Article pubs.acs.org/IECR
Optimization of Pressure-Swing Batch Distillation with and without Heat Integration for Separating Dichloromethane/Methanol Azeotrope Based on Minimum Total Annual Cost Xin Li, Yongteng Zhao, Bin Qin, Xia Zhang, Yinglong Wang, and Zhaoyou Zhu* College of Chemical Engineering, Qingdao University of Science and Technology, Qingdao 266042, China ABSTRACT: Pressure-swing batch distillation was investigated to separate the binary azeotrope of dichloromethane/ methanol. Based on the robust control of a composition−flow rate cascade control structure, minimum total annual cost, instead of conventional energy consumption, was studied to optimize the process. Different cooling mediums of cooling water and refrigerated water were applied and compared based on a 323 K reflux-drum criterion. Heat integration was also studied to realize further reduction of total annual cost. The results show that the difference of the costs is small when the lower-pressure column was operated at 1 and 1.6 atm. However, heat integration reduces the cost by 11.3% and shows advantage in energy consumption, which is 53.8% lower than that of the process without heat integration. types and summarized their features. Sørensen and Skogestad26 compared the regular and inverted BD and found that the inverted BD yields the shortest operating time when there is a small amount of light component in the feed. Gruetzmann et al.27 theoretically analyzed the cyclic operation of middle vessel batch distillation and experimentally investigated the operating behavior, and the proposed process is able to save time and for industrial use. Babu et al.28 studied a novel energy-efficient middle vessel batch distillation (MVBD) and proposed an adaptive heat pumping system which achieves 2.18% savings in total annual cost. Zhao et al.29 compared two operation policies of multivessel batch distillation and found that optimal operation policy is better than regular constant reflux operation in time savings and operational simplicity. However, when azeotropes need to be separated to high purities, common BD processes perform poorly and should be improved. Pressure-swing batch distillation (PSBD) is explored to separate azeotropes in relevant industries.30 Modla and Lang31 proposed new PSBD methods to produce components simultaneously with heat integration. Klein and Repke32 investigated regular and inverted modes of PSBD to separate acetonitrile/water system, and the regular batch process with a big reboiler requires the least time and energy. Modla and Lang33 introduced a double-column PSBD process to separate a pressure-sensitive azeotropic mixture and compared the effects of operational parameters on the process under the situation with and without thermal integration. These studies
1. INTRODUCTION Dichloromethane/methanol mixture, whose components are great solvents, is a common binary azeotrope in pharmaceutical processes such as prednisone production and antibiotic synthesis.1 The wastewater discharged from the processes contains large amounts of the two components. The emission of dichloromethane jeopardizes both human health and the environment.2,3 As for methanol, it is an important carbon feedstock in the chemical industry4 and can be used as a convenient fuel;5 therefore, the recycle of the component is of great economy. Hence, it is meaningful and important to separate the binary azeotrope and reuse the components. Distillation receives heavy emphasis when separating mixtures.6 Pressure-swing distillation (PSD), as a special distillation process, is generally used to separate different types of azeotropes such as maximum-boiling azeotropes,7,8 minimum-boiling azeotropes,9−12 and the azeotropes which present both of the above azeotropic behaviors with pressure changes.13 The process has the advantage of avoiding the third component and can implement different heat-integrated methods.14−17 Bessa et al.18 studied double-effect integration of multicomponent alcoholic distillation columns and reduced the specific steam consumption. Heat-integrated methods are also applied in other distillation processes. Kiran et al.19 applied the pinch technology to identify the feasible range for thermal coupling, and the scheme showed lower energy consumption and better economic figures. Maiti et al.20 introduced a thermally integrated batch distillation with chemical reaction, and the proposed process provided an energy saving of 37.08%. Batch distillation (BD) is widely applied in the pharmaceutical and fine chemical industries because of its great flexibility.21−23 BD has been developed for some new processes.20,24 Diwekar25 classified the BD processes into six © XXXX American Chemical Society
Received: Revised: Accepted: Published: A
February 1, 2017 March 5, 2017 March 22, 2017 March 22, 2017 DOI: 10.1021/acs.iecr.7b00464 Ind. Eng. Chem. Res. XXXX, XXX, XXX−XXX
