Design and Control of Reactive Distillation Process ... - ACS Publications

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Design and Control of Reactive Distillation Process for the Production of Methyl Valerate Cheng-Liang Chen,† Yao-Hsien Chung,† and Hao-Yeh Lee*,‡ †

Department of Chemical Engineering, National Taiwan University, Taipei 10617, Taiwan Department of Chemical Engineering, National Taiwan University of Science and Technology, Taipei 10607, Taiwan

‡

S Supporting Information *

ABSTRACT: Methyl valerate (VAME), also known as methyl pentanoate, is a methyl ester of pentanoic acid (valeric acid). VAME is usually used as a fragrance in the production of beauty care, soap, and laundry detergents. High purity VAME can also be used as a kind of plasticizer. This study presents design details of the process for the manufacture of VAME. A reactive distillation column (RDC) is used in the production process to overcome equilibrium limitation of the esterification reaction. There is no bottom outlet of this RDC. Both products, VAME and water, are withdrawn from the distillate of RDC and then can be separated by two strippers and a decanter. A thermally coupled design is then developed to reduce the remixing effect in the rectifying section of the RDC. The simulation results show that 30% energy saving can be achieved by using the proposed thermally coupled configuration, but only 17% of total annual costs can be saved due to the use of a compressor. Control strategies of both conventional and thermally coupled configurations in neat design are investigated. The simulation results show that a thermally coupled configuration can reject disturbances faster with smaller steady state deviations from the specifications of the VAME product.

1. INTRODUCTION Process intensification is an important trend in chemical process technology which has drawn a lot of attention in both industry and academic communities. According to Tsouris and Porcelli’s definition, process intensification refers to technologies that replace large, expensive, energy-intensive equipment or processes with ones that are smaller, less costly, or more efficient or that combine multiple operations into fewer devices (or a single apparatus).1 Since the crude oil price has increased almost four times during the past decade, reduction of energy consumption during the production processes has become one of the main research interests.2 Distillation is most widely used in chemical and petrochemical industries for separation despite its high energy consumption. It contributes about 40% to the total energy of these industries.3 Thermal coupling between two distillation columns is one of energy-integrated methods, which can save energy. Specifically, it contains heat and mass transfer through the material streams. Thermally coupled distillation (TCD) has been proven that it can save, on average, about 30% of energy consumption compared with conventional arrangement.4 Hernández et al.5 showed that TCD can decrease energy consumption by eliminating a remixing effect, in contrast to conventional distillation sequences. The remixing effect makes the conventional sequences inefficient due to the energy required for repurifying the intermediate component in another column. Of three common configurations, fully coupled is described as the most energy-efficient6 and has been known for nearly 50 years. However, it is not widely used in industry despite its attractive features, which can be attributed to the lack of design procedures and the difficulty in control. Agrawal and Fidkowski7 also showed that, among the three configurations, the side stripper and side rectifier tend to be the most efficient configuration than the fully © XXXX American Chemical Society

coupled configuration over a wide range of relative volatilities and feed compositions. Replacement of some single function units with a multifunctional unit is an effective way for process intensification. Reactive distillation (RD) is one of the major applications. RD configuration is the combination of a reactor and a separation unit which can increase the conversion by Le Chatelier’s principle. In recent years, more and more papers studying RD have been published. As for the applications of RD, Malone and Doherty8 gave a good review paper of RD technology. From the book of Luyben and Yu,9 it was shown that there were 1105 related publications and 814 U.S. patents between 1971 and 2007. Luyben and Yu9 also highlighted 236 reaction systems which can be designed with RD configuration. Also, in Sundamacher and Kienle’s book,10 more than 100 industrially important reactions for applications are summarized by Sharma and Mahajani in one chapter. Tung and Yu11 studied the effects of relative volatility ranking on the design of RD processes and provided a systematic design procedure to determine the process configuration. For an ideal quaternary system (A + B ⇔ C + D), all possible relative volatilities can be categorized into six configurations based on relative volatilities of reactants and products.11 Muller et al.12 was the first researcher who combines the RD with the divided wall column configuration, also known as reactive divided wall column (RDWC), for the simultaneous esterification of methanol and n-butanol with acetic acid to produce methyl acetate, n-butyl acetate, and water. There are a Received: September 21, 2015 Revised: January 7, 2016 Accepted: January 15, 2016

