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Effect of Solvent Flow Rates on Controllability of Extractive Distillation for Separating Binary Azeotropic Mixture Yinglong Wang, Shisheng Liang, Guangle Bu, Wei Liu, Zhen Zhang, and Zhaoyou Zhu Ind. Eng. Chem. Res., Just Accepted Manuscript • DOI: 10.1021/acs.iecr.5b03666 • Publication Date (Web): 07 Dec 2015 Downloaded from http://pubs.acs.org on December 15, 2015

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Effect of Solvent Flow Rates on Controllability of Extractive Distillation for Separating Binary Azeotropic Mixture Yinglong Wang,* Shisheng Liang, Guangle Bu, Wei Liu, Zhen Zhang, and Zhaoyou Zhu College of Chemical Engineering, Qingdao University of Science and Technology, Qingdao 266042, China Corresponding Author *E-mail: [email protected]. ABSTRACT: The mixture of n-heptane and isobutanol creates a minimum-boiling azeotrope under atmospheric conditions. Extractive distillation was used for separating n-heptane and isobutanol. A sequential iterative optimization procedure optimized the steady state extractive distillation process to obtain a minimal total annual cost (TAC). The dynamic control of the optimal extractive distillation process showed poor controllability. It was found that the dynamic control performance could be improved if the solvent flow rate was increased properly with a small increase in TAC. Hence, the choice of the optimal extractive distillation process for separating a binary azeotropic mixture should be considered from the perspectives of both TAC and the dynamic control performance. Keywords: n-heptane; isobutanol; TAC; dynamic control; extractive distillation

1. INTRODUCTION The separation of binary azeotropic mixtures is an important process. Special distillation methods such as extractive distillation,1-6 pressure-swing distillation,7-11 and azeotropic distillation12-16 are widely used to separate binary azeotropes. Pressure-swing distillation is useful only if the azeotropic composition has an obvious shift as the pressure changes. Extractive distillation, in which a third component called the solvent is added to achieve an effective separation, is another commonly used method to separate binary azeotropes. Organic solvents are common entrainers in extractive distillation. An appropriate solvent can make extractive distillation easier for separating the binary azeotropic mixture. For example, ethylene glycol was selected as the solvent in the separation of tetrahydrofuran and ethanol;17 dimethyl sulfoxide was used to separate an azeotrope of acetone and chloroform;18 2-methoxyethanol was chosen as the entrainer to separate isopropyl alcohol from diisopropyl ether;19 and water was used as the solvent for acetone-methanol separation.20 Recently, the use of ionic liquids as the entrainers and extractive dividing-wall columns has aroused much research interest.21-24 The principle of extractive distillation involves a high boiling point component, which is added to the separation system to change the relative volatility between the light and heavy components. Finally the azeotrope is separated by two columns; one is called the extractive column and the other the solvent recovery column. Two high-purity products are obtained at the top of the two columns, 1



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and the recovered solvent at the bottom of the solvent recovery column is recycled to the extractive column. In the steady state design of extractive distillation, investment and operating costs are the two critical factors that should be taken into account. Many scholars and researchers have focused on the minimal economic cost of the extractive distillation process. Luo et al.19 investigated the optimal extractive distillation process from an economic standpoint, in the separation of isopropyl alcohol and diisopropyl ether. R. Munoz25 explored the simulation and optimization of separating isobutyl alcohol and isobutyl acetate to obtain a minimal TAC. Bao et al.26 conducted the economic optimization for the separation of trimethyl borate and methanol with a sequential iterative optimization procedure. Wang et al.27 optimized the extractive distillation process for separating a binary azeotrope of ethanol and tetrahydrofuran and obtained the optimal design variables with a minimal TAC. The effective control structure to ensure a robust dynamic control performance is another critical factor in distillation design. Luyben28 provided a detailed description of making an effective control structure with controllers and schemes that involved the fixed reflux ratio and the ratio of reflux flow rate/ feed flow rate（R/F）. Wang et al.4 studied the temperature control for separating methylal and methanol with extractive distillation and found that good dynamic responses can be shown with R/F schemes. Yang et al.29 explored the dynamic control of an extractive distillation system for benzene/acetonitrile separation and concluded that the improved control scheme with fixed reflux-to-feed ratio and the dual temperature control structure were both advantageous in controlling the system. Luyben30 compared the dynamic performance of the acetone/methanol system with different solvents and reported that all the systems were controllable, but product quality variability was poorest when methanol was driven overhead in the extractive column. Wang et al.31 investigated the dynamic control of extractive distillation for separating the azeotropic mixture of ethanol and tetrahydrofuran and found an improved control structure when the solvent flow rate/ feed flow rate (S/F) was controlled. The control structure exhibited good controllability. All of the above mentioned studies promoted the development of extractive distillation. To date, there are no studies investigating the tradeoff between TAC and controllability for an optimal extractive distillation process used in separating binary azeotropes. In this paper, the steady state design and effective dynamic control structure were studied to determine an optimal extractive distillation process for separating an azeotrope of n-heptane and isobutanol from the perspectives of both TAC and controllability. The mixture of n-heptane and isobutanol is generated in the production of triisobutyl vanadate32. The investment costs for the materials can be reduced if the n-heptane and isobutanol mixture can be separated, and the materials can be reused 2