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Industrial & Engineering Chemistry Research promote the development of PSBD processes and lay a solid foundation for further research. Robust control strategies are essential for batch distillation and are meaningful for achieving high-purity products. Luyben34 designed the MVBD simulation process and explored a composition control structure to separate a benzene/toluene/ xylene ternary system. Zhu et al.35 compared two control structures of MVBD and found that temperature control performs better than composition control. As for the control structures of PSBD process, Repke et al.36 used level control and composition control to investigate the regular batch and inverted batch and found that inverted batch requires less time. Kopasz et al.37,38 presented simple composition control schemes for the double-column batch rectifier process and double-column batch stripper process. The vapor division ratio and the liquid division ratio were optimized to implement the minimum specific energy consumption. The high-purity products in the two processes were obtained by good controllability of the control systems. The studies make contributions to the control of PSBD. All of the above works make great contributions to the development of PSBD. Total annual cost (TAC), however, is less found in the process optimization as the objective function. In this paper, we introduce a composition−flow rate cascade control structure into the PSBD for separating dichloromethane/methanol binary azeotropic mixture by use of the AspenTech simulation platform. Based on the 323 K refluxdrum criterion,39 different operating pressures of the lowpressure column (LPC) and high-pressure column (HPC) were investigated to minimize the TAC. Heat integration was also studied to realize further cost and energy savings.
Figure 2. T−xy diagrams of dichloromethane/methanol azeotrope at 1 and 11 atm.
azeotropic mixtures are withdrawn as distillate back to the feed vessel. (2) When the initial mole fraction of dichloromethane in the feed vessel is greater than 87.1 mol %, the products from the bottoms of both columns are dichloromethane, and the azeotropic mixtures are withdrawn as distillate to the feed vessel. (3) When the initial mole fraction of dichloromethane in the feed vessel is between 70.3 and 87.1 mol %, the product from the bottom of the LPC is methanol and that of the HPC is dichloromethane, and the azeotropic mixtures are also withdrawn as distillate to the feed vessel. The operating pressures of the two columns must be determined by the rule that the actual feed composition should be located between the two azeotropic compositions so that the PSBD could separate the mixture efficiently. In addition, different operating pressures lead to the difference of TAC; therefore, the pressures need to be optimized. There is another column configuration called double-column batch rectifier (DCBR) which is quite different from DCBS (Figure 1b). DCBR is not suitable for separating minimum azeotrope like the dichloromethane/methanol system in this paper because when the initial mixture is vaporized and fed into the bottom of the two rectifiers, the azeotrope, which has the lowest boiling point temperature, is distilled as products at the top and pure components are sent back to the feed vessel. The azeotrope cannot be separated by the DCBR. Hence, this configuration will not be considered in the rest of the paper.
2. FEASIBILITY OF PSBD FOR SEPARATING DICHLOROMETHANE/METHANOL MIXTURE Dichloromethane and methanol form minimum pressuresensitive azeotrope whose azeotropic composition varies with pressure changing so that it is possible to separate the azeotrope by PSBD. Two column configurations for PSBD process are shown in Figure 1. Double-column batch strippers
Figure 1. Pressure-swing batch column configurations: (a) DCBS and (b) DCBR.
(DCBSs) were applied to implement the separation in this work, and the flowsheet is shown in Figure 1a. Figure 2 shows the T−xy diagram for the dichloromethane/methanol mixture at 1 and 11 atm. There are three feed conditions, and the corresponding processes are as follows: (1) When the initial mole fraction of dichloromethane in the feed vessel is less than 70.3 mol %, the products from the bottom of the both columns are methanol, and the
3. TAC MODEL OF PSBD TAC that includes capital cost and operating cost is an essential index to judge whether the process is economical for separating the mixtures. The TAC models used for the PSBD are shown in B
DOI: 10.1021/acs.iecr.7b00464 Ind. Eng. Chem. Res. XXXX, XXX, XXX−XXX
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Industrial & Engineering Chemistry Research Figure 3, and the analytical expressions of the TAC with a three-year payback period are as follows:
Table 1. Basis of Economics and Equipment Sizing
Figure 3. TAC of PSBD.
