid, and

Figure 3. T�xy diagram for ethylene/VAM at 27.2 atm.

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water. The pressure is 0.33 atm, which gives a heterogeneousazeotropic temperature of 312 K that is close to the decantertemperature. The bottoms from the column contains 5 mol %water. The composition of the overhead vapor from the columnis close to the azeotropic composition. The column feed is 535.3kmol/h and the organic reflux is 390 kmol/h, giving themix pointshown in Figure 4 between the feed composition and the organiccomposition. The component balance line runs through the mixpoint 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 absorberwould come from the base of a stripping column, which is notincluded in the simulation because the stripping column has littleeconomic effect on the major units of the process.

4. FLOWSHEET CONVERGENCE

It might be useful to say a few words about the issue offlowsheet convergence for this complex multiunit process. Theexistence of two recycle streams can present significant problemsin these types of processes. However, the two recycles weresuccessfully closed by employing the technique of fixing the totalflows (fresh plus recycle) and permitting the makeup streamsto vary.

For example, the total flow rate of gas to the vaporizer (freshethylene feed plus recycle gas) is held constant using an AspenPlus Flowsheet Design Specification that varied the ethylene freshfeed. A second Aspen Plus Flowsheet Design Specification variedthe acetic acid fresh feed to maintain a specified total acetic acidflow rate (fresh plus recycle) to the vaporizer. This convergencemethod worked quite well and was successful in handling a widerange of values of design variables.

5. EFFECT OF DESIGN VARIABLES

The four major design variables in the vinyl acetate monomerprocess are reactor temperatures (set by the temperature of the

coolant and the flow rate of gas recycle that maintains a 455 Kpeak temperature), the reactor size (set by the number of tubes),the pressure in the gas loop (vaporizer, reactor, separator, andabsorbers), and the flow rate of the recycle acetic acid. Aniterative exhaustive enumeration of these four variables was usedto find the optimum values of the design variables. The economicobjective function was incremental return on incremental invest-ment. For example, as more tubes are used in the reactor, capitalinvestment increases due to the cost of the vessel (more heat-transfer area) and the cost of the catalyst. Production of VAMproduct increases as more tubes are added, but the increase is atan increasingly smaller rate. The point that gives an incrementalROI of at least 30% was selected.

The three constraints that must be satisfied are a maximumpeak temperature in the reactor, an upper limit on the oxygenconcentration of the feed to the reactor (assumed to be 7 mol %oxygen), and the requirement that thematerial in the reactor is allgas phase (temperature must be above the dewpoint temperatureof the mixture at all points in the reactor). Keep in mind that theflow rate of fresh oxygen is fixed for all cases.5.1. Effect of Reactor Size.The catalyst in the VAMprocess is

quite expensive because it contains palladium. A catalyst cost of$800 per kg is assumed (2.27 wt % palladium at $800 oz), whichfavors 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 inthe plant as low-pressure steam. Operating with lower coolanttemperature is explored in a later section. The peak reactortemperature is maintained at 455 K by varying gas recycle flowrate using a Flowsheet Design Specification. The acetic acid recycleis initially set at 375 kmol/h. Reactor inlet pressure is initially setat 6.5 atm.With these design variables fixed, the effect of changing the

number of reactor tubes is shown in Figure 5. The production ofVAM and the consumption of reactants increase as more tubesare used in the reactor, but these flow rates begin to level out for

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

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more than about 1200 tubes. Since the product is worth morethan the reactants, profit increases. But capital cost of the reactorvessel 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 perkmol, ethylene = $0.442 per kmol and acetic acid = $0.596 perkmol. The cost of oxygen is not considered because the freshoxygen flow rate is the same in all cases. A “pseudo-profit”parameter is defined as the value of the VAM producedminus thecost of the ethylene and acetic acid reactants minus the energycosts of the compressor, vaporizer, and azeotropic columnreboiler plus the value of the reactor steam:

profitð$=hÞ ¼ ðproduce kmol=hÞð$0:971Þ� ðfresh ethylene kmol=hÞð$0:442Þ� ðfresh HAc kmol=hÞð$0:596Þ� ðcompressor work kWÞ 3600 s=h

