Design and Control of a Heat-Integrated Reactive Distillation Process

Dec 15, 2010 - (RD) to produce methanol and n-butyl acetate are investigated in the study. ... n-butyl acetate by the transesterification reaction of ...
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Ind. Eng. Chem. Res. 2011, 50, 1321–1329

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Design and Control of a Heat-Integrated Reactive Distillation Process to Produce Methanol and n-Butyl Acetate San-Jang Wang,*,† Hsiao-Ping Huang,‡ and Cheng-Ching Yu‡ Department of Chemical and Material Engineering, Ta Hwa Institute of Technology, Chiunglin, Hsinchu 307, Taiwan, and Department of Chemical Engineering, National Taiwan UniVersity, Taipei 106, Taiwan

The process design and control of the transesterification of methyl acetate and n-butanol by reactive distillation (RD) to produce methanol and n-butyl acetate are investigated in the study. The products from the RD column include n-butyl acetate and a methanol/methyl acetate mixture, which is separated by an additional separation column. In the study, the transesterification RD process with thermal coupling is redesigned from a previous paper. An innovative process, heat-integrated RD, which combines RD and heat-integrated distillation, is also designed to produce high-purity products of the reaction. The heat-integrated RD process provides the best energy efficiency. Proper selection and pairing of controlled and manipulated variables chosen for three control objectives were determined by using steady-state analysis for the heat-integrated RD process. A simple temperature control scheme is sufficient to maintain product purities and stoichiometric balance between the reactant feeds without resetting the set points of temperature loops. 1. Introduction Reactive distillation (RD) has attracted much attention in the last two decades and has become one of many extensively researched process intensification methods because of its ability to increase reaction conversion by removing the product and simplify separation by reacting away azeotropes. The number of papers in this field has grown rapidly in recent years for process design,1-7 steady-state behavior description,8-10 and dynamics and control.11-14 Research on RD has been comprehensively reviewed by Doherty and Malone,15 Sundmacher and Kienle,16 and Luyben and Yu.17 One of the processes that can make effective use of RD is the production of methanol and n-butyl acetate by the transesterification reaction of methyl acetate and n-butanol.18-20 A large amount of methyl acetate with a low market value is produced as byproduct in the production of polyvinyl alcohol. The methyl acetate can be transformed into a more valuable product, n-butyl acetate, by the transesterification reaction. Like esterification, transesterification is a typical equilibrium-limited reaction. Therefore, RD is used to improve conversion. However, the presence of a minimum-boiling azeotrope between methyl acetate and methanol prevents the use of only a single RD column because any unreacted methyl acetate leaving from a reaction zone in the RD will go out the column top with the methanol. Three plantwide designs using RD from four to two column configurations can be found in the literature. Jime´nez and Costa-Lo´pez18 examined a four-column process configuration; o-xylene was added to a RD column so that extractive distillation and RD were carried out in the single column. It means that the extractive solvent must be recovered in a subsequent separation unit, and unpreventable contamination of products with small amounts of the solvent occurs. Luyben et al.19 showed that the use of o-xylene is not necessary and proposed a three column design without using extractive solvent. The top and bottom products from the RD column are fed to two separation columns, respectively, to purify methanol and n-butyl acetate. Wang et * To whom correspondence should be addressed. Tel.: +886-35927700 × 2853. Fax: +886-3-5927310. E-mail: [email protected]. † Ta Hwa Institute of Technology. ‡ National Taiwan University.