TAC =
Capital cost + Operating cost Payback period
(1)
Capital cost = Column vessel cost (C1) + Plate cost (C2) + Heat exchangers cost (C3)
(2)
Operating cost = Annual steam cost (C4) + Annual cooling water cost (C5)
(3)
The capital cost includes the costs of column vessels, plates, and heat exchangers. The column parameters were calculated by “Tray Sizing” function in Aspen Plus. The detailed calculation methods of sizing parameters and price of each part provided by Douglas40 are shown in Table 1. The mixtures are set at 450 batches per year. Additional costs, such as valves, pumps, and pipes, are much lower than those of column vessels, trays, and heat exchangers and are ignored. It should be noted that refrigerated water was used in this paper, and its price was taken from Turton et al.41
4. SIMULATION PROCESSES AND CONTROL STRUCTURE DESIGN 4.1. PSBD Process Design. In the steady-state simulation with UNIQUAC property method, 100.0 kmol binary mixture with the composition of 83.0 mol % dichloromethane and 17.0 mol % methanol was fed into the feed pot. Two columns operated at different pressures were set to separate the binary azeotrope using the Radfrac model (Figure 4). The mixture was fed into the top of the two columns simultaneously, and the high-purity methanol and dichloromethane were withdrawn from the bottom of the LPC and HPC, respectively. Green and Perry42 concluded that when the difference of the two azeotropic compositions is greater than 5.0 mol % and the pressure difference is lower than 10 atm, the pressure-swing process can be economical. Based on the above conclusion and feed composition, the operating pressure of LPC was set to 1 atm initially, and that of HPC was selected between 3 and 15 atm. The feed pot was operated at 1 atm as well as the methanol vessel (V-LP) and dichloromethane vessel (V-HP), and their volumes were set to 7.69, 0.76, and 6.85 m3, respectively. The stage number of the columns was set to 30, which is large enough to separate the components to 99.8 mol %, and the column diameter was calculated by the “Tray Sizing” function. The pumps and valves were used to change the
streams’ pressure because the dynamic simulation is based on a pressure driven mode. 4.2. Design of Control Structures. 4.2.1. Initial Dynamic Process. Two pressure controllers were set to maintain the operating pressures in the columns by manipulating the heat duty of condensers (Figure 4). Figure 5 shows that the methanol purity in the V-LP vessel is up to 99.9 mol %, whereas the dichloromethane purity in the V-HP vessel is only 96.1 mol %. The heat duty of the reboiler in the HPC is 373.0 kW. The flow rate of the stream TO-HP remains at 50.0 kmol/h; therefore, the dichloromethane purity should be up to 99.9 mol % if the composition of the stream TO-HP remains constant. However, the composition of methanol in the FEEDPOT increases from 18.8 to 21.5 mol % gradually, and more methanol is fed into the HPC without good control; therefore, the dichloromethane purity in the V-HP vessel is only 96.1 mol % at the end of the batch. To realize high dichloromethane purity in the vessel, the most important point is that almost all the methanol should be distilled at the top of the HPC with azeotropic composition. C
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Figure 4. Pressure-swing batch distillation flowsheet.
Figure 5. Responses of initial dynamic batch process.