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in the reaction process to produce triisobutyl vanadate. For the steady state design, sequential iterative optimization was used to obtain a minimal TAC. From the point of dynamic control, control structures with different solvent flow rates were evaluated to ensure a proper extractive distillation process.

2. STEADY STATE PROCESS CONSIDERING TAC 2.1. Solvent Selection The determination of the solvent is critical for extractive distillation.30,

33-36

The relative

volatility is a criterion used for solvent selection. Relative volatility is the ratio of volatilities between the light and heavy components after the addition of solvent. A higher relative volatility results in an easier separation. A preliminary screening was performed, and the three solvents NMP (N-methyl-2-pyrrolidone), DMSO (dimethyl sulfoxide) and 1,2-propanediol were studied according to the principle of solvent selection. Figure 1 shows the influence of the different solvents on VLE (vapor liquid equilibrium) at the fixed S/F ratio of 2. The three solvents were able to effectively break the azeotropic phenomenon of n-heptane and isobutanol. The VLE curve with the addition of NMP showed the largest deviation from diagonal, which means that the relative volatility was the highest. The residue curve map (RCM) of the n-heptane/isobutanol/NMP ternary system drawn by Aspen Plus using the UNIQUAC model is shown in Figure 2. From the figure, we can see that both n-heptane and isobutanol were the saddles, while NMP was the stable node, and the azeotrope of n-heptane/isobutanol was the unstable node. The resulting residue curves with arrows pointing to pure NMP indicate the absence of the distillation boundary in the RCM. This is an ideal situation for the selection of an extractive distillation process. The blue line stands for the material balance line. It was observed that F1 separated into D1 and B1, and B1 separated into D2 and B2, which indicated that the entrainer NMP helped the feed to be separated into fairly pure products. The addition of a small supplementary stream of NMP could balance the tiny loss of entrainer in both the D1 and D2 streams.

2.2. Process Design and Economic Analysis The separated mixture was set at 100 kmol/h with 50 mol % n-heptane and 50 mol % isobutanol. The two product specifications were set at a purity of no less than 99.90 mol %. The bottom stream of the solvent recovery column consisted of pure NMP with a trace amount of n-heptane and isobutanol. This was recycled back to extractive column. The condenser pressures of both the extractive column and solvent recovery column were set at 1 atm, with a tray pressure drop of 0.0068 atm. The commercial software Aspen Plus was used to simulate the process. Before the rigorous steady state simulation, several studies were made to find the suitable solvent flow rate. Figure 3 shows the influence of solvent flow rate and reflux ratio (RR1) on the 3



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purity of n-heptane and on the distillation product in the extractive column. The purity of n-heptane initially showed an increasing trend and then decreased with the rise in RR1. This indicated that highest purity could be achieved with an appropriate RR1. To achieve a 99.90 mol % purity, the solvent flow rate should be larger than 95 kmol/h. The sequential iterative optimization procedure19,

26

was used to obtain the optimal design variables including the total stages of

extractive column (NT1), fresh feed tray location (NF1), solvent feed stage (NFE), total stages of solvent recovery column (NT2), fresh tray location (NF2) of the solvent recovery column, and solvent flow rate. The results from Figure 4 indicated an optimal solvent flow rate of 100 kmol/h with a minimal TAC. The flowsheet for the extractive distillation system is shown in Figure 5 with detailed equipment sizes, design variables, stream information, heat duties and operating conditions.