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DOI: 10.1021/acs.iecr.5b03495 Ind. Eng. Chem. Res. XXXX, XXX, XXX−XXX

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Industrial & Engineering Chemistry Research lot of open literature sources showing that the thermally coupled reactive distillation (TCRD) has higher energy efficiency and is more economically favorable than the conventional RD process. Wang et al.13 studied TCRD for the production of methanol and n-butyl acetate by transesterification reaction of methyl acetate and n-butanol. Both energy reduction and controllability are better in the thermally coupled reactive distillation process. Lee et al.14 showed that using TCRD can achieve 23.14% energy reduction in the i-propyl acetate production process. Their work includes the decanter in the original design flowsheet. Wu et al.15 developed the TCRD system for the esterification of an alcohol mixture containing n-amyl alcohol and n-hexanol. This has a 34% reduction in total annual cost as compared with the original indirect sequence. As for experimental validation of RDWC, Sander et al.16 conducted the laboratory and miniplant scale experiments for the methyl acetate hydrolysis process. Hernández et al.,17 and Delgado−Delgado et al.18 also provided experimental results for the production of ethyl acetate in a RDWC. All these results showed good correspondence to the simulated processes. Both reactive distillation and thermally coupled distillation are very promising technologies and of great interest to the industry and the research community. Also, they can achieve substantial and economical benefits from process intensification. In this work, a technology combining RD with thermally coupled configuration is investigated for methyl valerate (VAME) production process. VAME is commonly used in fragrances. High purity VAME is also used as a plasticizer in the production of plastics. Traditionally, VAME is produced by the esterification reaction of valeric acid (VA) and methanol (MeOH). However, most esterification reactions have a chemical equilibrium limitation which results in low conversion. In view of this

Figure 1. Relationship between VA conversion and reaction temperature.

problem, reactive distillation should be a suitable technology for the VAME production process. However, it is very challenging in designing such kind of reactive distillation process because reactants are the lightest and heaviest components, respectively, and products are components of middle boiling point in the system. Also, this is an example with temperature constraint in the RD column. In this work, conventional RD process and thermally coupled RD process will be developed and optimized for the production of VAME. After the proposed design flowsheets are developed, it is of interest that whether or not these flowsheets can be properly controlled in spite of various disturbances. Consequently, overall control strategies for two processes are also studied. Commercial simulators Aspen Plus and Aspen Plus Dynamics are used for steady-state and dynamic simulations, respectively.

2. KINETICS AND THERMODYNAMICS 2.1. Kinetic Data. In the VAME production process, valeric acid (VA) and methanol (MeOH) are reacted to yield VAME and byproduct water. Amberlyst 35 is used as the catalyst.

( −5374.07 )aVAaMeOH − 7.2424 × 107 exp( −3280.47 )aVAMEawater T T

1.1246 × 1012 exp −rA =

[1 + 10.2930aVA + 79.1875aMeOH + 3.8427aVAME + 274.48a water ]2

Kinetics expressed as in eq 1 was provided by Lee.19 The kinetic equation is with the unit of (mol/kg min), where “ai” represents activity, which can be predicted by nonrandom twoliquid (NRTL) model. T is absolute temperature in Kelvin. R is the ideal gas constant, 8.314 kJ·kmol−1. The esterification reaction usually has a very slow reaction rate and low conversion. Figure 1 shows the relation between the fractional conversion of VA and the reaction temperature in an extremely large continuous stirred-tank reactor (CSTR) with equal amounts of two reactants. It can be noted that high operating temperature is needed if high conversion of VA is to be obtained, which will exceed the high temperature limitation of the catalyst. Therefore, RD is worth considering in the production of VAME. 2.2. Thermodynamic Data. VAME is a new component in the Aspen Plus databank. It is needed to be created for the estimation of physical properties. Besides the molecular structure of VAME, the boiling point, freezing point, and the specific