Figure 6. Composition−flow rate cascade control structure of PSBD process.
delivered to the flow rate controllers (FC1, FC2), and the opening of the valves (VF-LP, VF-HP) would decrease. Therefore, the flow rate of the feed stream TO-HP reduces and that of methanol in the stream TO-HP to the HPC decreases when dichloromethane purity decreases because of the direct acting of the composition controller. Similarly, the flow rate of the feed stream TO-HP and that of methanol to the HPC increase with the dichloromethane purity increasing. As a result, the product purity of dichloromethane is ensured to reach a high level under the control of composition−flow rate cascade control structure. The parameters of the controllers are shown in Table 2. The tuning parameters of pressure controllers were defaults, and those of flow rate controllers were set empirically as in the existing works.43−45 As for the composition controllers, the parameters were determined by trial- and-error method46 until the controllers satisfied the goals of product purity.
4.2.2. Composition and Flow Rate Cascade Control. Composition controllers and flow rate controllers were cascaded to achieve robust control on the product purity (Figure 6). Figure 7 shows the Aspen flowsheet of the PSBD process with the cascade control structure. The detailed control structure is as follows: (1) The heat duty of the condensers in the columns was manipulated (reverse acting) to control the operating pressure of the columns. (2) The flow rate of the feed streams to the columns was manipulated (direct acting) to control the composition of products at the bottom. When the product purities at the bottom of the columns were lower than the set point of composition controllers (CC1, CC2), the output signal, which reduced the flow rate of the feed stream (TO-HP) to the columns (LPC, HPC), was D
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Figure 7. ASPEN flowsheet of the PSBD process with cascade control structure.
5. PROCESS OPTIMIZATION BASED ON TAC MINIMIZATION 5.1. LPC with Operating Pressure at 1 atm. The LPC was operated at 1 atm, and the HPC operated in the pressure range of 3−15 atm was investigated to optimize the process by minimizing the TAC. The azeotropic temperatures of dichloromethane/methanol binary system are changed from 37.4 to 134.8 °C; therefore, the operating cost is quite different under various pressures when the heating steam and cooling water were taken into consideration. Because of the low temperature, which is 37.4 °C at the top of the LPC, refrigerated water with the temperature of 5 °C was used to condense the vapor efficiently. As for the HPC, the top temperature is at least 70.4 °C so the cooling water can be utilized. Different operating pressures of the HPC cause changes of process parameters, such as tray diameters and compositions of the streams. The composition−flow rate cascade control structure is tested to implement robust control under each condition of pressure change, and the TAC is also recalculated. The calculated TAC under different operating pressures is shown in Figure 9a. When the operating pressure was set to 3 atm in the HPC, the TAC was 374 632 $/y. With the pressure increasing, the TAC decreases quickly to 285 224 $/y when the column was operated at 11 atm. After that, the TAC increases gradually to 291 363 $/y at 15 atm. The minimum TAC was acquired at 1 atm for LPC and 11 atm for HPC. The heating energy required is 13.36 GJ per batch. In addition, the optimized pressure difference of the two columns is 10 atm, and the results in this paper extend the application range of pressure differences compared with Green and Perry’s work.42 5.2. LPC with Operating Pressure at 1.6 atm. Refrigerated water was used to condense the top vapor of LPC. However, the refrigerated water is expensive, and its price is 4.43 $/GJ, which is much higher than 0.354 $/GJ of cooling water; therefore, other condensing methods should be tried to decrease the TAC. The boiling point of components or azeotropes would increase with increasing pressure, and this feature could be used to reduce the cost. According to the 323 K reflux-drum criterion, the operating pressure of LPC was set to 1.6 atm; the azeotropic boiling point is 50.6 °C (323.6 K)
Table 2. Parameters of the Controllers in PSBD parameter
PC1
FC1
CC1
PC2
FC2
CC2
gain, Kc integral time, τI (min)
10 12
0.5 0.3
10 20
10 12
0.5 0.3
10 20
Figure 8 shows the results of the PSBD process with composition−flow rate cascade control structure. The
Figure 8. Results of composition−flow rate cascade control structure.
