Design and Control of Conventional and Reactive Distillation

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PROCESS DESIGN AND CONTROL Design and Control of Conventional and Reactive Distillation Processes for the Production of Butyl Acetate William L. Luyben,* Kristin M. Pszalgowski, Michael R. Schaefer, and Cory Siddons Process Modeling and Control Center, Department of Chemical Engineering, Lehigh University, Bethlehem, Pennsylvania 18015

Jimenez and Costa-Lopez (Ind. Eng. Chem. Res. 2002, 41, 6735) recently discussed a reactive distillation process for the production of butyl acetate in which an extractive agent is used. Their process requires four distillation columns. No details of their plantwide control structure are given. This paper presents a quantitative comparison of two alternative processes for the production of butyl acetate in which no extractive agent is employed. The first is a conventional reactor/separation process with a reactor, three distillation columns, and two recycles. The second uses a reactive distillation column with two other conventional columns and two recycles. Each process is economically optimized, and the reactive distillation process has a 20% lower total annual cost. Plantwide control structures are developed for each process, and their effectiveness in the face of very large disturbances is demonstrated by dynamic simulation. 1. Introduction Reactive distillation has been the subject of many papers in recent years because of its potential for process intensification and improved economics. Applications of reactive distillation are limited to systems in which the reaction rates are fairly high and there is no mismatch of temperatures favorable for reaction and separation as these two operations occur simultaneously in a reactive distillation column. A recent paper by Jimenez and Costa-Lopez1 reported work on the production of butyl acetate and methanol by reacting methyl acetate and butanol in a reactive distillation column. The feed stream is a mixture of methyl acetate and methanol, which comes from an upstream polyvinyl alcohol process. They propose reacting the methyl acetate with butanol to form butyl acetate and methanol. This liquid-phase reaction is reversible with an equilibrium constant of about unity. Kinetics are quite fast, and heat effects are small. In their first paper,2 Jimenez and Costa-Lopez provide useful information dealing with reaction equilibrium and kinetics, which are used in this paper. In their second paper,1 they study a four-column process in which the key unit is a reactive distillation column. The presence of a minimum-boiling azeotrope between methyl acetate and methanol prevents the use of just a single reactive column because any unreacted methyl acetate that leaves the reaction zone will go out the top of the column with the methanol. Jimenez and Costa-Lopez1 claim that an extractive agent (o-xylene) must be added to the reactive distillation column. Of course, this means that the entrainer must recovered in a subsequent column. In addition, the unpreventable contamination of products with small amounts of the entrainer occurs. * To whom correspondence should be addressed. E-mail: [email protected]. Tel.: 610-758-4256.

Gangadwala et al.3 studied a reactive distillation process for the production of butyl acetate from the reaction of butanol with acetic acid. In this paper, we demonstrate that specification products can be economically produced without the use of an extraction agent, which simplifies the process. Two alternative processes are explored: (1) a conventional process with a reactor and three distillation columns, in which reactants leaving the reactor are recovered and recycled and methanol and butyl acetate products are produced at 99% purity, and (2) a process with a reactive distillation column and two conventional columns. In the latter process, the overhead of the reactive column is a methyl acetate/methanol mixture that is separated in a downstream column into a methanol bottoms product and a distillate with a composition close to the binary azeotrope (66.7 mol % methyl acetate at 15 psia), which is recycled back to the reactive distillation column. The bottoms of the reactive column is separated in another column into a butyl acetate bottoms product and a distillate that is about 90 mol % butanol, which is recycled back to the reactive column. Fresh butanol feed and the methyl acetate/methanol fresh feed streams are also fed to the reactive column. The steady-state economic optimum designs of these two processes are presented, and plantwide dynamic control structures are developed and tested. The commercial simulation products AspenPlus and AspenDynamics are used in this study. 2. Chemical Kinetics and Phase Equilibrium 2.1. Chemical Kinetics and Chemical Equilibrium. The liquid-phase reversible reaction considered is