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gravity are also inputted into the simulator. The boiling point and freezing point under 1 atm are 128 °C and −91 °C, respectively. The specific gravity is 0.875. NRTL activity coefficient model is chosen to describe the phase equilibrium behavior in this process because this model can handle a nonideal system with liquid phase splitting. There are totally four components in the system and six pairs of binary parameters are needed. Of all of them, VA−water and MeOH− water are used Aspen built-in parameters. Others do not exist in Aspen Plus databank and are estimated by UNIFAC group contribution method. All the values are shown in Table 1. The result of the calculation of NRTL model shows that there is one azeotropic composition with VAME and water. The ranking of boiling point and azeotrope under 1 atm is shown in Table 2. Figure 2 is the T−xy diagram for VAME and the water binary system. From Figure 2, it is found that the VAME/water azeotrope is a heterogeneous azeotrope. It can also be observed that there is liquid−liquid splitting behavior between VAME and B


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Industrial & Engineering Chemistry Research Table 1. Values of NRTL Binary Parameters component i component j source temp. unit Aij Aji Bij Bji Cij

MeOH water ASPEN-VLE °C −0.693 2.7322 172.9871 −617.2687 0.3

MeOH VAME UNIFAC °C 0 0 283.7547 213.8551 0.3

VA MeOH UNIFAC °C 0 0 422.54962 −286.488262 0.3

a

temp (°C)

composition (mole basis)

MeOH VAME/watera water VAME VA

64.53 86.77 100.02 128.00 185.84

1 (0.2729/0.7271) 1 1 1

VA VAME UNIFAC °C 0 0 −237.8980 512.8860 0.3

VA water ASPEN-LLE °C −8.4137 7.8996 2286.1772 −711.9017 0.2

reactive distillation for ethyl acetate process.20 (4) The weir height is set as 0.1016 m. As a result, the liquid holdup of each tray is about 0.09 m3. This RD column is composed of five reactive stages and four rectifying stages. The temperature of the last stage will be over 150 °C, which cannot be served as reactive stages. Since there is no bottom outlet, only one degree of freedom remains. As the only degree of freedom, reflux rate is adjusted to reach 99% conversion of VA. Simulation result shows that 3.23 Gcal/h is needed in the reboiler duty. There are two distillation columns operating at 1 bar in the separation sections. The outlet of the RD column is sent to the first distillation column, C1. VAME is obtained at 99 mol % from the bottom of C1. The overhead of C1 with azeotropic composition is condensed and then sent into a decanter. Two phases, organic phase and aqueous phase, are separated in the decanter, which is operated at about 80 °C. The organic phase is sent back to the first stage of C1, while the aqueous phase is fed on the first stage of the second distillation column, C2. High purity water comes out from the bottom of C2, with 99 mol % of water. Both C1 and C2 are stripping columns so there is only one degree of freedom remaining in each column. Reboiler duties of both C1 and C2 are used to reach design specifications. The results of C1 and C2 reboiler duties show 1.75 Gcal/h and 0.06 Gcal/h, respectively. The temperature of the decanter is an important design variable. In some processes, the inlet of a decanter will be cooled down first because the lower temperature will provide better liquid−liquid separation for two phases. In other words, concentrations of dominant components in both organic and aqueous phases are higher if the decanter is operated at lower temperature. If this is the case, reboiler duties required to obtain high purity products in distillation columns are lower as well. However, in the VAME process, the inlet flow of the decanter is not cooled down because, as seen from Figure 2, there is not much difference in phase splitting from 80 and 40 °C. Therefore, the energy savings will be not significant if the decanter is operated at a lower temperature. Instead, recycle streams of two distillation columns with a lower temperature may lead to higher energy consumption in two columns. As a result, the decanter is operated at about 80 °C in this design. 3.2. Optimization of Design Flowsheet. Since there is no recycle stream back to the RD column from the separation section, optimization of the RD column and the separation section can be conducted separately. The total annual cost (TAC) analysis is used to find the optimal design. Equation 2 depicts the TAC which consists of the annual operating cost (AOC) and the annualized total capital cost (TCC) with a payback period of 3 years. The calculation is based on the work of Douglas,21 ignoring costs of piping and pumps.