al.20 proposed a flowsheet, which includes a RD column followed by only one conventional distillation column. In the flowsheet, a methyl acetate/methanol mixture, rich in methyl acetate, is fed to a RD column to react with n-butanol to produce a bottom product of n-butyl acetate. The n-butyl acetate/nbutanol azeotrope is reacted away in the column. The top product from the column is fed to a separation column to produce a distillate approximating the methyl acetate/methanol azeotrope, which is then recycled into the RD column. Methanol is recovered at the bottom of the separation column. In their study, a transesterification RD process with thermal coupling was also presented to produce n-butyl acetate and methanol. High product purities can be obtained at the bottoms of a RD column and a side stripper column. When a RD column is used, complete conversion of reactants is possible and reactant recirculation can be avoided. However, in the transesterification reaction of methyl acetate and n-butanol to produce methanol and n-butyl acetate, methyl acetate is a reactant that forms a minimum-boiling azeotrope with a product, methanol, and its recirculation is unavoidable. RD in face of azeotrope(s) in the separation of a mixture from a RD column may need debottleneck design and more columns to obtain high purity products. Pressure-swing distillation is an efficient way to perform the separation of azeotropes when the azeotropic composition varies at least 5% (preferably 10% or more) over a moderate pressure range (not more than 10 atmospheres between the two pressures).21 The azeotropic composition of methanol/methyl acetate is sensitive to pressure and is rich in methanol (bottom product in the succeeding separation column) at high pressure rather than at low pressure. This characteristic drives the RD column operation at higher pressure to increase reaction conversion and reduce energy consumption. Heat-integrated distillation is one way to improve the economic performance of a traditional distillation configuration. Its basic feature is the utilization of the heat content of the overhead vapor generated in one column to supply the heat required in the reboiler of another column. To provide the necessary temperature difference, the columns must often be operated under different pressures. Overhead vapor temperature from the high-pressure column should generate a reasonable temperature difference from the reboiler temperature of the low-

10.1021/ie100677x  2011 American Chemical Society Published on Web 12/15/2010

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pressure column. In the transesterification RD process, heat integration can be implemented between a high-pressure RD column and a low-pressure separation column to reduce energy consumption further. This heat integration reduces the overall heat demand for distillation. In this study, we will redesign the thermally coupled RD process for the transesterification reaction of methyl acetate and n-butanol. Energy efficiency and design robustness of the process are investigated by adjusting the numbers of stages of the side stripper column and the rectification zone in the RD column, used for the separation of methanol/methyl acetate azeotrope. We will also show that the RD process with the RD column operated at higher pressure can give better energy efficiency than the thermally coupled RD process. The RD process with heat integration via pressurized operation provides the best energy efficiency. The control strategy of the heatintegrated RD process will also be investigated. 2. Design of RD Process without and with Thermal Coupling The transesterification reaction of methyl acetate and nbutanol proceeds as follows: CH3(CH2)3OH + CH3COOCH3 h n-butanol methylacetate n-BuOH (MeOAc)

(1) CH3(CH2)3COOCH3 + CH3OH n-butylacetate methanol (n-BuOAc) (MeOH) The yield for the transesterification reaction is strongly limited by the equilibrium conversion. Jime´nez et al.22 studied the chemical equilibrium and kinetics of strong acid catalyzed reaction and gave a pseudohomogeneous model to predict the reaction rate: r ) 2.018 × 108 exp(-71960/(RT))Cn-BuOHCMeOAc 2.839 × 108 exp(-72670/(RT))Cn-BuOAcCMeOH (2) where r is the reaction rate and Ci is concentration of the ith component. There is only little difference between the forward and backward reaction activation energies. The reaction rate equation indicates that the heat of reaction is small and also that the equilibrium constant has little temperature dependence. The simulation of the transesterification distillation was carried out using the rigorous distillation model provided by ChemCad software. The liquid phase activities were calculated by using modified UNIFAC,23 as in Wang et al.20 The quaternary reaction system presents two minimum boiling binary azeotropes (methyl acetate/methanol and n-butyl acetate/nbutanol). The predicted azeotropic temperatures and compositions agree well with experimental results provided by Gmehling et al.24 Figure 1a is a conventional RD process used in Wang et al.,20 including a RD column and a separation column. There are three zones in the RD column with 27 (including a total condenser), 16, and 7 (including a partial reboiler) equilibrium stages in the rectification, reaction, and stripping zones, respectively (50 tray RD column). The column is operated at 1 atm. The pressure drop in the column is assumed to be 0.2 atm. High-boiling and low-boiling reactants, n-butanol, and methyl acetate at a feed rate of 100 kmol/h each, are fed immediately above and below the reactive zone, respectively. A reaction volume of 50 L and catalyst of 500 g are available on every tray in the reactive zone. n-Butyl acetate, the high-boiling