MeAc + BuOH T MeOH + BuAc The kinetics for the forward and reverse reactions are

10.1021/ie040167r CCC: $27.50 © 2004 American Chemical Society Published on Web 11/10/2004

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Figure 1. Methyl acetate/methanol vapor-liquid equilibrium at 15 psia.

based on those given by Jimenez and Costa-Lopez.1 Their reaction rates are given in units of moles per minute per gram of catalyst (mol min-1 gcat-1) and must be converted to units acceptable to AspenPlus, which are kilomoles per second per cubic meter of reactor volume (kmol s-1 m-3). Assuming a catalyst bulk density of 2000 kg/m3, the kinetic expressions used are given below, with the overall reaction rate R having a first-order dependence on each of the two reactants with concentration in molality (kmol/m3)

R ) kFCMeACCBuOH - kRCMeOHCBuAC kF ) 7 × 106e-71960/RT kR ) 9.467 × 106e-72670/RT where the activation energies are in kilojoules per mole and temperature is in kelvin. Notice that the activation energies are essentially the same, which means that the heat of reaction is small and also that the equilibrium constant has little temperature dependence. The temperature in the reactor of the conventional process is set at 200 °F, and this is also about what the temperatures are in the reactive zone of the reactive distillation column when operating at 15 psia in the condenser. 2.2. Vapor-Liquid Equilibrium. The phase equilibrium of this four-component system is complex because of the existence of two binary azeotropes. Using NRTL physical properties (as recommended by Jimenez and Costa-Lopez1), AspenPlus predicts two binary azeotropes: (1) Methyl acetate and methanol form a homogeneous minimum-boiling azeotrope with a composition of 66.7 mol % methyl acetate at 15 psia and 129.5 °F. (2) Butanol and butyl acetate form a homogeneous minimum-boiling azeotrope with a composition of 78.1 mol % butanol at 14.7 psia and 242.4 °F. However, the azeotrope disappears when the pressure is raised to about 50 psia. The first azeotrope means that any columns that are separating methyl acetate from methanol can produce high-purity methanol in the bottoms, but can only produce a distillate that has a composition near the azeotropic composition. In this work, these columns are specified to produce a distillate with a composition of 64 mol % methyl acetate. Note that a column pressure

of 15 psia permits the use of cooling water in the condenser. Figure 1 gives a temperature-composition (T-x-y) diagram for methyl acetate and methanol at 15 psia. The second azeotrope has an unusual pressure dependence. In most chemical systems, increasing pressure moves the composition of the azeotrope to the left (becomes less rich in the lighter component). However, the opposite effect is displayed in the butanol/butyl acetate system, as shown in Figure 2. In this study, the columns separating butanol and butyl acetate are operated at 50 psia and specified to produce a distillate with a composition of about 90 mol % butanol. This composition is selected because of the “pinch” in the x-y curves that occurs at the high-butanol end of the diagrams. This separation is difficult and therefore requires a high reflux ratio (2-5 depending on the column feed composition) and many trays (50 are used). If this separation were conducted at low pressure, the butanol recycled back to the reactor or to the reactive distillation column would have a higher concentration of butyl acetate. This would reduce the per-pass conversion and require larger recycle flows. However, the disadvantage in operating at high pressure is a high base temperature. It would be 260 °F if the column were operated at 15 psia, but would increase to 346 °F for operation at 50 psia. This means that steam at about 400 °F must be used in the reboiler (250 psia steam), which is more expensive than the 310 °F (80 psia) steam that could be used with a lowpressure column. However, there is an additional advantage of operating a high-pressure butanol/butyl acetate column. The reflux drum temperature at 50 psia is about 316 °F, which is high enough to permit heat integration with one of the other columns in the process (the methanol column, which has a reboiler temperature of about 157 °F). Thus, the process is much more efficient with a high-pressure column making this separation. 3. Alternative Process Configurations 3.1. Reactive and Extractive Process. Before we describe in detail the two processes studied in the paper, it might be useful to describe the process proposed by Jimenez and Costa-Lopez.1 Figure 3 gives a sketch of this four-column process.