Table 2. Boiling Points Ranking under 1 atm component

VAME water UNIFAC °C 0 0 570.3880 1778.6885 0.3

Heterogeneous azeotrope.

Figure 2. T−xy diagram for VAME−water at 1 bar.

water while the temperature is lower than the VAME/water azeotrope. It implies that the heterogeneous azeotropic distillation with a decanter can be easy to apply for the separation of these two components.

3. STEADY STATE DESIGN 3.1. Conventional Reactive Distillation Configuration. Figure 3 shows the base case design of conventional reactive distillation configuration with stoichiometric feed ratio of MeOH and VA. The goal of this process is to convert 99% of VA and obtain 99 mol % of VAME and water. VA and MeOH, enter a RD column from the bottom. The RD column is operated at a vacuum condition, 0.7 bar. The reason is that the maximum temperature the catalyst, amberlyst 35, can tolerate is 150 °C. As a result, the temperature in the reactive zone cannot exceed this limitation. Operating pressure is chosen to meet the constraint on reactive zone temperature. Four assumptions are made while designing the RD column: (1) The downcomer area occupies 10% tray area. (2) Half of the liquid holdup is filled with catalyst Amberlyst 35 in each reactive stage. Catalyst density is assumed as 800 kg/m3. (3) There is no need to design stripping trays in the RD column. The reason is that the heaviest reactant VA stays at the lower section in the column; therefore, the lower section of the column should serve as reactive trays. The concept can also be seen in the design of C


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Figure 3. Base case design.

Figure 4. (A) Iterative optimization procedure of reaction section. (B) Iterative optimization procedure of separation section.

D


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Figure 5. Results of optimization.

TAC = AOC +

TCC 3

4. Go back to step 2, change Nr until minimum TAC is obtained. 5. Go back to step 1, change NF until minimum TAC is obtained. 3.2.2. Optimization of the Separation Section. In the separation section, reboiler duties of two columns are used to meet the design specifications. Consequently, there are totally three remaining design variables which need to be optimized, including (1) numbers of C1 total stages (NC1); (2) numbers of C2 total stages (NC2); and (3) feed stage of C1 (NF1). The optimization procedure is shown in Figure 4B. The steps of optimization are as the following: 1. Guess NC2. 2. Guess NC1. 3. Change the value of feed stage of C1 (NF1) until minimum TAC is obtained. 4. Go back to step 2, change NC1 until minimum TAC is obtained. 5. Go back to step 1, change NC2 until minimum TAC is obtained.

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3.2.1. Optimization of the RD Column. For a typical RD column of a quaternary system, there are a total of seven design variables, including (1) numbers of reactive stage (Nrxn); (2) numbers of rectifying stage (Nr); (3) numbers of stripping stage (Ns); (4) feed stage of one reactant; (5) feed stage of the other reactant; (6) reflux rate; and (7) reboiler duty. In this design flowsheet, it is assumed that there is no stripping section and the lightest reactant, MeOH, should be fed to the RD column from the bottom. Also, there is only one degree of freedom since only one outlet exists and it is used to reach the design specification. Thus, the remaining design variables to be optimized are (1) numbers of reactive stage (Nrxn); (2) numbers of rectifying stage (Nr); and (3) feed stage of VA (NF). The optimization procedure is shown in Figure 4A. The steps of optimization are as follows: 1. Guess NF. 2. Guess Nr. 3. Change Nrxn until minimum TAC is obtained. E


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Figure 6. Optimal flowsheet of conventional RD configuration.