106 kW=GW

� �ð$16:8=GJÞ

� ðvaporizer energy MWÞ 3600 s=h103 MW=GW

� �ð$8:22=GJÞ

� ðazeo column reboiler MWÞ 3600 s=h103 MW=GW

� �ð$7:78=GJÞ

þ ðreactor steam MWÞ 3600 s=h103 MW=GW

� �ð$7:88=GJÞ ð4Þ

The energy costs in the de-ethanizer (reboiler heat input andcondenser refrigeration) and in the absorbers remain essentiallythe same as conditions in the reactor are changed, so they are notconsidered.Capital investment is the sum of the installed costs of the

reactor vessel, catalyst, vaporizer, and compressor. The capitalcosts of all the other units remain essentially the same asconditions in the reactor change. They are not considered inthe economics since an incremental return on investment is used

to select the optimum values of design variables. Cost equationsfrom Douglas5 and Turton et al6 are used to calculate thesecapital costs.Figure 5 and Table 1 show that profit is increasing but capital

investment is also increasing as more reactor tubes are used. Theincremental return on investment decreases. Moving from 1000to 1100 tubes gives an incremental return on investment of 35%.Adding more tubes yields a lower incremental return on invest-ment. Note that the required total gas recycle is set to achievea 455 K peak reactor temperature. Under these conditions, witha 433 K reactor coolant temperature, the dewpoint limitationis avoided (the process stream leaving the reactor has a vaporfraction >1). Reactor inlet oxygen concentration (2.66 mol %) is

Figure 5. Effect of reactor size.

Table 1. Effect of Reactor Sizea

Ntubes = 1000 Ntubes = 1100 Ntubes = 1200

product kmol/h 57.538 57.983 58.202

gastot kmol/h 821.6 796.9 786.3

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 MW 2.638 2.658 2.668

compressor kW 266.3 257.2 253.3

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 6.5096aConditions: 6.5 atm;HAc total = 375 kmol/h;Tcool = 455K;Tmax = 455K.

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well below the flammability limit of 7mol %. However, the effectsof other important design optimization variables must be con-sidered. 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 feedshould promote the desired VAM reaction. However, moreacetic acid recycle means higher energy consumption in thevaporizer. The azeotropic column energy consumption is alsohigher because more organic reflux is required to maintain theVAM 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 orhigher 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 VAMincreases, as does the consumption of reactants (ethylene andacetic acid). Gas recycle flow rate increases to keep the 455 Kpeak temperature, which increases compressor work. Naturallyvaporizer energy consumption increases, as does the requiredorganic 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 increasessteadily but quite slowly as acetic acid recycle is increased. Table 2shows a large incremental return on an incremental investment ingoing from 375 to 385 kmol/h. Profit increases by $15,000 peryear 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 significantlydifferent than that in the original (11 mol % HAc) because ofthe much larger acetic acid recycle flow rate (329.2 compared to84.2 kmol/h). The higher acetic acid concentrations in the reactorpromote 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 theoriginal (7.4 mol % O2). This inhibits the undesired combustionreaction. Finally, the reactor size is significantly different (1100versus 622 tubes), which means higher capital investment in thereactor vessel and catalyst. However, the valuable VAM produced ismuch larger (58.19 versus 49.56 kmol/h), which can justify asignificant 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 azeo-tropic column reboiler duty is lower (2.806 MW in the modi-fied versus 4.65 MW). Since the recycle gas flow rate is smaller inthe modified design, the compressor work is smaller (258 versus350 kW).