compositions of methanol and dichloromethane in the vessels are 99.9 and 99.8 mol %, respectively. The variation tendencies of methanol flow rate in the stream TO-HP and dichloromethane purity in the V-HP vessel are similar because of the control of the composition−flow rate cascade structure. The control structure could control the PSBD process robustly to separate the binary azeotrope. The operating pressures impact the capital cost and operating cost directly, and the azeotropic compositions influence the controllability of the control structure because the difference of the initial feed composition and azeotropic compositions are not constant. However, the composition− flow rate cascade control structure realizes robust control with the operating pressure changes. E
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LPC was operated at 1.6 and 1 atm, and the difference is only 0.57%. The TAC is composed of operating cost and capital cost. When the operating pressure is increased, the operating cost and the capital cost change. Then, the TAC as the summation of the two costs changes and has potential to obtain a minimum value. For example, when the LPC is operated at 1 atm and the operating pressure of the HPC changes from 9 to 13 atm, the operating cost decreases from 80 246 $/y to 62 233 $/y while the capital cost increases from $640 245 to $673 933 (Figure 10). Therefore, the minimum TAC of 285 224 $/y is obtained at 11 atm. 5.3. Heat Integration in the PSBD. Heat integration is an efficient method for using the redundant energy of the processes themselves, and then the TAC can be reduced. In the PSBD process, the different operating pressures of the two columns make it possible to implement heat integration. The temperature of the vapor at the top in HPC is high enough to heat the liquid partially or fully at the bottom of the LPC in the reboiler. Partial heat integration with an auxiliary reboiler was investigated in this paper. The process is shown in Figure 11.
Figure 9. TAC variation under different operating pressures of HPC when LPC is operated at (a) 1 atm and (b) 1.6 atm.
Figure 11. Heat integration process of PSBD.
There are four key parameters to implement the heat integration: the heat duties of condenser in the HPC and reboiler in the LPC, the top temperature of HPC, and the bottom temperature of LPC. Unlike conventional continuous pressure-swing distillation in which the above four factors always remain constant, in the batch process, some of the factors are changing. Hence, the temperature differences and heat duties must be suitable for the whole process, and then heat integration can be applied. Figure 12 shows the variation tendency of the four parameters. The heat duty of the reboiler
under this pressure. Therefore, the cooling water can be applied in the two columns, and it is possible to decrease the TAC. Figure 9b shows the TAC when the LPC was operated at 1.6 atm. The highest TAC is 430 891 $/y at 3 atm in the HPC, and it decreases quickly in the range of 3−11 atm. When the operating pressure in the HPC is greater than 11 atm, the TAC increases gradually. The variation tendency of Figure 9b is similar to that of Figure 9a, and the minimum TAC of 283 594 $/y is acquired at 11 atm. The heating energy required is 19.45 GJ per batch. The minimum TAC is slightly different when the
Figure 10. Relationship of TAC, capital cost, and operating cost. F
DOI: 10.1021/acs.iecr.7b00464 Ind. Eng. Chem. Res. XXXX, XXX, XXX−XXX
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minimum TAC was adopted as the objective function for the first time to optimize the PSBD process instead of the minimum energy consumption. A composition−flow rate cascade control structure was introduced to achieve robust control. The methanol content in the feed stream to the HPC was the key point to ensure the high purity of dichloromethane in the V-HP vessel. As for the purity of methanol in the V-LP vessel, the dichloromethane content in the feed stream to the LPC plays an important role. The composition−flow rate control structure controls the product purities by manipulating the impurity flow rate to the columns, and the purities of separated dichloromethane and methanol are greater than 99.8 mol %. Pressure is an important parameter for the PSBD process. Different operating pressures of HPC were studied to minimize the TAC with the LPC operating at 1 and 1.6 atm. The conventional conclusion is that the pressure-swing process is economical when the pressure difference of the two columns is within 10 atm. We studied a wider operating pressure range from 3 to 15 atm and calculated the TAC. The results show that the minimum TAC was acquired when the pressure of HPC was set to 11 atm. Although the difference of the TAC under the two conditions is slight, the optimized pressure difference extends the range of pressures worthy of consideration. Heat integration was applied to the PSBD process to achieve further TAC reduction. Different from continuous pressureswing distillation, some of the key parameters, such as heat duties of reboiler in the LPC and condenser in the HPC and their temperature difference, change in the batch process. The temperature difference is at least 50 °C, and the heat duty of the condenser is always lower than that of the reboiler during the whole process; therefore, it is feasible to implement heat integration, and an auxiliary reboiler was set. The results show that through application of the energy-saving method to the PSBD process, the minimum TAC was reduced by 11.3%, and the heating energy required was 53.8% lower than that of the process without heat integration.