3.2.3. The Result of Optimization. Figure 5 shows the result of flowsheet optimization. The capital cost will increase when numbers of rectifying and reactive stages increase. On the contrary, the operating cost will decrease due to the reduction of energy consumption in the RD column. The feed stage of VA does not affect TAC much. However, the reaction rate at the lower stages will decrease a lot if VA enters the column from upper stages. Therefore, it is suggested that VA be fed to the RD column from the bottom. The increase on numbers of total stages of C1 and C2 both improve the separation ability of distillation columns. In other words, energy required for the separation in the columns will decrease. However, the increase of capital cost is obvious. Through TAC analysis, an optimal flowsheet is shown in Figure 6. The RD column has a total of seven reactive stages and six rectifying stages. VA is fed from the bottom of the RD column. C1 and C2 have 11 and 4 total stages, respectively. The inlet stream of C1 is fed on stage 7. Figure 7 is the composition profile of the RD column. It can be seen that the composition profile is quite flat in the reactive

stages. This implies that the temperature change is not significant in this section. Notice that the temperature of the reactive zone ranges from 141.0 to 146.6 °C. The equilibrium constant ranges from 99.1 to 105.9. Another observation from the composition profile is that the concentration of VAME increases when it gets closer to the top of the column. However, it decreases between stages 2 and 1. This is so-called remixing effect. Remixing of the product, VAME, results in the waste of energy in the column. As a result, this phenomenon indicates a very good potential of using a thermally coupled configuration to save energy consumption. Therefore, thermally coupled reactive distillation will be discussed in the next section. 3.3. Thermally Coupled Configuration. Figure 8 shows the optimal flowsheet of thermally coupled reactive distillation. To see its ability of energy saving, all equipment sizes have remained the same as those in the conventional configuration. In this way, all energy savings are attributed to the thermal-coupling instead of column sizes. Vapor out of the RD column is first compressed to higher than 1 bar in order to feed it to third stage of C1. The liquid flow on the third stage is partially recycled back to the RD column. The separation section is the same as that of the conventional configuration. In this flowsheet, C1 has totally two degrees of freedom, including reboiler duty and feed location. Reboiler duty is used to satisfy the purity specification while the remaining degree of freedom, feed location, is used to minimize the TAC. The reboiler heat inputs for the RD column and C1 are 2.88 Gcal/h and 0.146 Gcal/h. The reboiler duty of the RD column remains almost the same as that in the conventional configuration. In contrast, much of the reboiler duty of C1 is reduced by implementing thermal coupling between the RD column and C1. The composition profile is shown in Figure 9. It is found that the composition at the bottom section of an RD column profile is almost the same to the conventional RD design. However, it can be noted that the remixing effect in the upper section of the RD column is eliminated; 30.3% energy consumption is saved in thermally coupled configuration as shown in Table 3. However,

Figure 7. Composition profile of the RD column of a conventional RD configuration. F


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Figure 8. Flowsheet of thermally coupled configuration.

Control loops can be divided into two levels. One is inventory control, including liquid level, flow rate, and pressure control. It is related to the material balance in the process and ensures the process safety. The other one is quality control, which consists of temperature and composition control. It is used to maintain the product purity. Though the product specifications are usually product purity, a temperature controller is used instead of composition controller. The reason is that most composition analyzers, such as gas chromatographs, have large measurement delay, high capital cost, and maintenance cost. 4.1. Design of Control Loop. In this paper, the conventional proportional-integral (PI) controllers are used for flow, temperature, and pressure control loops. Simple proportional (P) controllers are used for level control. The measurement points of temperature in three distillation columns are found out through the sensitivity test. The controlled stage temperature should have the largest gain toward the manipulated variables. Moreover, the autotune variation (ATV) method is used to tune controller parameters (Åström and Hägglund).22 ATV starts from the relay-feedback method which can help us get ultimate gain and ultimate period. Then controller parameters are calculated using the Tyreus−Luyben turning relations.23 The sequential iterative tuning procedure is used to find the final settings. Moreover, first order time lags are added in the temperature control loop for improving the realism of the simulation. The ATV method can be conducted in the simulator Aspen Plus Dynamics. To find out the stage with the highest sensitivity of temperature toward manipulated variables, an open loop temperature sensitivity test is conducted. Feed ratios ±1% are given to the RD column and reboiler duties ±0.1% are given to C1 and C2. When the steady state is reached, temperature sensitivity is analyzed. The stages with largest temperature variation are chosen to be temperature-controlled because they are the most sensitive. The results are shown in Figure 10. After control structures are developed, they are tested with throughput and composition disturbances in order to see their