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 kmol/h 57.983 58.188 58.388

gastot kmol/h 796.9 800.6 804.3

ethylene kmol/h 56.61 56.74 56.88



Qvap MW 3.807 3.881 3.968

Tvap K 422.2 423 423

QRX MW 2.784 2.789 2.795

QR2 MW 2.658 2.806 2.986

compressor kW 257.2 258.2 259.1

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 6.0622aConditions: 6.5 atm; Ntubes = 1100; Tcool = 455 K; Tmax = 455 K.

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5.3. Effect of Reactor Pressure. The reaction rates of bothreactions depend directly on pressure, so higher pressures increaseboth reaction rates, as shown in the kinetics given in eq 2, whichshould decrease reactor size for a given conversion. The undesirablecombustion reaction is assumed to have first-order dependence onoxygenwhile the desirableVAMreaction is assumed to have a square-root dependence. Therefore, lower pressure should favor the desiredreaction, and selectivity should improve lower pressures. Both freshgas streams are available at high pressure, so feed gas compression isnot required over the range of system pressures considered.

The results of running several pressures are shown in Figure 7and Table 3. The other three design variables are fixed at Tcool =433 K, Ntubes = 1000 and HAc total = 375 kmol/h. The mostimportant effect of changing pressure is the flow rate of gasrecycle, which is set to maintain a peak reactor temperature of455 K. The middle left graph in Figure 7 shows a very largechange in the Gastot required. As the pressure at the inlet of thereactor is varied from 5.5 to 7 atm, the total gas increases by afactor of 2.The reactor heat removal changes only slightly, and the reactor

temperature profiles are very similar at all pressures. The higherreaction rates at higher pressure require more gas recycle so thatthe maximum peak temperature in the reactor is not exceeded.Naturally, the largergas recycle increases compressor costs.Table3

shows that vaporizer energy consumption also increases. Since thegas recycle goes through the vaporizer in order to lower vaporizertemperatures, 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 inthe compressor and the vaporizer.5.4. Effect of Reactor Coolant Temperature. The kinetic

relationships indicate that lower reactor temperatures shouldimprove selectivity because the activation energy of the desiredVAM reaction is smaller than the activation energy of theundesired combustion CO2 reaction. However, lower reactortemperatures decrease conversion so bigger reactors are re-quired. Lower reactor temperatures also require higher gasrecycle flow rate to keep from forming liquid in the reactor.The optimum design involves an economic trade-off amongthese various effects. Running the reactor hot enough to generatesteam would appear to be the best design, but several othercoolant 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 1100tubes, 6.5 atm, and 385 kmol/h HAc total. In most cases the gas

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 kmol/h 676.0 821.6 966.1

ethylene kmol/h 56.03 56.27 56.37



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 kW 233.2 266.3 294.3

yRXin mol % O2 3.21 2.71 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 5.6781aConditions: HAc total = 375 kmol/h; Ntubes = 1000; Tcool = 455 K;Tmax = 455 K.

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recycle is set to maintain a 455 K peak reactor temperature, but atlow coolant temperatures the gas recycle must be increased to keepthe 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 coolanttemperatures below 433 K. Gas recycle is set by the reactor peaktemperature limitation for coolant temperature down to 427 K.Below this coolant temperature, the dewpoint limitation comesinto effect, which requires a higher gas recycle flow rate (secondleft graph in Figure 8) and more compressor work (third rightgraph in Figure 8).

Capital investment increases with reactor coolant tempera-ture, but profit takes a jump up at 433 K because of the reactorsteam 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, itsdynamic controllability is explored. Figure 9 gives the plantwidecontrol structure developed. It is significantly different than thoseproposed for the original design. Since the oxygen concentra-tions are well below the flammability limit, the flow rate of freshoxygen can be used as the throughput handle (sets the produc-tion 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 thetotal of the fresh and recycle acetic acid. This configurationprevents 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 ofethylene.