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Figure 2. Butanol/butyl acetate vapor-liquid equilibrium at (A) 15 and (B) 50 psia.

Figure 3. Reactive and extractive distillation process.

Insufficient details are provided in their paper to know the process conditions in detail, but it appears that

the o-xylene entrainer is being used to break the methyl acetate/methanol azeotrope (see their Figure 3 in their


Figure 4. Conventional process.

Part I paper2) and perhaps affect the butanol/butyl acetate separation as well. However, the composition profiles given in Figure 5 of their Part II paper1 show a significant amount of methyl acetate leaving the top of the reactive column (20 mol % methyl acetate) even with the entrainer being added to the column. Results given later in this paper show composition profiles that are quite similar without use of the entrainer. The reactive distillation column (T-200) is fed with fresh butanol, recycle butanol, recycle o-xylene, and an azeotropic mixture of methyl acetate and methanol. The last of these comes from a column (T-100) that produces a bottoms methanol product and the azeotrope distillate product. The T-100 column is fed with the fresh methyl acetate/methanol feed and with the distillate from the reactive column. All of the methanol produced comes out the bottom of T-100. The bottoms from the reactive column goes to two columns in series. The first (T-300) apparently recycles butanol and xylene from the bottom back to the reactive column. The T-300 distillate is fed to T-400 in which the butyl acetate is removed as the distillate product and the bottoms is recycled back to the reactive column. It is not clear from the discussion in their paper how this separation works as butyl acetate is higher-boiling than butanol. These authors presented only a very sketchy description of the process conditions, so it is difficult to understand how the process really works. They also provide little information about their plantwide control structure other than stating that PI controllers and ratio control are used. They also state that an on-line composition analyzer must be used. Only temperature measurements are used in this paper. The authors provide no P&I (process and instrumentation) drawing showing their control structure. There is no discussion of how the fresh feeds are managed to ensure an exact stoichiometric balance. Nor is there any discussion of how the recycle streams are controlled to prevent “snowballing.” The authors claim to have tested

their process with 5-10% disturbances. Much larger disturbances are used in this paper. 3.2. Conventional Process. Figure 4 gives a sketch of the conventional process studied in this paper. Flowsheet stream conditions, equipment sizes, and reboiler heat duties are provided. Fresh feed with a composition of 60 mol % methyl acetate and 40 mol % methanol and a flow rate of 128 lb-mol/h is fed into a 200-ft3 reactor, which operates at 200 °F and 100 psia. The residence time in the reactor is about 15 min. Fresh butanol (76.8 lb-mol/h) is also fed to the CSTR (continuous stirred tank reactor), along with a butanol recycle stream (208 lb-mol/h, 90 mol % butanol, 10 mol % butyl acetate) and a methyl acetate/methanol recycle (157 lbmol/h, 64 mol % methyl acetate, 36 mol % methanol). The per-pass conversion of methyl acetate is about 42%. A small amount of heat (0.09 × 106 Btu/h) must be added to the reactor. Reactor effluent is fed to column C1 in which methyl acetate and methanol are taken overhead and butanol and butyl acetate leave in the bottoms. The column has 30 trays. The specifications for this column are 0.01 mol % methanol in the bottoms and 2 ppm butyl acetate in the distillate. The required reflux ratio is 0.7, and the reboiler heat duty is 6.92 × 106 Btu/h. The column operates with a condenser pressure of 15 psia, which gives a reflux drum temperature of 132 °F and permits the use of cooling water in the condenser. The distillate stream is fed to column C2, which produces high-purity (99 mol %) methanol in the bottoms and a distillate stream whose composition (64 mol % methyl acetate) is near that of the azeotrope. The distillate is recycled back to the reactor at a rate of 157 lb-mol/h. The column has 30 trays. To achieve these specifications, the reflux ratio is 1.0, and the reboiler heat duty is 4.39 × 106 Btu/h. The column operates with a condenser pressure of 15 psia, which gives a reflux drum temperature of 136 °F and permits the use of cooling water in the condenser. The bottoms from C1 is fed to column C3, which produces high-purity (99 mol %) butyl acetate in the

