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Control of Extractive Distillation and Partially Heat-Integrated Pressure-Swing Distillation for Separating Azeotropic Mixture of Ethanol and Tetrahydrofuran Yinglong Wang,* Zhen Zhang, Yujun Zhao, Shisheng Liang, and Guangle Bu College of Chemical Engineering, Qingdao University of Science and Technology, Qingdao 266042, China S Supporting Information *

ABSTRACT: Dynamic controls of an extractive distillation (ED) process and a partially heat-integrated pressure-swing distillation (PHIPSD) process for separating a mixture of 50 mol % ethanol and 50 mol % tetrahydrofuran are investigated. Comparisons between the control structures of the two distillation processes are made. The results show that PHIPSD has advantages of economic savings and control compared with the ED process. In addition, some control structures of PHIPSD for feed streams with different composition and some comparisons between the performances of the control structures are also studied. It is concluded that the feed composition mainly impacts the dynamic control structures of PHIPSD through the different combinations of auxiliary heat exchanger and that a composition controller can improve the control performances.

1. INTRODUCTION Extractive distillation (ED) and pressure-swing distillation (PSD) are commonly used for the separation of azeotropic mixtures.1−3 In our previous paper,4 the separations of a binary azeotrope of ethanol and tetrahydrofuran (THF) were investigated by means of ED and partially heat-integrated pressure-swing distillation (PHIPSD). However, the dynamic controls of ED and PHIPSD for separating the mixture were not investigated. For further understanding of two different separation processes from the dynamic point of view, the dynamic control schemes of ED and PHIPSD will be studied on the basis of steady-state simulation results. Dynamic control has received a lot of researchers’ interest for its practical application in industry, and many achievements have been made to date.5−8 Simulators permitting the quantitative assessment of the effectiveness of different control structures are now available. Simulation results provided by the simulators give an efficient way of comparing control structures, and tuning rules are less expensive and less time-consuming than those from experiments in plant tests.9 Thus, many researchers give dynamic control schemes of distillation processes for separating azeotropes using the simulators based on the reliable thermodynamic data of the components between mixtures.10−15 For example, Luyben,10 Li and Bai,12 and Wei et al.,13 used the simulators of Aspen Plus and Aspen Plus Dynamics to simulate different distillation processes, such as ED, PSD, and heterogeneous azeotropic distillation. In addition, many researchers are aware of the importance of tuning rules for controllers on the dynamic control performances. Thus, tuning rules for the controllers are also developed in the field of automation in addition to the studies of the control structures of the separation processes.16,17 For example, Pavkovic et al.17 presented a new tuning method for the PID controllers for higher-order aperiodic processes, which aims at step response-based autotuning applications. The combination of the simulators with these tuning rules would result in vast improvements in the predictabilities for the control structures. © XXXX American Chemical Society

ED is widely used for separating close boiling point mixtures or azeotropes by introducing a solvent (entrainer) into the system to alter the relative volatilities of the key components, which drives one component overhead to become a distillate product with high purity.18−20 For the binary azeotropes in which the component exhibits pressure-insensitivity or with the relative volatilities of the key components below 1.1, the ED process is more challenging.21,22 The application of the ED process to practical industry is affected by the effective control scheme corresponding to the specific azeotrope system. To date, many researchers have focused on making designs and controls of the ED processes, and various laudable accomplishments have been reported.23−27 For instance, Gil et al.24 gave the design and control strategies for an ED process for producing dehydrate ethanol using glycerol as entrainer. Segovia-Hernández and coworkers26 proposed and simulated a cryogenic ED for ethane and carbon dioxide azeotrope separation using a multiobjective stochastic optimization procedure. Apart from the efforts made on the exploration of control structures, the influences of controlled variables on dynamic performances have also been investigated.28−30 These achievements will promote the development of ED for separating various kinds of azeotropes and facilitate the later to make further investigation on the ED process. PSD, which utilizes the property of binary mixture of azeotropic composition shifting with pressure, is another process frequently used for the separation of azeotropic mixtures because of its superiorities of simple configuration and no third solvent being introduced to the system.31−34 A means of energy saving of the PSD process by combining the condenser of the highpressure column (HPC) with the reboiler of the low-pressure column (LPC) make this distillation process more competitive than conventional distillation.35,36 The dynamic control published Received: May 2, 2015 Revised: July 16, 2015 Accepted: August 14, 2015

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DOI: 10.1021/acs.iecr.5b01642 Ind. Eng. Chem. Res. XXXX, XXX, XXX−XXX

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Industrial & Engineering Chemistry Research

Figure 1. (a) Extractive distillation column temperature profile, (b) entrainer recovery column temperature profile, (c) extractive distillation column liquid composition profile, and (d) entrainer recovery column liquid composition profile.