Figure 1. Process flowsheet of transesterification RD: (a) conventional configuration, (b) thermally coupled configuration, and (c) thermodynamically equivalent configuration of setup b.

product, with high-purity 99.5 mol % is withdrawn from the column bottom. The distillate product, a mixture of methanol and unreacted methyl acetate, is fed to a separation column with 10 stages. High methanol purity, 99.5 mol %, is obtained from the column bottom. The column distillate, approximating the methanol-methyl acetate azeotrope is recycled back to the fresh methyl acetate feed location. For this conventional RD process, there are four design degrees of freedom: two reflux ratios and two reboiler duties for the RD process shown in Figure 1a. The two reboiler duties are used to satisfy the purity specifications while the other two are used as the remaining degrees of freedom to minimize the total reboiler duty of the RD process. Figure 1a also shows the optimal operating condition with the minimal total reboiler duty. The total reboiler duty is minimum (8.48 Gcal/h) when the reflux ratios of the RD column and the separation column are 2.2 and

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2.84, respectively. The reboiler duties of the RD column and the separation column are 5.31 Gcal/h and 3.17 Gcal/h, respectively, under the optimal condition. Figure 1b shows the flowsheet of a thermally coupled RD process by directly interchanging the liquid and vapor between the RD column and the separation column shown in Figure 1a. A condenser is then reduced by implementing the thermal coupling. The same reaction volume and catalyst on every tray in the reactive zone are also used. To maintain parity with the conventional RD process design, the same total number of trays (60) is used in the thermally coupled RD process. The configuration in Figure 1c is the thermodynamic equivalent of the configuration in Figure 1b. The thermally coupled RD process consists of a RD column and a side stripper column with 55 and 5 trays, respectively. The maximum boiling component, n-butyl acetate, is withdrawn from the bottom of the RD column. The approximate homogeneous azeotrope of methanol and methyl acetate is distillated at the column top and recycled back to the fresh methyl acetate feed location. A side draw is designed to purge the methanol product to the side stripper column. High purity methanol (99.5 mol %) is obtained from the bottom of the side stripper column. There are five design degrees of freedom for the thermally coupled RD process shown in Figure 1c: reflux ratio, liquid split ratio, and side draw location of the RD column and reboiler duties of both columns with only two product specifications. Liquid split ratio is specified as the ratio of the liquid flow directed to the side stripper column to the flow directed to the RD column. The two reboiler duties are used to satisfy the purity specifications while reflux ratio, liquid split ratio, and side draw location of the RD column are used as the remaining degrees of freedom to minimize the total reboiler duty of the RD process with thermal coupling. Figure 1c also shows the optimal operating condition with the minimal total reboiler duty. The total reboiler duty is minimum (8.29 Gcal/h) when reflux ratio, liquid split ratio, and side draw location of the RD column are 8.0, 2.045, and 19, respectively. The reboiler duties of the RD column and the side stripper column are 4.16 Gcal/h and 4.13 Gcal/h, respectively, under the optimal condition. As compared with the configuration without thermal coupling, only 2.2% reduction of total reboiler duty is obtained by the use of thermal coupling. To improve the energy efficiency further, the number of stages in the separation column or that in the RD rectification zone are adjusted. These stages are mainly used for the separation of methanol/methyl acetate azeotrope. The number, 37, of total stages for the separation of the azeotrope is held constant. The numbers of stages in the RD reactive and stripping sections also remain unchanged. The corresponding thermally coupled distillation process can be designed by directly interchanging the liquid and vapor between the RD column and the separation column. Figure 2a shows the optimal configuration with the minimum total reboiler duty (6.95 Gcal/h) for the conventional RD process. The numbers of stages in the RD rectification section and separation column are 12 and 25, respectively. As compared with the configuration shown in Figure 1a, energy consumption can be substantially reduced (18.0%) by adjusting the number of stages mainly used for the separation of methanol/methyl acetate azeotrope in the separation column or in the RD rectification zone. Figure 2b gives the corresponding thermally coupled RD process designed by the same procedures as those used in the process shown in Figure 1c. The reboiler duties of the RD column and the side stripper column are 3.92 and 2.86

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Figure 2. (a) Conventional RD process with the minimum total reboiler duty and (b) its corresponding thermally coupled RD process.