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Figure 5. Reactive distillation process.

bottoms and a butanol-rich (90 mol %) distillate that is recycled back to the reactor at a rate of 208 lb-mol/h. The column has 50 trays. To achieve these specifications, the reflux ratio is 2.0, and the reboiler heat duty is 11.1 × 106 Btu/h. The column operates with a condenser pressure of 50 psia, which gives a reflux drum temperature of 316 °F. As noted earlier, this temperature is high enough to permit heat integration with the low-temperature reboiler (157 °F) in column C2. The total energy consumption of the three columns in this conventional process is 22 × 106 Btu/h, not accounting for any heat integration. 3.3. Reactive Distillation Process. Figure 5 gives the flowsheet for the process in which the reaction occurs in reactive distillation column C2. The fresh feed mixture of methyl acetate and methanol is fed to the first distillation column, which has 30 trays. This column is also fed by a recycle stream from the top of reactive column C2. The first column produces highpurity (99 mol %) methanol in the bottoms and a distillate stream whose composition (64 mol % methyl acetate) is near that of the azeotrope. The distillate is fed to the reactive column C2 at a rate of 137 lb-mol/h. To achieve these specifications, the reflux ratio is 0.56, and the reboiler heat duty is 3.13 × 106 Btu/h. The column operates with a condenser pressure of 15 psia, which gives a reflux drum temperature of 130 °F and permits the use of cooling water in the condenser. Reactive column C2 has two feed streams. The lighter methyl acetate/methanol stream is introduced lower in the column (stage 22 using Aspen notation of numbering from the top). The heavy butanol stream is introduced higher in the column (stage 12). The reactive zone is from stage 12 to 22. The molar holdup on the reactive trays is specified to be 1.5 lb-mol. This corresponds to a weir height of about 1.8 in. for a column that is 4.8 ft in diameter, so the hydraulics are quite reasonable. The butanol stream is the combination of the fresh butanol feed (76.9 lb-mol/h) and the butanol recycle stream from the overhead of the third column, C3. The distillate from C2 is fed back to C1. There must be a negligible amount of butanol or butyl acetate in this

distillate stream because these will all go out the bottom of C1 and reduce the methanol product purity. Only small amounts of methyl acetate or methanol should go out the bottom of C2 because they will affect the purity of the distillate recycle stream (methyl acetate should be put high up in reactive column C2 where it cannot react and will increase the recycle back to C1). Figure 6 gives the composition profiles in the reactive column C2. The butanol is fed on stage 12, and there is a high concentration of butanol in the reactive zone (stages 12-22). Despite this large butanol concentration, there is a small concentration of methyl acetate at the top of the reactive zone. This methyl acetate goes out the top of the column, building up to about 10 mol % in the top trays. The bottoms from C2 is fed to the last column, C3, which separates the butanol from the butyl acetate. High-purity (99 mol %) butyl acetate leaves in the bottoms, and a butanol-rich (96 mol %) distillate is recycled back to the reactor at a rate of 54 lb-mol/h. The column has 50 trays. To achieve these specifications, the reflux ratio is 5.6, and the reboiler heat duty is 5.96 × 106 Btu/h. The column operates with a condenser pressure of 50 psia, which gives a reflux drum temperature of 316 °F. As noted earlier, this temperature is high enough to permit heat integration with the lowtemperature reboiler (157 °F) in column C1. The total energy consumption of the three columns in this reactive distillation process is 18 × 106 Btu/h, not accounting for any heat integration. This should be compared with the 22 × 106 Btu/h required in the conventional process. In addition, the conventional process has a fourth process vessel, the reactor (5 ft in diameter and 10 ft in length). 3.4. Flowsheet Convergence. As any user of flowsheet simulators knows, the convergence of steady-state simulators when recycle streams are present can be very difficult. Such is the case with both of these processes since they each involve two recycle streams. An alternative approach is to use a dynamic simulation to converge the process flowsheet to a steady state.