structure of the ED process was reported relative to the PSD process. In this work, the dynamic control schemes of the ED and PHIPSD processes for separating an azeotropic mixture of ethanol and THF are given, respectively. The final effective control structures are presented, and the comparisons between the dynamic performances of the two processes are made as well. Furthermore, we provide the control structures of the PHIPSD processes handling different feed streams with 10 mol % ethanol and 90 mol % ethanol in this work.

in the open literature includes the control structures for the PSD with no, partial, and full heat integration.37−39 For example, Luyben37 presented some dynamic control structures for PSD with different heat integration in the separation of THF and water mixture. Yu et al.38 provided an improved control structure with the reflux ratio of the HPC adjusted by a pressurecompensated temperature control structure, which can work well for fully heat-integrated PSD for separating a mixture of methylal and methanol. In addition, the design and control structure for a new PSD process for separating pressure-insensitive maximum boiling azeotrope were shown by Li et al.40 The control schemes of PSD with partial and full heat integration are more complicated than that of PSD without heat integration because of the different controllable degrees of freedom.37 The engineers should make some trade-offs between the energy savings and the complexity of the PSD processes with different heat integration. Comparisons between ED and PSD were reported based on the steady-state simulation results and/or the dynamic control performances.41−43 For example, Luyben41 compared the steadystate design and dynamic control of ED and PSD with and without heat integration that applied to the separation of acetone and methanol. Hosgor43 revealed that PSD is significantly more economical than the ED process when the mixture of methanol and chloroform was separated, and the dynamic control schemes of PSD were investigated. However, only the basic control

2. DYNAMIC CONTROL OF TWO PROCESSES 2.1. ED Process. 2.1.1. Selection of Temperature Control Stage. The steady state of the ED process for separating the mixture containing 50 mol % ethanol and 50 mol % THF using ethylene glycol (EG) as the entrainer was designed in our previous paper.4 There are some differences of the steady-state results in this work from the previous work for reasons of different simulator versions (Aspen Plus V7.2 in this work and Aspen Plus V2006.5 in the previous work) and consideration of pressure drops between adjacent stages in two columns. Figure 1 gives the temperature profiles and the liquid composition profiles for the ED column and the entrainer recovery column in the ED process, respectively. For the selection of the temperature control stage in the ED column, stage 56 displays a fairly steep slope in the temperature profile, as shown in Figure 1a; it is B


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Table 1. Tuning Parameters of the Temperature Controllers in the Basic Control Structure with Fixed Reflux Ratios of the ED Process parameter

TC1

TC2

TCHE

controlled variable manipulated variable transmitter range controller output range ultimate gain ultimate period gain, KC integral time, τI

T1,56 QR1 273.2−446.1 K 0−24.57 GJ/h 24.62 4.80 min 7.70 10.56 min

T2,6 QR2 273.2−598.8 K 0−10.91 GJ/h 0.69 4.20 min 0.21 9.24 min

Trecycle QHE 273.2−366.9 K −13.21−0 GJ/h 0.33 2.40 min 0.10 5.28 min

Figure 2. Improved control structure with S/F control of the ED process.

2.1.2. Basic Control Structure with Fixed Reflux Ratios. The basic control structure with fixed reflux ratios of the ED process for separating the mixture of ethanol and THF is shown in Figure S1. Eleven controllers added in the control structure are summarized as follows. (1) Feed is flow controlled (reverse acting). (2) Reflux drum levels in both columns are held by manipulating the flow of distillates (direct acting). (3) Base level in the ED column is held by manipulating the flow of the bottoms (direct acting). (4) Base level in the entrainer recovery column is held by manipulating the makeup EG flow rate (reverse acting). (5) The total entrainer is flow-controlled by manipulating the flow of the bottoms of the entrainer recovery column (reverse acting). (6) The pressure in the two columns is controlled by manipulating the heat removal rate in the condenser of the two columns (reverse acting).

noticed that the composition of THF in stage 56 shows a sharp decrease from Figure 1c. Therefore, stage 56 in the ED column was selected as the temperature control stage. Similarly, stage 6 was selected as the temperature control stage for the entrainer recovery column, as shown in Figure 1b,d. Note that the composition of THF increases to 99.90 mol % and the composition of EG reduces close to zero quickly in the rectifying section that has only three stages in the ED column. The composition of three components have sharp changes on the feed stage from the extractive section to the stripping section. The heuristic method that provides 5 min of liquid holdup when half full was used to calculate the size of the reflux drums for two columns and the size of the sump for the ED column.44 Meanwhile, the volume of the sump for the entrainer recovery column is specified to provide holdup for 15 min when it is filled with ∼50% to handle the inherent swing. The ratios of height or length to diameter of reflux drums and sumps are both set to 2. In addition, the pressure drops of all the pumps and valves are sized to handle changes in flow rates. C


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Figure 3. Dynamic responses for the improved control structure with S/F controlled: (a) feed flow rate disturbances and (b) feed composition disturbances.