Figure 3. Comparison of the total reboiler duties for the conventional RD process and the corresponding thermally coupled RD process for different numbers of stages of a thermally coupled RD column.

Gcal/h, respectively. As compared with the conventional configuration shown in Figure 2a, only 2.4% reduction of total reboiler duty is obtained by use of thermal coupling. Figure 3 shows the comparison of the total reboiler duties for the conventional RD process and the corresponding thermally coupled RD process for different numbers of stages in the thermally coupled RD column. The flatness of the total reboiler duty curve for the thermally coupled RD process designs, particularly when the number of RD column trays is reduced below 46, suggests the process design is more robust compared to the conventional RD process with no thermal coupling.

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Figure 4. Configuration of the RD process with high-pressure RD column.

Figure 6. Relationship between optimal total reboiler duty and pressure of RD column.

Figure 5. Relationship between pressure and MeOH azeotropic composition of MeOH/MeOAc azeotrope.

3. Design of Heat-Integrated RD Process The distillate product of the RD column is a mixture of methanol and methyl acetate. The methanol-methyl acetate azeotrope is sensitive to pressure and becomes richer in methanol as the pressure is increased. This indicates the reaction conversion in the RD column and separation efficiency in the separation column might be enhanced with the RD column operated at higher pressure. Figure 4 shows this configuration of the transesterification RD process. It is similar to that in the conventional RD process shown in Figure 2a but with a different operating pressure for the RD column. Pressure simultaneously affects reaction and separation in the RD column. As the pressure is varied, the reaction heat effects remain small with only slight variation in the equilibrium constant, which has only mild temperature dependence. Figure 5 gives the relationship between pressure and methanol azeotropic composition of methanol/methyl acetate azeotrope. Methanol azeotropic composition increases with pressure. It indicates that the low-boiling product, methanol, with higher purity may be obtained at the top of the RD column operated at higher pressure. This effect increases the reaction conversion in the RD column and reduces the total reboiler duty of the transesterification RD process. In the RD process with high-pressure RD column, there are five design degrees of freedom: two reflux ratios, two reboiler duties, and pressure of the RD column. The two reboiler duties are adjusted to satisfy the purity specifications while the two reflux ratios and the pressure of the RD column are used as the remaining degrees of freedom to minimize the total reboiler duty of the process. Figure 6 shows the pressure effect of the

RD column on the optimal total reboiler duty. Reboiler duty of the separation column decreases under the RD column operated at higher pressure due to more pure methanol fed to the separation column. By using n-butyl acetate as base component, the relative volatilities between the other components and n-butyl acetate in the RD stripping section decrease when column pressure increases. Reboiler duty increases with column pressure from the analysis of pressure effect on relative volatility. However, reaction conversion can be increased when the RD column is operated at higher pressure, and then less reboiler duty is necessary to obtain high-purity product at the column bottom and more pure methanol at the column top. These conflicting conditions, due to the complicated interaction of reaction and separation in the RD column, render the existence of an optimal pressure, minimizing the reboiler duty of the RD column. The RD column gives the minimum reboiler duty when the pressure of the RD column is equal to 3 atm. The optimal condition for the complete process with the minimum total reboiler duty is obtained when the RD column is operated at this pressure. The total reboiler duty is minimum (4.31 Gcal/h) when the two reflux ratios are 1.15 and 1.4, respectively. Figure 4 also shows the stream information under the optimal condition where the reboiler duties of the RD column and the separation column are 3.13 and 1.18 Gcal/h, respectively. As compared with the conventional RD configuration shown in Figure 2a, not only purer methanol but also lower distillate flow rate of methanol/methyl acetate mixture can be obtained at the top of the RD column in the RD process with high-pressure RD column. There is only little difference between the distillate compositions of separation columns in these two configurations. However, the distillate flow rate of the separation column in the RD process with high-pressure RD column is lower than that in the conventional RD process. These result in higherpurity reactant methyl acetate fed to the RD column. Both reboiler duties (3.13 and 1.18 Gcal/h, respectively) in the RD process with high-pressure RD column are lower than those (5.29 and 1.66 Gcal/h, respectively) in the conventional RD process. A 38% reduction of total reboiler duty is obtained for the RD process with high-pressure RD column. The RD process with high-pressure RD column can also provide better energy efficiency than the thermally coupled RD process designed in section 2. Higher energy-saving potential is offered by operating a RD column at higher pressure than by implementing thermal coupling between two columns for the transesterification reaction.