Figure 9. Composition profile of the RD column of a thermally coupled configuration.

Table 3. Comparison of Conventional Configuration and Thermally Coupled Design Energy consumption (Gcal/h) conventional RD design thermally coupled design Difference (%)

AOC (106 USD)

TIC/3 (106 USD)

TAC (106 USD)

4.66

2.14

0.62

2.76

3.25

1.46

0.82

2.28

−30.3

−31.8

+32.0

−17

total installed cost increases by 17% because an extra compressor is required in the thermally coupled configuration.

4. PROCESS CONTROL This section will investigate control strategies for both conventional and thermally coupled processes. The purpose of the control strategies is to keep the conversion of VA at 99% and purity of VAME and water at 99 mol %. G


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Figure 10. (A) Open loop sensitivity test of conventional configuration; (B) open loop sensitivity test of thermally coupled configuration.

4.2. Control Structure of Conventional Configuration. The control structure is provided in Figure 11A. In the RD column, the sump level is controlled by the flow rate of the steam of the reboiler. Since it is operated at vacuum condition, the pressure should be controlled by the recycle flow rate at the

abilities of rejecting disturbances. Throughput disturbances are the throughput ±20%. Since feeds are assumed to be pure components in steady state design, composition disturbances include (1) 3% water in feed MeOH; (2) 5% water in feed MeOH; (3) 3% water in feed VA; and (4) 5% water in feed VA. H


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Figure 11. (A) Control structure of conventional configuration. (B) Control structure of thermally coupled configuration.

outlet of the vacuum pump. However, it is not able to simulate this pressure control loop in the simulator. Therefore, the pressure is controlled by the condenser duty instead in the

simulation work. The temperature of the fourth stage is controlled by manipulating the feed ratio. In fact there are two possible manipulated variables for the temperature control, the I


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Figure 12. continued

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Figure 12. Dynamic responses of conventional configuration: (A) throughput disturbances; (B) feed MeOH composition responses; (C) feed VA composition responses.

feed ratio and the reflux rate. The reason for choosing feed ratio is to reject composition disturbances. When there are feed composition disturbances, in order to keep two reactants at stoichiometric ratio, feed ratio should not be maintained at one. Reflux rate is set to be proportional to the total feed flow rate with the ratio equal to 32, which applies the concept of feedforward control. Units of feed flow rate and reflux rate are respectively kmol/h and kg/h. In both C1 and C2, the sump level is controlled by bottom flow rate and the pressure is controlled by the condenser duty. In reality, there is always a valve after the condenser which can be used for pressure control by changing the liquid level in the condenser, thus changing the heat transfer area of the condenser. When the pressure is too high, the opening of the valve will increase. With lower liquid level in the condenser, heat transfer area increases and more vapor is condensed, resulting in a decrease on column pressure. Nevertheless, in the dynamic model, there is no way to change the liquid level or heat transfer area in the condenser. Instead, condenser duty can be used to control the column pressure by changing the reflux temperature. But that cannot be applied to this case. The reason is that the decanter temperature should also be controlled if the outlet temperature of the condenser changes, there will be a conflict between the pressure control and the temperature control of the decanter. Consequently, two virtual valves at the top of two distillation columns are added in this work. The pressure of two columns is controlled by manipulating these two virtual valves. When the pressure is too high, valve opening increases and more vapor comes out. It may be questioned that whether or not reflux rates of two columns are thus affected. However, this is a