Figure 6. Composition profiles in the reactive distillation column. Table 2. Basis of Economicsa,b

Table 1. Economic Comparison conventional process

parameter

reactive distillation process

C1 ID (ft) NT QC (106 Btu/h) AC (ft2) QR (106 Btu/h) AR (ft2) shell cost (106 $) HX Cost (106 $) energy cost (106 $/year)

4.2 30 7.0 2330 6.92 1384 0.270 0.412 0.303

3.5 30 2.93 977 3.21 642 0.157 0.241 0.141

3.0 30 4.34 1454 4.39 878 0.191 0.304 0.192

4.8 30 8.24 2750 8.88 1180 0.315 0.469 0.389

5.3 50 9.86 3290 11.1 2220 0.523 0.534 0.486

3.7 50 5.34 1780 5.96 1190 0.360 0.357 0.261

C2 ID (ft) NT QC (106 Btu/h) AC (ft2) QR (106 Btu/h) AR (ft2) shell cost (106 $) HX cost (106 $) energy cost (106 $/year) C3 ID (ft) NT QC (106 Btu/h) AC (ft2) QR (106 Btu/h) AR (ft2) shell cost (106 $) HX cost (106 $) energy cost (106 $/year) DxL (ft) cost (106 $)

Reactor 5 × 10 0.203 Costs

capital (106 $) energy (106 $/year) TACa (106 $/year) a

2.44 0.982 1.79

1.90 0.791 1.42

TAC ) total annual cost.

The process conditions shown in Figures 4 and 5 are obtained in this way. Luyben4 discusses this procedure in detail. 3.5. Economic Comparison. The capital and energy costs of the conventional process and the reactive distillation process are compared in Table 1. As expected, because there is no mismatch between the reaction temperatures and the distillation temperatures,

Condensers heat-transfer coefficient 150 Btu h-1 °F-1 ft-2 differential temperature 20 °F capital cost 1557(area)0.65 Reboilers heat-transfer coefficient differential temperature capital cost

100 Btu h-1 °F-1 ft-2 50 °F 1557(area)0.65

column vessel capital cost reactor vessel capital cost energy cost

1917D1.066L0.802 3(1917D1.066L0.802) $5 per 106 Btu

a TAC ) (capital cost/payback period) + energy cost period ) 3 years

b

Payback

the reactive distillation process has a 20% lower total annual cost. The economic bases for these calculations are given in Table 2. The reactive distillation process requires less capital investment than the conventional ($1,900,000 versus $2,440,000) and has a lower energy cost ($791,000 per year versus $982,000 per year). 3.6. Optimization. Rigorous comprehensive optimizations of the flowsheets discussed above were not performed during this conceptual design phase of the work. Our main objective was to determine whether an entrainer-free process was feasible. A heuristic, engineering-judgment optimization approach was employed. The numbers of trays in the columns were selected to be approximately twice the minimum number required for the separation so as to have reasonable reflux ratios. Thirty-tray columns are used for the fairly easy separations (in C1 in both processes and in C2 in the conventional process), which gave low reflux ratios. For example, increasing the number of trays in C1 in the conventional process from 30 to 40 only reduced the reflux ratio from the already low value of 0.70 to 0.55, whereas decreasing the number of trays 30 to 27 increased the reflux ratio to 1.1. A 50-tray C3 column was selected because of the more difficult separation as a result of the pinch in the x-y curve near the azeotropic composition. The resulting reflux ratio of 2 is reasonable. Exploratory runs with more or fewer trays showed little economic sensitivity. Feed tray locations were also varied but had a fairly minor effect on energy consumption.

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Figure 7. (A) Plantwide control structure and (B) controller faceplates for the conventional process.