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4. Control of Heat-Integrated RD Process

Figure 7. Configuration of the heat-integrated RD process.

Figure 8. Reactant composition profiles of RD column operated at 1 and 3 atm.

In the RD process with high-pressure RD column shown in Figure 4, the condenser duty is 2.80 Gcal/h for the high-pressure RD column while the reboiler duty is 1.18 Gcal/h for the lowpressure separation column. The condenser temperature (88.6 °C) of the RD column is higher than the base temperature (64.2 °C) of the separation column. The difference between the condenser temperature of the RD column and the base temperature of the separation column is large enough to make heat integration an attractive way to save energy consumption in the RD process with high-pressure RD column. The heat removed from the condenser of the RD column can be used as the heat input to the base of the separation column and an auxiliary condenser is required with a duty of 1.62 Gcal/h. Figure 7 shows the heat-integrated RD process. The total reboiler duty can be further reduced by 27.4% via the use of heat integration between the RD column and the separation column. As compared to the conventional RD process, a 55% reduction of the total reboiler duty can be obtained for the heat-integrated RD process. The heat-integrated RD process provides the best energy efficiency among different designs for the transesterification reaction. Figure 8 shows the reactant composition profiles of RD columns operated at 1 and 3 atm. Poorer reactant distribution is observed for the RD column operated at 1 atm. Deviation from stoichiometric condition (n-butanol/methyl acetate ) 1:1) is higher in the reaction zone. Higher reflux ratio and reboiler duty are required to achieve the desired product purity. However, reactant distribution is improved by operating the RD column at 3 atm. The control strategy of the heat-integrated RD process with the best energy efficiency is discussed in the following section.