5. Reactor exit temperature is controlled by manipulating thecooling medium temperature.

6. Separator temperature is controlled by manipulating heatremoval in the partial condenser.

7. Liquid level in the separator is controlled by manipulatingthe exiting liquid stream.

8. The acetic acid wash stream to the top of the absorber isflow controlled and its temperature is controlled by manip-ulating heat removal in the cooler.

Figure 8. Effect of reactant coolant temperature.

Table 4. Effect of Reactor Coolant Temperaturea

Tcool = 430 K Tcool = 433 K Tcool = 435 K

product kmol/h 58.801 58.188 57.564

gastot kmol/h 665.8 800.6 919.3

ethylene kmol/h 57.19 56.74 56.31



Qvap MW 3.797 3.881 3.964

Tvap K 427 423 420

QRX MW 2.984 2.789 2.619

QR2 MW 3.090 2.806 2.673

compressor kW 209.2 258.2 301.2

yRXin mol % O2 2.90 2.63 2.46

mol % C2H4 59.40 63.41 66.23

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 6.1553aConditions:HAc total = 385 kmol/h;Ntubes = 1100;Tmax = 455K; 6.5 atm.

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9. The liquid level in the base of the absorber is controlled bymanipulating the exiting liquid stream.

10. The MEA solvent to the top of the CO2 absorber is flowcontrolled.

11. The liquid level in the base of the CO2 absorber iscontrolled by manipulating the exiting liquid stream.

12. The ethane concentration in the purge stream is con-trolled 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 thereflux drum.

15. The reflux flow rate is flow controlled.16. Reflux drum level is controlled by manipulating conden-

ser heat removal.17. Stage-8 temperature is controlled by manipulating re-

boiler 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 AspenPlus results, with one exception. Themajor difference is theflow rateof 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 tomaintain an ethane concentration in the recycle gas of 3.89 mol %.The reason for this difference is thought to be the result of theconvergence inaccuracies in steady-state simulations of individualunits where the very minute concentrations of ethane can lead toincorrect component balances. The higher purge flow rate gives agreater loss of ethylene, so theVAMproduced is reduced from58.19

Figure 9. Plantwide control structure.

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to 57.73 kmol/h, which is still significantly greater than the original49.56 kmol/h.

The performance of the control structure was tested for through-put and feed composition disturbances. Figure 10 gives results for15% step disturbances in the set point of the oxygen flow controller.Solid lines are for an increase anddashed lines are for a decrease.Thedisturbance is handled well with VAM product purity (third rightgraph in Figure 10A) held very close to the desired purity. The othertwo fresh feed streams eventually change in the appropriate direc-tion (ethylene brought in to hold system pressure and acetic acid

brought in to hold the total acetic acid flow to the reactor, which isratioed to the oxygen flow rate).

Figure 10B shows that there are some dynamic changes in thepurge flow rate as the composition controller attempts to main-tain the ethane composition in the recycle gas, but the finalsteady-state purge flow rate changes only slightly from the initialsteady-state value.

Figure 11 gives result when the ethane impurity in the freshethylene feed undergoes a drastic change from 0.1 to 1 mol %.The purge flow rate must increase to remove this 10-fold increase

Figure 10. Throughput changes.

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

7. CONCLUSION

A modified vinyl acetate flowsheet has been developedthat looks very favorable from a steady-state economic point ofview. Selectivity is significantly improved by lowering oxygenconcentrations and increasing acetic acid concentrations in thereactor.

An effective plantwide control structure is developed that issignificantly different than that proposed for the original processand provides stable regulatory control in the face of large dis-turbances.

’AUTHOR INFORMATION

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

’REFERENCES

(1) Luyben, M. L.; Tyreus, B. D. An industrial design/control studyfor 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.

(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: UpperSaddle River, NJ, 2003.

Figure 11. Ethane disturbances.

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vinyl acetate

Documents

original nonoptimized

reactor pressure

original paper

original owsheet

reactor size

conceptual design

reactor inlet conditions

original designis