(7) Entrainer feed temperature is held by manipulating the cooler HE heat duty (reverse acting). (8) The temperature for stage 56 in the ED column is controlled by manipulating the reboiler heat input into the ED column (reverse acting). (9) The temperature for stage 6 in the entrainer recovery column is controlled by manipulating the reboiler

heat input into the entrainer recovery column (reverse acting). Notice that the bottom level of the entrainer recovery column is controlled according to the inventory control loop suggested by Grassi and Luyben.45,46 The controller “FCtot” is on “cascade” because it receives the signal from the multilier “S/F” (solvent flow rate/feed flow rate), and its value equals the solvent D


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Figure 4. (a) LPC temperature profile and (b) HPC temperature profile.

flow rate/feed flow rate. The reflux ratios of 4.97 for the ED column and 0.48 for the entrainer recovery column are held constant in the control loop. The basic controllers, such as the flow controller, level controllers, and pressure controllers, are implemented using the conventional PI controller. The proportional and integral (PI) settings of the flow control and the pressure control loops are set at KC = 0.5, τI = 0.3 min and KC = 20, τI = 12 min, respectively. All level controllers are used with KC = 2. Three deadtime elements with the deadtime of 1 min are inserted into the control loop for the three temperature controllers. After the relay−feedback tests for the temperatures are run, the Tyreus−Luyben tuning rule is used to calculate the gains and integral times. The tuning parameters are listed in Table 1. The dynamic responses of the basic control structure are tested by introducing the disturbances in feed flow rate and feed composition. In the following disturbance figures, the solid lines stand for the increases of feed flow rate or the composition of ethanol in the feed and the dashed lines stand for the decreases of feed flow rate or the composition of ethanol in feed without further specific clarifications. Figure S2 gives the performances of the basic control structure for the disturbances. As can be seen from Figure S2a, the purities of two products can return to initial values when the feed flow rate is reduced by 20%; however, the structure cannot control the disturbances of 20% increase of feed flow because the purities of ethanol and THF have some deviations from its specifications when the new steady state arrived. Figure S2b indicates the performances of the control structure that handle the disturbances of the composition of ethanol in the feed flow. It is found that the purities of two products deviate from the initial value of 99.90 mol % (99.54 mol % of ethanol and 99.67 mol % of THF) when the composition of ethanol in the feed flow is changed from 50 mol % to 40 mol %. Therefore, the basic control structure with fixed reflux ratios cannot handle the disturbances efficiently. 2.1.3. Improved Control Structure with S/F Controlled. On the basis of the steady-state results and the performances of the basic control structure, the improved control structure with S/F controlled in the control loop is presented in Figure 2. The reflux flow rate of the entrainer recovery column is in proportion to the feed instead of holding constant reflux ratio. The reflux ratio of the ED column remains constant. The solvent flow rates that

proportioned with the feed flow rate is not sufficient for the increasing feed flow rate and the changing composition in the feed stream to maintain the purity of THF in the distillate stream in the basic control structure. Thus, a composition controller that detects the purity of THF in the distillate of the ED column is added in the loop, which controlled the value of S/F. A deadtime element with the deadtime of 3 min is inserted into the control loop for the composition controller “CC”. Relay−feedback test and the Tyreus−Luyben tuning rule are performed for the CC controller; the achieved parameters of the CC controller are KC = 365.78 and τI = 112.20. Figure 3 gives the dynamic performances of the improved control structure that handle the disturbances in feed flow rate and feed composition. Notice that the positive change of the feed flow rate is 15% and the decrease of the composition of ethanol in feed flow changed from 50 mol % to 42 mol %. When the feed is changed, the ratio of S/F immediately varies the solvent fed to the extractive column. The changed flow rate of the solvent that is proportional to the feed flow rate in the basic control structure is not sufficient for the feed disturbances, especially for the feed composition disturbances. A composition controller, CC, is used to vary the flow rate of the solvent by detecting the purity of THF. As can be seen from Figures S2 and 3, the improved control structure maintains the purities of two products well, especially for the purity of ethanol when the negative change of ethanol in feed composition occurs. As shown in Figure 3, the purity of THF produced from the ED column has transient deviations at about t = 1.5 h and then goes back to the desired value smoothly with the improved control structure equipped. The purity of ethanol taken from the top of the entrainer recovery column has some quite large transient deviations at about t = 1.5 h then returns to the initial value soon, at about t = 2.2 h, with the negative changes of feed flow rate and ethanol composition in feed stream. In the improved control structure, the reflux ratio of the ED column and the R/F ratio of the entrainer recovery column are kept constant in terms of the values of the reflux ratios of two columns (4.97 for the ED column and 0.48 for the solvent recovery column). The control structure with S/F controlled is selected for this binary azeotrope system according to the mutual effect of the entrainer and the mixture. If the relative volatility of both the light E