Normally, for an equilibrium-stage column, the temperature and composition profiles are independent of the feed rate, as long as the quality specifications are the same. Such invariance, however, is not valid for a kinetically controlled RD,25,26 since the residence time in the reactive zone changes as production rate varies. Hence there are three keys to controlling such a column: (1) maintain the correct stoichiometric balance between the feeds, (2) maintain the product quality, (3) account for possible changes in temperature/composition set-point changes for throughput changes. When operating a RD column, the feed must be kept in correct stoichiometric ratio. The simplest way to maintain this balance is to use a feed ratio control. However, a feed-ratio control scheme is a feedforward scheme that has no guarantee of the final product quality. When there is a measurement bias in the feed flow rate, the feed-ratio control scheme will not be able to maintain stoichiometric balance. To overcome this problem, Al-Arfaj and Luyben27 suggested that the reactant composition of a certain column stage be controlled by the reactant feed flow. An internal composition control scheme requires the use of online analyzer, which is usually expensive and slow in response. In industrial applications, temperature control is usually used instead of composition control. Stage temperatures are employed in the study to infer product compositions. For a faster response, a feed-ratio + temperature control scheme is utilized in the study to maintain stoichiometric balance. Methyl acetate and n-butanol feed flows are two candidates for the manipulated variable of the internal temperature control to maintain stoichiometric balance. RD processes often exhibit input multiplicities, which increase the difficulties of operation and control.28,29 Input multiplicity occurs for the transesterification RD process with thermal coupling.20 The stage temperatures with input multiplicity or interaction multiplicity with manipulated variables should not be used as controlled variables to avoid the problem of control stability.28 Figure 9 panels a and b show the effect of changes in n-butanol feed flow and methyl acetate feed flow, respectively, on the changes of stage temperatures in the RD column. It was found that there is no input multiplicity between these two reactant feed flows and stage temperatures for the heat-integrated RD process. There are thus several possible candidates for tray temperature control. To maintain high-purity n-butyl acetate at the bottom of the RD column, one temperature loop can be used, in which a stage temperature is controlled by manipulating the reboiler duty of the column. Figure 9c shows the effect of changes of reboiler duty on the change of stage temperatures in the column. Input multiplicity also does not occur between reboiler duty and stage temperatures in the column. For the RD column in the kinetic regime, different desired temperature profiles can be found for various throughput rates due to residence time change on the reactive stages as the throughput rate changes. Hence the set point of the temperature control loop must be adjusted as long-term disturbances occur. This will cause a lot of difficulty in column operation. Figure 10 shows the deviations of the desired steady-state temperature profiles of RD column and separation column from their nominal ones when both reactant feed flows are simultaneously increased or decreased by the same amounts, 20%, and with bottom n-butyl acetate and methanol compositions kept at their respective design values. There is almost no change for the desired steady-state temperature profile of the separation column. In the RD column, only very small changes are observed between the temperatures of stages 10

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Figure 10. Effect of feed flow rate on the deviations of desired steadystate temperature profiles for (a) RD column and (b) separation column in the heat-integrated RD process.

Figure 9. Effect of changes of (a) n-butanol feed flow, (b) methyl acetate feed flow, and (c) reboiler duty on changes of stage temperatures in RD column.

and 26 and no changes are necessary for the other stage temperatures. It indicates that there is almost no need to change the set points of temperature loops to maintain stoichiometric balance and product purities as the throughput rate changes for the heat-integrated RD process. To select the most appropriate location used for control purpose in the RD column, singular value decomposition

method30 is used in the study. The controlled temperatures are selected at stages 12 (T12) and 15 (T15) when reboiler duty and n-butanol feed rate or reboiler duty and methyl acetate feed rate are used as manipulated variables for the RD column. In the thermally coupled RD process,20 n-butanol feed flow was used as a manipulated variable to maintain stoichiometric balance, and inverse temperature loop pairing was designed in the RD column. Inverse loop pairing means that control loop pairing has a controlled variable near the top of the column manipulated by reboiler duty and a controlled variable near the column bottom manipulated by n-butanol feed flow. The following shows the steady-state gain and relative gain array (RGA) analysis of controlled stage temperatures and manipulated variables in the RD column for the heat-integrated RD process.

[ ] [ [ ] [

T12,RD -0.6135 ) T15,RD -1.2611 T12,RD -1.2443 ) T15,RD -1.8104

][ ] ][ ]

1.9631 m1 4.3790 m3 1.9631 m2 4.3790 m3

(3)

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Figure 11. Control scheme for the heat-integrated RD process.

Figure 12. Dynamic responses of controlled stage temperatures, methyl acetate feed flow, and product purities under temperature control for (20% changes in n-butanol feed flow.

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T12,RD T15,RD 12.74 -11.74 m1 RGA ) -11.74 12.74 m3

][ ]

[

(4)

T12,RD T 15,RD 2.88 -1.88 m2 RGA ) -1.88 2.88 m3

][ ]