3. DYNAMIC CONTROL OF EXTRACTIVE DISTILLATION PROCESS It is necessary to explore and evaluate the control structure after the steady state design of the process. Before starting a dynamic simulation in the Aspen Plus Dynamics software, the sizes of reflux drums and column sumps were designed to provide a 10 min holdup if full. Pumps and valves were given proper pressure drops to ensure the dynamic operation. Aspen Plus file was exported to Aspen Plus Dynamics after running the pressure checker.

3.1. Dynamic Control of The Optimal TAC Design 3.1.1. Basic Control Structure The basic control structure for the optimal TAC design of extractive distillation with a solvent flow rate of 100 kmol/h is shown in Figure 6. The stage 32 in the extractive column and stage 13 in the solvent recovery column were chosen as temperature sensitive stages, according to the slope criterion of the temperature control stage selection.28, 37-39 The detailed control structure is as follows： (1) The heat removal rate in the condenser of two columns was manipulated (reverse acting) to control the operating pressure of two columns. (2) The total solvent flow rate was in proportion to the feed flow rate. (3) The flow rate of distillation (direct acting) was manipulated to control the reflux drum levels in both columns. (4) The flow rate of the bottom (direct acting) was manipulated to control the base level in the extractive column. (5) The flow rate of the makeup stream (reverse acting) was manipulated to control the base level in the solvent recovery column. (6) The temperature of stage 32 in the extractive column was controlled by manipulating the QR/F ratio (reboiler heat duty /fresh feed flow rate) of the extractive column (reverse acting). 4


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(7) The reboiler heat duty of the solvent recovery column (reverse acting) was manipulated to control the temperature of stage 13 in the solvent recovery column. (8) The cooler HEDC heat duty (reverse acting) was manipulated to hold the temperature of the solvent feed. (9) The reflux ratio of each column was fixed. (10) Fresh feed was flow-controlled (reverse acting). It should be noted that when the solvent flow rate controller was set as reverse acting, the system stopped running when disturbances were introduced. Hence, the controller was set as direct acting, and the system could run without stopping, even with disturbances. All three deadtime blocks were inserted with the deadtime of 1 min, then relay-feedback tests were conducted on the three temperature controllers to determine the ultimate gains and periods. The parameters of the three temperature controllers using Tyreus-Luyben tuning40 are shown in Table 1. All the four level controllers were given gains (Kc) of 2 and integral times (τI ) of 9999 min, and the flow rate controllers had Kc = 0.5 and τI = 0.3 min. To assess the dynamic control performances of the basic control structure, a feed flow rate of ±20% and composition disturbances were introduced to the control system. The corresponding dynamic responses are shown in Figure 7. The purities of n-heptane and isobutanol returned to their specified values approximately 3 h after ±20% feed flow rate disturbances and -20% composition disturbances occurred at 0.5 h. When faced with +20% composition disturbances, the purity of n-heptane in the extractive column showed a value of 99.67 mol % at the new steady state. This was a large deviation compared with the specified purity of 99.90 mol %. Hence, it can be concluded that the basic control structure cannot handle +20% composition disturbances well. 3.1.2. Improved Control Structure In this section, an improved control structure to control the extractive distillation system is presented. Basic control structure schemes such as R/F41, reflux ratio controlling the temperature of temperature sensitive stages, and R/F controlling the temperature of temperature sensitive stages were attempted but failed to improve the controllability. Finally, a composition controller was used to detect the purity of n-heptane at the top of the extractive column, and the output signal of the composition controller was the S/F ratio. The improved control structure of the optimal design is shown in Figure 8, and the parameters of the composition controller are shown in Table S1. The detailed improved control structure was the same as the basic control structure except for a composition controller. The composition controller was set as reverse acting. This improved control structure handled both ±20% feed flow rate and feed composition disturbances well, and the purity of n-heptane and isobutanol returned to their specified values as shown in Figure 9. The control stages’ temperatures of both columns were brought back to their 5



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set points at the new steady state. However, when faced with a ±20% feed composition disturbance, the improved control structure took approximately 15 h to return back to its desired values. The improved control structure using a composition controller is not the best choice for the extractive distillation system, due to the high installation cost and the yearly maintenance expenses.