Exploratory runs were made with the reactive distillation column in which the numbers of reactive, stripping, and rectifying stages were varied. Both the performance of the column and the economic factors were found to be fairly insensitive to these stage numbers. Having about 10 stages in each section produced reasonable conversions and recycle flow rates. The purities of the BuOH recycle streams in the two processes are additional design optimization variables. These were set at levels that seemed appropriate for the particular case. A somewhat higher purity is used in the reactive column because reaction holdup is more expensive in a column compared to a reactor vessel. The 90 mol % BuOH used in the conventional process and the 96 mol % BuOH used in the reactive distillation process were assumed to provide reasonable compromises in the tradeoff between the costs of reaction and the costs of separation. Certainly, more rigorous optimization methods could improve these designs, but the relative differences in steady-state economics and in dynamic control should not be affected greatly.

4. Plantwide Control Plantwide control schemes were developed and tested for both the conventional and the reactive distillation processes. Portions of the control structures are identical, but there are significant differences between the two control schemes, particularly regarding recycle streams and fresh feed makeup flows. Proportional-integral controllers were used for flows, pressures, and temperatures. Proportional controllers were used for all liquid levels. Dead times of 1 min were inserted in the column temperature controllers. Relayfeedback tests were run to determine ultimate gains and periods of these column temperature controllers, and the Tyreus-Luyben settings were used. The location of the tray whose temperature is to be controlled was selected by finding the location in the temperature profile in each column where there is the largest change in temperature from tray to tray. In the conventional process, these locations were stages 26, 29, and 44 in columns C1, C2, and C3, respectively. In the reactive distillation process, these locations were stages 25, 8,


Figure 8. (A) Flow disturbances and (B) feed composition disturbances for the conventional process.

and 44 in columns C1, C2, and C3, respectively. No on-line composition measurements were used. 4.1. Conventional Process. The control scheme developed for the reactor/three-column conventional process is shown in Figure 7A. The important loops from a plantwide control perspective are the two flow controllers FC1 and FC2 that control the flow rates of the total methyl acetate and total butanol streams fed to the reactor. (1) The flow rate of the stream MEACTOT is measured, and the flow rate of the fresh methyl acetate/ methanol feed is manipulated (using valve V1) to keep the total flow at its desired set point (285 lb-mol/h at base-case conditions). Note that this total flow is the

sum of the flows of the fresh feed and the distillate from column C2. This distillate flow is manipulated to control the level in the C2 reflux drum. If this level increases, the distillate flow rate will increase. Because the total flow rate is fixed, however, the flow rate of the fresh feed will be reduced by the FC1 flow controller. (2) The flow rate of the stream BUOHTOT is measured and the flow rate of the fresh butanol feed (stream FBUOH) is manipulated to keep the total flow at its desired set point (285 lb-mol/h at base-case conditions). Note that this total flow is the sum of the flows of the fresh butanol feed and the distillate from column C3. This distillate flow is manipulated to control the level in the C3 reflux drum. If this level increases, the

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Figure 9. (A) Plantwide control structure and (B) controller faceplates for the reactive distillation process.

distillate flow rate will increase. Because the total flow rate is fixed, however, the flow rate of the fresh butanol feed will be reduced by the FC2 flow controller. The set point of the FC2 controller comes from a multiplier, which means the total butanol flow rate is ratioed to the total methyl acetate flow rate. As shown on the FC2 controller faceplate in Figure 7B, this controller is “on cascade.” The production rate handle in this process is the set point of the FC1 controller. The other loops are standard unit-operations control schemes for the individual units (reactor and columns) and are displayed in Figure 7A. The effectiveness of this plantwide control structure is demonstrated by subjecting the process to feed flow rate and feed composition disturbances. Figure 8A presents results for a series of step changes in the set point of FC1. At time 0.2 h, the set point is increased from 285 to 342 lb-mol/h (a 20% increase). At time 4 h, the set point is decreased to 228 lb-mol/h (a 40% decrease). Finally, at time 8 h, the set point is increased to 342 lb-mol/h (a 40% increase).