[

These results reveal that the simple temperature control scheme is sufficient to maintain product purities and stoichiometric balance between the reactant feeds without changing the set points of temperature loops for the heat-integrated RD process. 5. Conclusions

where m1, m2, and m3 represent the n-butanol feed flow, methyl acetate feed flow, and reboiler duty of RD column, respectively. Much more interaction exists between the temperature loops when n-butanol feed flow is used as the manipulated variable to maintain stoichiometric balance. Better controllability is obtained if methyl acetate feed flow is used as the manipulated variable. Hence the control loop pairing used in the RD column is that the temperatures at stages 12 and 15 in the column are controlled by manipulating the methyl acetate feed flow and reboiler duty, respectively, to maintain stoichiometric balance and bottom product purity. To maintain high-purity methanol at the bottom of the separation column, one temperature loop is used, in which a stage temperature is maintained by manipulating the reboiler duty of the column. The controlled stage temperature can be selected from an open-loop sensitivity analysis. The open loop test selects the temperature of stage 21 in the separation column as the controlled variable of the temperature loop due to its maximum sensitivity to reboiler duty change. In the following discussions on control performance, the controllers were tuned using a sequential design approach31 after the pairing was determined. For each controller, relay-feedback test32 is performed to obtain ultimate gain and ultimate frequency. The following equations are used to calculate the tuning parameters of PI controllers:

The thermally coupled RD process in Wang et al. (2008) is redesigned to produce n-butyl acetate and methanol by the transesterification reaction of methyl acetate and n-butanol. Energy efficiency and design robustness of the RD process are investigated by adjusting the numbers of stages used for the separation of methanol/methyl acetate azeotrope. Simulation results indicate that the thermally coupled process can provide better energy efficiency and design robustness than the conventional RD process without thermal coupling. The pressure effect is also exploited in the transesterification reaction to increase the reaction conversion in the RD column and reduce the total reboiler duty of the process. The RD process with the RD column operated at higher pressure is shown to give better energy efficiency than the thermally coupled RD process. The RD process with heat integration via pressurize operation provides the best energy efficiency. In the heat-integrated RD process, methyl acetate feed, different from the feed used in the thermally coupled RD process, is selected by the steady-state analysis to maintain the reactant inventory in the RD column due to its better temperature controllability. Dynamic simulation results show that the temperature control strategy, without changing the set points of temperature loops, can maintain reactant inventory in the RD column and two bottom product purities almost at their desired operating values.

KC ) KCU /3

(5)

Acknowledgment

TI ) PU /0.5

(6)

This work is supported by the National Science Council of ROC under Grant No. NSC 99-2221-E-233-008.

where KC and TI represent proportional gain and integral time, respectively, and KCU and PU are ultimate gain and ultimate period, respectively. Figure 11 shows the control scheme for the heat-integrated RD process. Three temperature control loops are designed by the above steady-state analysis. In the RD column, column pressure is controlled by manipulating coolant flow rate in the auxiliary condenser. Reflux-drum level and base level are maintained by changing reflux flow rate and bottom flow rate, respectively. Reflux ratio is controlled by manipulating distillate flow rate. The n-butanol feed is flow controlled. In the separation column, column pressure is controlled by manipulating coolant flow rate in the condenser. The control strategies of reflux-drum level, base level, and reflux ratio are the same as those used in the RD column. In the distillation control system, pressure, level, and flow control belong to inventory control maintaining the basic operation of column. Thus in the following discussion, emphasis is placed on the response of temperature control strategy used to maintain stoichiometric balance and product quality. In the RD column, the temperature loop to maintain bottom product purity is tuned first, followed by that to maintain stoichiometric balance. Figure 12 shows the dynamic responses of controlled stage temperatures, methyl acetate feed flow, and the corresponding product purities under temperature control for (20% n-butanol feed flow disturbances. The controlled stage temperatures are settled at their corresponding set points. Stoichiometric balance is maintained and two product compositions can almost return to their desired operating values under temperature control.

Nomenclature Ci ) molar concentration of i component (mol/L) Kc ) proportional gain Kcu ) ultimate gain m1 ) n-butanol feed flow m2 ) methyl acetate feed flow m3 ) reboiler duty of RD column Pu ) ultimate period R ) gas constant (J/mol · K) RD ) reactive distillation r ) reaction rate (mol/L · min · gcat) T ) temperature (K) TI ) integral time

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ReceiVed for reView March 19, 2010 ReVised manuscript receiVed December 4, 2010 Accepted December 6, 2010 IE100677X