3.2. Dynamic Control of the Process with a Modified Solvent Flow Rate Although, adding a composition controller effectively handled the ±20% feed flow rate and feed composition disturbances, the improved control structure with the composition controller had some shortcomings, such as high costs and long time delays to reach the steady state. Hence, we further investigated obtaining better controllability at the cost of increasing a small TAC. To this end, we changed some design variables of the economic optimal design to obtain better control performance. In this section, the dynamic control performances with an increased solvent flow rate of 110, 120 and 130 kmol/h were investigated. The dynamic control performance for the optimal design flow rate of 100 kmol/h was compared with those of the increased flow rates. The other design variables NT1, NF1, NFE, NT2, and NF2 were fixed at the same values as that for the optimal design flow rate of 100 kmol/h. The control structures of the three processes with higher flow rates were the same as that of optimal design flow rate of 100 kmol/h. The temperature controller parameters with solvent flow rates of 110 and 120 kmol/h are shown in Table S2. The temperature controller parameters and the steady state flowsheet for the solvent flow rate of 130 kmol/h are shown in Table 2 and Figure 10, respectively. The control structure with a solvent flow rate of 130 kmol/h as shown in Figure 11 was the same as the optimal basic control structure, except that stage 12 was selected as the temperature control stage of the solvent recovery column. This is due to the steep slope at stage 12. To investigate the dynamic control performances of the process with modified solvent flow rate, ±20% feed flow rate disturbances and composition disturbances were introduced in to the control systems. The dynamic responses with solvent flow rates of 110 and 120 kmol/h are shown in Figure S1 and S2, respectively. The dynamic response with a solvent flow rate of 130 kmol/h is shown in Figure 12. The control structure addressed both the feed flow rate and feed composition disturbances well. Product purities were held fairly close to their desired values at the new steady state at approximately 3 h, and the two controlled tray temperatures were brought back to their set points. The purities of n-heptane and isobutanol experienced small deviations but quickly returned to their specified values.

4. PROCESS COMPARISONS In this section, we compared the basic control structure with a solvent flow rate of 100 kmol/h (CSA), the improved control structure with a solvent flow rate of 100 kmol/h (CSB), and the basic 6


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control structure with a solvent flow rate of 130 kmol/h (CSC). It can be seen from Figure 13 that the CSA could not maintain the purity of n-heptane with a deviation at the new steady state (99.67% vs 99.90%) when +20% feed composition disturbance is introduced. The CSB and the CSC maintained the purity of both products at acceptable levels after several oscillations when ± 20% feed flow rate and feed composition disturbances were introduced. The CSB required a longer time of 15 h to reach the new steady state, while the CSC reached a new steady state at 3 h in maintaining the purity of n-heptane when ±20% feed flow rate and feed composition disturbances were introduced. Although the CSC experienced larger transient deviations than CSB, the transient deviations were at acceptable levels and could quickly return to their desired values when ±20% feed flow rates and feed composition disturbances were introduced. In industrial production, ±20% feed flow rate and feed composition disturbances are quite large. Hence, we reduced the disturbances to ±10% feed flow rate and feed composition disturbances, and the corresponding dynamic response are shown in Figure 14. When faced with ±10% feed flow rate disturbances, the CSB and the CSC both handled the disturbances well, with the purities of n-heptane and isobutanol held fairly close to their desired values at the new steady state with similar transient deviations. When ± 10% feed composition disturbances were introduced, the CSB took a longer time (of approximately 15 h) to reach the new steady state in maintaining the purities of n-heptane. In comparison to the CSC, the CSB had larger transient deviations in maintaining the n-heptane product purity. Hence, the CSC showed better controllability than CSA and CSB when faced with feed flow rate and feed composition disturbances.