The immediate effect of increasing the FC1 set point is an increase in both the fresh methyl acetate/methanol feed and the butanol fresh feed. However, as the flow rates increase through the process, the two recycle streams (D2 and D3) increase. As a result, the two flow controllers cut back on the two fresh feed streams. The process ends up at a new steady state, producing more methanol (B2) and more butyl acetate (B3). The process rides through these very large disturbances with the purities of the two product streams held quite near their desired values. The variable xB3 is the mole fraction of butyl acetate in the butyl acetate product. The variable xB2 is the mole fraction of methanol in the methanol product. The Aspen default control structure of a constantmass-reflux flow rate is used on all columns. Even less variability in product compositions could be achieved by using either reflux-to-feed ratio or reflux ratio control. Figure 8B presents results for a series of step changes in the composition of the fresh methyl acetate/methanol


Figure 10. (A) Flow disturbances for the reactive distillation process. (B) Flow disturbances for the reactive distillation process with ratio. (C) Feed composition disturbances for the reactive distillation process.

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feed. At time 0.2 h, the feed composition is increased from 60 to 70 mol % methyl acetate. At time 4 h, the feed composition is decreased to 50 mol % methyl acetate. The process rides through these very large disturbances with the purities of the two product streams held quite near their desired values. The larger methyl acetate concentration in the feed produces an increase in the amount of butyl acetate produced (stream B3) and a decrease in the amount of methanol produced (stream B2). Notice that the two fresh feed flow rates also change because the flow rates of the total streams to the reactor are controlled, not those of the fresh feeds. 4.2. Reactive Distillation Process. The control scheme developed for the three-column reactive distillation process is shown in Figure 9A. The important loops from a plantwide control perspective are the two flow controllers FC1 and FC2 that control the flow rates of the methyl acetate azeotrope and total butanol streams fed to the reactive distillation. (1) The flow rate of the distillate stream D1 from column C1 is flow controlled at its desired set point (137 lb-mol/h at base-case conditions). (2) The flow rate of the stream HVYREC is measured and flow controlled by manipulating valve V32 (131 lbmol/h at base-case conditions). The set point of this controller comes from a multiplier, so the flow rate of HVYREC is ratioed to the D1 flow rate. As shown on the FC2 controller faceplate in Figure 9B, this controller is “on cascade.” (3) The reflux drum level in column C3 is controlled by manipulating the flow rate of the fresh butanol feed (stream FBUOH) using valve V2. Changing this flow has an immediate effect on the level because the total flow rate is fixed. Therefore, if the drum level is increasing and the controller decreases FBUOH, the flow rate of distillate from C3 will increase. (4) The reflux drum level in column C1 is controlled by manipulating the flow rate of the fresh methyl acetate/methanol feed. If the level decreases, more feed is added to the column, which reduces the tray temperature and increases the reboiler heat input, thus bringing the reflux drum level back up. The production rate handle in this process is the set point of the FC1 controller. The other loops are standard unit-operations control schemes for the individual units and are shown in Figure 9A. The effectiveness of this plantwide control structure is demonstrated by subjecting the process to feed flow rate and feed composition disturbances. Figure 10A gives results for a series of step changes in the set point of FC1. At time 0.2 h, the set point is increased from 137 to 164 lb-mol/h (a 20% increase). At time 4 h, the set point is decreased to 110 lb-mol/h (a 40% decrease). Finally, at time 8 h, the set point is increased to 164 lb-mol/h (a 40% increase). The process rides through these very large disturbances, providing stable baselevel regulatory control. However, the purity of the methanol product, xB1, goes through a dynamic transient, dropping to about 97.5 mol % methanol for the very large 40% increase in flow rates at 8 h. This occurs because the step change in the distillate D1 cause the reflux drum level in column C1 to decrease quickly, so the level controller increases the fresh feed fairly quickly. This rapid change in the liquid feed to the column affects the bottoms composition until the temperature controller can bring