common configuration that is often used in the dynamic simulation work. Liquid levels of the two phases in the decanter are also controlled by the flow rate of organic and aqueous streams. Temperatures of the eighth stage in C1 and the first stage in C2 are controlled by the reboiler duties of C1 and C2, respectively. Figure 12 shows the dynamic responses under throughput disturbances and composition disturbances. Disturbances are given to the process at the third hour. The system has fast responses. At 13th hour, all kinds of disturbances can be rejected and the process reaches steady state again. Composition disturbances can be rejected with very small steady state value deviation, which can almost be neglected. While handling throughput disturbances, the largest steady state value deviation of VAME purity is 0.3% when the throughput −20%. In contrast, water purity has the largest deviation of 0.01% when the throughput +20%. 4.3. Control Structure of Thermally Coupled Configuration. The control structure is provided in Figure 11B. In the RD column, sump level is controlled by the reboiler duty, and the pressure is controlled by the work of the compressor. The temperature of the third stage is controlled by the feed ratio, and the flow rate of the liquid from C1 with unit in kmol/h is equal to the product of 0.53 and total feed mole flow rate. As for C1 and C2, the sump level is controlled by the bottom flow rate, and the pressure is controlled by the condenser duty. However, two virtual valves at the top of these two columns are used to control the pressure which is the same as that in the conventional configuration. The reason has been mentioned in section 4.2. Two reboiler duties are used to control temperature K


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Figure 13. continued

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Figure 13. Dynamic responses of thermally coupled configuration: (A) throughput disturbances; (B) feed MeOH composition responses; (C) feed VA composition responses.

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of the fifth stage in C1 and the first stage in C2, respectively. The liquid level of two phases in the decanter is controlled by the organic flow rate and the aqueous flow rate. All parings of quality control loops for temperature controllers are the same as the conventional configuration. Figure 13 shows the dynamic responses under throughput disturbances and composition disturbances. All disturbances are given to the process at the third hour, and the system can reach steady state again at the 10th hour. Composition disturbances can be rejected very well. When there are throughput disturbances, VAME purity has negligible deviation from the set point while water purity has approximately 0.01% deviation.

ASSOCIATED CONTENT

S Supporting Information *

The Supporting Information is available free of charge on the ACS Publications website at DOI: 10.1021/acs.iecr.5b03495. Table S1, controller parameters for inventory control loops; Table S2, controller parameters of temperature controllers in conventional configuration; Table S3, controller parameters of temperature controllers in thermally-coupled configuration (PDF)

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AUTHOR INFORMATION

Corresponding Author

5. CONCLUSION Process intensification is an important issue in the industry and research community. Reactive distillation and thermal coupling are two applications. In this article, the methyl valerate process is designed with a conventional reactive distillation column and a thermally coupled reactive distillation column. The simulation result shows that the thermally coupled configuration can save 30% energy consumption in the whole system. However, due to an extra compressor required in the thermally coupled design, only 17.6% of total annual cost savings is reached. An overall control strategy is also proposed for both configurations with only one temperature control loop in each column. Variations in feed composition and throughput changes are investigated. All product compositions can be maintained at high purity. However, the thermally coupled configuration can reach new steady state faster than the conventional configuration with smaller deviation of product purities.

*Tel: +886-2-2737-6613. Fax: +886-2-2737-6644. E-mail: [email protected]. Notes

The authors declare no competing financial interest.

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ACKNOWLEDGMENTS This research was funded by the Ministry of Science and Technology, Taiwan under Grant NSC 102-2622-E-011-006CC1.

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REFERENCES

(1) Tsouris, C.; Porcelli, J. V. Process IntensificationHas Its Time Fnally Come? Chem. Eng. Prog. 2003, 99, 50. (2) Annual Energy Outlook 2010; U.S. Energy Information Administration: 2010.

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