5. CONCLUSIONS Extractive distillation for the separation of the azeotropic mixture of n-heptane and isobutanol was studied in this paper. According to the criterion of solvent selection, NMP was selected as a suitable solvent. Then, the design and optimization were conducted with the minimal TAC as the objective function. The optimal design of the extractive distillation process was ensured with a total of 36 stages for the extractive column and 18 stages for the solvent recovery column. The two columns were operated under atmospheric conditions. The optimal solvent flow rate was determined as 100 kmol/h with a TAC of 108,432.8 $/y. The dynamic control performances of the optimal steady state design were explored using the Aspen Plus Dynamics software. Feed flow rate and composition disturbances of ±20% were introduced, and it was found that the CSA could not handle the +20% composition disturbance well. The CSB kept the product purities at their desired values but took a long time (of approximately 15 h) to reach the new steady state when±20% composition disturbances were introduced. Dynamic responses indicated that the control structure with modified solvent flow rate showed excellent controllability. In the 7



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separation of isobutanol and n-heptane, the CSC had advantages in terms of the shorter time to reach the new steady state and smaller transient deviations when faced with feed flow rate and composition disturbances. However, this improved performance came at the cost of a TAC increase of 7.66%. Hence, in the design of an extractive distillation process for separating a binary azeotropic mixture, it is essential to conduct a study about the tradeoffs between the minimal TAC and the controllability. This paper reveals that better control performance can be obtained with a modified solvent flow rate following a small increase in TAC. This was confirmed for the example of the separation of the binary azeotrope of isobutanol and n-heptane. More examples are needed to explore and confirm the conclusion. Further studies should be conducted to analyze the reasons.

ASSOCIATED CONTENT Supporting Information Dynamic responses with solvent flow rate of 110 and 120 kmol/h (Figures S1 and S2), composition controller parameters of the improved control structure (Table S1), and temperature controller parameters with solvent flow rate of 110 and 120 kmol/h (Table S2).

AUTHOR INFORMATION Corresponding Author E-mail: [email protected]. Notes The authors declare no competing financial interest.

ACKNOWLEDGEMENTS Editing services of Kunal K. and Marcie J. from the ACS ChemWorx Authoring Services are gratefully acknowledged. Financial supports from the National Natural Science Foundation of China (Project 21306093) and Project of Shandong Province Higher Educational Science and Technology Program (Project J13LD16) are gratefully acknowledged.

NOTATION TAC = total annual cost R/F = reflux flow rate/feed flow rate NMP = N-methyl-2-pyrrolidone DMSO = dimethyl sulfoxide VLE = vapor liquid equilibrium RCM = residue curve map RRn = reflux ratio of column n 8


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NT1 = total stages of extractive column NT2 = total stages of solvent recovery column NF1 = fresh feed tray location NF2 = fresh tray location of the solvent recovery column NFE = solvent feed stage QR/F = reboiler heat duty /fresh feed flow rate TC1, TC2 = temperature controller of column 1, 2 HEDX = temperature controller of recycle solvent KC = gain of controller τI = integral time of controller CSA = basic control structure with solvent flow rate of 100 kmol/h CSB = improved control structure with solvent flow rate of 100 kmol/h CSC = basic control structure with solvent flow rate of 130 kmol/h