the temperature back to its set point (temperature drops 180 to 165 °F during the transient). Notice that the butyl acetate product purity is not affected as much as the methanol product purity. This is because the flowrate changes that the third column C3 sees occur more gradually (B2 is not changed in a step function, but takes about 1 h to come to its new steady state). In the conventional process, both products come from columns that are not subjected to the rapid step increases in their feed flow rates. Therefore, the compositions of the product streams are not disturbed as much as when the rapid feed disturbance occurs in a column whose bottoms stream is a product. An obvious way to improve the load response of column C1 is to apply a feed-to-steam ratio control structure. The feed flow-rate signal is fed to a multiplier whose other input is the desired reboiler heat input to feed flow-rate ratio, which is the output signal of the TC1 temperature controller. Thus, the output of the multiplier sets the reboiler heat input. The effectiveness of this approach is shown in Figure 10B. The changes in the purity of the methanol xB1 are greatly reduced. With the ratio in operation, the initial response of the methanol purity actually moves in the opposite direction because an instantaneous ratio is used. These results indicate that a dynamic lag should be used in the ratio loop to delay the feed flow signal to the multiplier. Figure 10C presents results for a series of step changes in the composition of the fresh methyl acetate/ methanol feed. At time 0.2 h, the feed composition is increased from 60 to 70 mol % methyl acetate. At time 4 h, the feed composition is decreased 50 mol % methyl acetate. The process rides through these very large disturbances with the purities of the two product streams held quite near their desired values. 5. Conclusion Design and control of two alternative processes for the production of butyl acetate from methyl acetate have studied. Both processes are capable of producing highpurity butyl acetate and methanol without the use of an extractive agent. The reactive distillation process is more economical. Both processes can be effectively controlled using conventional PI controllers by using an appropriate plantwide control structure that manages the addition of fresh feeds and stabilizes recycle flow rates. Nomenclature AC ) heat transfer area of condenser (ft2) AR ) heat transfer area of reboiler (ft2) Bn ) bottoms flow rate from column n (lb-mol/h) BUOHTOT ) flow rate of total butanol stream to reactive column (lb-mol/h) Cj ) concentration of component j (kmol/m3) D ) diameter of vessel (ft) Dn ) distillate flow rate from column n (lb-mol/h) F ) flow rate of fresh methyl acetate/methanol feed (lbmol/h) FBuOH ) flow rate of fresh butanol feed (lb-mol/h) FCn ) flow controller n HX ) heat exchanger ID ) column diameter (ft)

Ind. Eng. Chem. Res., Vol. 43, No. 25, 2004 8025 kF ) forward reaction rate (m3 s-1 kmol-1) kR ) reverse reaction rate (m3 s-1 kmol-1) L ) length of vessel (ft) MEACTOT ) flow rate of total methyl acetate stream to reactive column (lb-mol/h) NT ) total number of trays in column QC ) condenser heat removal (Btu/h) QR ) reboiler heat input (Btu/h) R ) overall reaction rate (kmol s-1 m-3) TAC ) total annual cost (106 $/year) Tn ) temperature of control tray in column n (°F) xB1 ) composition of methanol product in reactive distillation process (mole fraction of methanol) xB2 ) composition of methanol product in conventional process (mole fraction of methanol) xB3 ) composition of butyl acetate product in reactive distillation process (mole fraction of butyl acetate)

Literature Cited (1) Jimenex, L.; Costa-Lopez, J. The production of butyl acetate and methanol via reactive and extractive distillation. II Process modeling, dynamic simulation, and control strategy. Ind. Eng. Chem. Res. 2002, 41, 6735. (2) Jimenex, L.; Costa-Lopez, J. The production of butyl acetate and methanol via reactive and extractive distillation. I Chemical equilibrium, kinetics, and mass-transfer issues. Ind. Eng. Chem. Res. 2002, 41, 6663. (3) Gangadwala, J.; Kienle, A.; Stein, E.; Mahajani, S. Production of butyl acetate by catalytic distillation: Process design studies. Ind. Eng. Chem. Res. 2004, 43, 136. (4) Luyben, W. L. Use of dynamic simulation to converge complex process flowsheets. Chem. Eng. Educ. 2004, 38, 2.

Received for review May 28, 2004 Revised manuscript received August 30, 2004 Accepted September 17, 2004 IE040167R

Design and Control of Conventional and Reactive Distillation

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