Figure 12. Variation tendency of heat-integrated parameters.
in the LPC is always greater than that of the condenser in the HPC; therefore, an auxiliary reboiler was set to supply the extra energy. The heat duty changes occurring in the condenser result from the changing feed flow rate to the HPC, which can be seen in Figure 8. The heat duty is one of the criteria for realizing heat integration, and the temperature difference is another important but easily neglected factor. If the difference is too small, the heat-transfer areas will become too larger and lead to the increase of capital cost. The appropriate difference is at least 20 °C so that the heat integration can be efficient. The temperature of the top in the HPC is much higher than that of the bottom in the LPC; therefore, the temperature difference, which is at least 50 °C, is large enough to implement heat integration. The total heat of the condenser in the HPC, which is 7.2 GJ per batch, was used as part of the required total heat of the reboiler in the LPC, which is 9.2 GJ per batch. Therefore, the TAC would be reduced. Table 3 shows the results of the costs with and without heat integration. The capital cost of the heat-integrated process is
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Table 3. Results of the Costs with and without Heat Integration
heat integration no heat integration
capital cost ($)
operating cost ($/y)
total annual cost ($/y)
638 567 657 529
39 925 66 047
252 780 285 224
AUTHOR INFORMATION
Corresponding Author
*E-mail:
[email protected]. ORCID
Yinglong Wang: 0000-0002-3043-0891 Notes
The authors declare no competing financial interest.
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638 567 $/y, and that of the process without heat integration is 657 529 $/y. The operating cost of the heat-integrated process is 39 925 $/y, and that of the process without heat integration is 66 047 $/y; heat integration can reduce the cost obviously. Applying heat integration to the PSBD process for separating the dichloromethane/methanol azeotropic mixture gives a TAC of 252 780 $/y, which is 11.3% lower than that of the PSBD process without heat integration. In addition, the heating energy required, including the heat duty of auxiliary reboiler, is 6.17 GJ per batch, which is lower than that of the process without heat integration by 53.8%.
ACKNOWLEDGMENTS Support from the National Natural Science Foundation of China (Project 21676152) is gratefully acknowledged.
6. CONCLUSIONS The PSBD process was investigated to separate dichloromethane/methanol binary azeotrope. The feasibility of the process was analyzed, and the initial composition in the feed vessel played an important role in the separation. The G
NOTATION AUXREB = auxiliary reboiler BD = batch distillation DCBR = double-column batch rectifier DCBS = double-column batch stripper HPC = high-pressure column LPC = low-pressure column MVBD = middle vessel batch distillation PSBD = pressure-swing batch distillation PSD = pressure-swing distillation TAC = total annual cost DOI: 10.1021/acs.iecr.7b00464 Ind. Eng. Chem. Res. XXXX, XXX, XXX−XXX
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V-HP = dichloromethane vessel V-LP = methanol vessel
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DOI: 10.1021/acs.iecr.7b00464 Ind. Eng. Chem. Res. XXXX, XXX, XXX−XXX
Article
Industrial & Engineering Chemistry Research (46) Robbins, L. Distillation Control, Optimization, and Tuning: Fundamentals and Strategies; CEC Press: New York, 2011.
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DOI: 10.1021/acs.iecr.7b00464 Ind. Eng. Chem. Res. XXXX, XXX, XXX−XXX