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dehydration system via heterogeneous azeotropic distillation. Chem. Eng. Sci. 2004, 59, 4547-4567. (13) Luyben, W. L. Control of the heterogeneous azeotropic n-butanol/water distillation system. Energy. Fuels. 2008, 22, 4249-4258. (14) Foucher, E. R.; Doherty, M. F.; Malone, M. F. Automatic Screening of Entrainers For Homogeneous Azeotropic Distillation. Ind. Eng. Chem. Res. 1991, 30, 760-772. (15) Widagdo, S.; Seider, W. D. Journal review. Azeotropic distillation. AIChE J. 1996, 42, 96-130. (16) Pham, H. N.; Doherty, M. F. Design and synthesis of heterogeneous azeotropic distillations—II. Residue curve maps. Chem. Eng. Sci. 1990, 45, 1837-1843. (17) Tang, K.; Bai, P.; Huang, C.; Liu, W. Separation of Tetrahydrofuran-Ethanol Azeotropic Mixture by Extractive Distillation. Asian J. Chem. 2013, 25, 2774. (18) Luyben, W. L. Control of the Maximum-Boiling Acetone/Chloroform Azeotropic Distillation System. Ind. Eng. Chem. Res. 2008, 47, 6140-6149. (19) Luo, H.; Liang, K.; Li, W.; Li, Y.; Xia, M.; Xu, C. Comparison of Pressure-Swing Distillation and Extractive Distillation Methods for Isopropyl Alcohol/Diisopropyl Ether Separation. Ind. Eng. Chem. Res. 2014, 53, 15167-15182. (20) Luyben, W. L. Comparison of Extractive Distillation and Pressure-Swing Distillation for Acetone-Methanol Separation. Ind. Eng. Chem. Res. 2008, 47, 2696-2707. (21) Meindersma, G. W.; Podt, A. J.; de Haan, A. B. Selection of ionic liquids for the extraction of aromatic hydrocarbons from aromatic/aliphatic mixtures. Fuel Process. Technol. 2005, 87, 59-70. (22) Orchilles, A. V.; Miguel, P. J.; Vercher, E.; Martínez-Andreu, A. Ionic liquids as entrainers in extractive distillation: isobaric vapor-liquid equilibria for acetone+ methanol+ 1-ethyl-3-methylimidazolium trifluoromethanesulfonate. J. Chem. Eng. Data. 2007, 52, 141-147. (23) Bravo-Bravo, C.; Segovia-Hernández, J. G.; Gutiérrez-Antonio, C.; Durán, A. L.; Bonilla-Petriciolet, A.; Briones-Ramírez, A. Extractive Dividing Wall Column: Design and Optimization. Ind. Eng. Chem. Res. 2010, 49, 3672-3688. (24) Kiss, A. A.; David, J.; Suszwalak, P. Enhanced bioethanol dehydration by extractive and azeotropic distillation in dividing-wall columns. Sep. Purif. Technol. 2012, 86, 70-78. (25) Munoz, R.; Monton, J.; Burguet, M.; de la Torre, J. Separation of isobutyl alcohol and isobutyl acetate by extractive distillation and pressure-swing distillation: Simulation and optimization. Sep. Purif. Technol. 2006, 50, 175-183. (26) Bao, Z.; Zhang, W.; Cui, X.; Xu, J. Design, Optimization and Control of Extractive Distillation for the Separation of Trimethyl Borate–Methanol. Ind. Eng. Chem. Res. 2014, 53, 14802-14814. (27) Wang, Y.; Cui, P.; Ma, Y.; Zhang, Z. Extractive distillation and pressure‐swing distillation for THF/ethanol separation. J. Chem. Technol. Biotechnol. 2014. 90, 1463-1472. (28) Luyben, W. L. Distillation Design and Control Using Aspen Simulation; Wiley: New York, 2013. (29) Yang, S.; Wang, Y.; Bai, G.; Zhu, Y. Design and Control of an Extractive Distillation System for Benzene/Acetonitrile Separation Using Dimethyl Sulfoxide as an Entrainer. Ind. Eng. Chem. Res. 2013, 52, 13102-13112. (30) Luyben, W. L. Effect of Solvent on Controllability in Extractive Distillation. Ind. Eng. Chem. Res. 2008, 47, 4425-4439. 10


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(31) Wang, Y.; Zhang, Z.; Zhao, Y.; Liang, S.; Bu, G. Control of Extractive Distillation and Partially Heat-Integrated Pressure-Swing Distillation for Separating Azeotropic Mixture of Ethanol and Tetrahydrofuran. Ind. Eng. Chem. Res. 2015, 54, 8533–8545. (32) Magee Jr, W. L. Method for preparation of alkyl vanadates. U.S. Patent, 4,351,775, Sep 28, 1982. (33) Van Dyk, B.; Nieuwoudt, I. Design of Solvents for Extractive Distillation. Ind. Eng. Chem. Res. 2000, 39, 1423-1429. (34) Xu, S.; Wang, H. A new entrainer for separation of tetrahydrofuran–water azeotropic mixture by extractive distillation. Chem. Eng. Process. 2006, 45, 954-958. (35) Prausnitz, J.; Anderson, R. Thermodynamics of solvent selectivity in extractive distillation of hydrocarbons. AIChE J. 1961, 7, 96-101. (36) Nieuwoudt, I.; Van Dyk, B. In the presence of an extractive distillation solvent comprising an amine, an alkylated thiopene, and paraffins. U.S. patent, 6,375,807, Apr 23, 2002. (37) Zhu, Z.; Wang, L.; Ma, Y.; Wang, W.; Wang, Y. Separating an azeotropic mixture of toluene and ethanol via heat integration pressure swing distillation. Comput. Chem. Eng. 2015, 76, 137-149. (38) Li, W.; Shi, L.; Yu, B.; Xia, M.; Luo, J.; Shi, H.; Xu, C. New Pressure-Swing Distillation for Separating Pressure-Insensitive Maximum Boiling Azeotrope Via Introducing a Heavy Entrainer: Design and Control. Ind. Eng. Chem. Res. 2013, 52, 7836-7853. (39) Yu, B.; Wang, Q.; Xu, C. Design and Control of Distillation System for Methylal/Methanol Separation. Part 2: Pressure Swing Distillation with Full Heat Integration. Ind. Eng. Chem. Res. 2012, 51, 1293-1310. (40) Luyben, W. L. Tuning Proportional-Integral-Derivative Controllers for Integrator/Deadtime Processes. Ind. Eng. Chem. Res. 1996, 35, 3480-3483. (41) Luyben, W. L. Plantwide dynamic simulators in chemical processing and control; Marcel Dekker: New York, 2002.

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Table 1. Transmitter ranges, controller output ranges, and tuning parameters of three temperature controllers of optimal design. parameters

TC1

TC2

HEDC

controlled variable

T1,32

T2,13

T recycle

manipulated variable

QR/F

QR2

QHX

transmitter range (K)

273.2-504.6

273.2-604.5

273.2-366.8

controller output range

0-0.14 GJ/kmol

0-10.03 GJ/h

-7.15-0 GJ/h

gain Kc

1.46

0.34

0.08

integral time τI (min)

10.56

13.2

5.28

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Table 2. Transmitter ranges, controller output ranges, and tuning parameters of three temperature controllers with solvent flow rate of 130 kmol/h. parameters

TC1

TC2

HEDC

controlled variable

T1,32

T2,12

T recycle

manipulated variable

QR/F

QR2

QHX

transmitter range (K)

273.2-507.1

273.2-571.4

273.2-366.8

controller output range

0-0.15 GJ/kmol

0-10.84 GJ/h

-9.21-0 GJ/h

gain Kc

0.99

0.47

0.08

integral time τI (min)

14.52

13.2

5.28

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Figure 1. The influence on VLE with different solvents at the fixed S/F of 2.

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Figure 2. The residue curve map (mole basis) of the n-heptane/isobutanol/NMP ternary system

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Figure 3. Effect of RR1 and solvent flow rates on n-heptane content in distillate product of extractive column.

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Figure 4. The influence of solvent flow rates on TAC.

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Figure 5. The optimal flow sheet with solvent flow rate of 100 kmol/h.

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Figure 6. The basic control structure for the optimal design.

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Figure 7. Dynamic performances of the basic control structure for the optimal design (a) feed flow rate disturbances (b) feed composition disturbances.

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Figure 8. The improved control structure of the optimal design.

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Figure 9. Dynamic performances of the improved control structure for the optimal design (a) feed flow rate disturbances (b) feed composition disturbances.

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Figure 10. The flow sheet with solvent flow rate of 130 kmol/h.

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Figure 11. The basic control structure for the extractive distillation with solvent flow rate of 130 kmol/h.

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Figure 12. Dynamic performances with solvent flow rate of 130 kmol/h (a) feed flow rate disturbances (b) feed composition disturbances.

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Figure 13. Comparison of the dynamic responses between solvent flow rate of 100 and 130 kmol/h (a) ±20% feed flow rate disturbances (b) ±20% feed composition disturbances.

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Figure 14. Comparison of the dynamic responses between solvent flow rate of 100 and 130 kmol/h (a) ±10% feed flow rate disturbances (b) ±10% feed composition disturbances.

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