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input the needed heat into the LPC. The heat duty of the auxiliary reboiler equals 985.48 kW, and the reboiler/condenser is sized by using an overall heat-transfer coefficient of 0.0020448 GJ/(h m2 K), resulting in an area of 24.49 m2. The basic control structure of the PHIPSD process is given in Figure S4. The “flowsheet equations” function is used to achieve the partial heat integration in Aspen Plus Dynamics (Figure S5a). The first equation is used to calculate the heat-transfer rate in the condenser of the HPC. The second equation is used to specify the total heat input in the LPC, which takes the summation of the negative of the value of the first equation and the heat input from the auxiliary reboiler to be the output signal from the “TC1” controller. The reflux ratios of two columns are kept constant in the basic control structure. It is noted that the pressure controller “PC2” is on manual and its value floats with operating conditions. Before the control structure is performed, the relay−feedback tests and the Tyreus−Luyben tuning rule are used to give the parameters for the temperature controllers listed in Table 2. Figure S6 presents the responses of the basic control structure with partial heat integration. It is found that the purities of both products have some deviations from the desired values after arriving at a new steady state when the disturbances of ±20% in feed flow rate and the composition of ethanol in feed stream are introduced at 1 h. For the positive 20% increase in feed flow rate and the decrease of the ethanol from 50 mol % to 40 mol %, the final purity of THF are 97.81 mol % and 98.44 mol %, respectively. 2.2.2. Pressure-Compensated Temperature Control Structure. Figure S7 shows the pressure-compensated temperature control structure based on the improvement of the basic control structure. In the PHIPSD process equipped with an auxiliary

Table 2. Tuning Parameters of the Temperature Controllers in the Basic Control Structure of PHIPSD parameter

TC1

TC2

controlled variable manipulated variable transmitter range controller output range ultimate gain ultimate period gain, KC integral time, τI

T1,36 Qaux 301.3−401.3 K 0−7.10 GJ/h 16.61 4.20 min 5.19 9.24 min

T2,34 QR2 381.8−481.8 K 0−8.98 GJ/h 7.52 5.40 min 2.35 11.88 min

and heavy components increased by the entrainer is obvious, the fixed S/F control structure would be efficient. 2.2. PHIPSD Process. Our previous paper reveals that PHIPSD has total annual cost (TAC) smaller than ED and conventional no heat-integrated PSD when the mixture of 50 mol % ethanol and 50 mol % THF is separated.4 It is necessary to give an efficient control structure for PHIPSD. The temperature profiles of the LPC and the HPC in the PSD process are shown in Figure 4. According to the slope criterion suggested by Luyben,44 the temperature control stages for the LPC and HPC are selected as stage 36 and 34, which is consistent with the results calculated by singular value decomposition (SVD) (Figure S3). In the PSD process without heat integration, the condenser duty in the HPC is 923.24 kW, which can be used to produce vapor in the base of the LPC (1908.72 kW for the reboiler). These results have some differences from the previous work due to different versions of simulators. 2.2.1. Basic Control Structure. To implement the partially heat-integrated process, an auxiliary reboiler should be added to

Figure 5. Improved pressure-compensated temperature control structure of the PHIPSD process. F


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Figure 6. Dynamic responses for the improved pressure-compensated temperature control of PHIPSD: (a) feed flow rate disturbances and (b) feed composition disturbances.

reboiler, the pressure of the HPC fluctuates along with the change of the feed stream. If a fixed temperature in the HPC is controlled, the purity of the product in the HPC will change along with the pressure changes. Thus, the pressurecompensated temperature control structure is often used in the control of PHIPSD with auxiliary reboiler.

A ratio block between the feed and heat input to the LPC can improve the response of the system. A summation block, one of whose input signals is from the multiplier block “Qtot/F” (i.e., the calculated total heat input to the LPC/feed flow rate), is added to the control loop. The other input signal into the summation block comes from the TC1 temperature controller G


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Figure 7. Comparisons of dynamic responses between ED and PHIPSD: (a) feed flow rate disturbances and (b) feed composition disturbances.

after a sign change in the multiplier block “negative”. The output of the summation is the required heat duty of the auxiliary reboiler. If an auxiliary condenser is used in PHIPSD, only the multiplier block can realize the proportion of the heat duty and the feed flow rate without the block for sign change. To set up a pressure-compensated temperature control structure, the bubble point temperatures of the liquids on stage 34 in the HPC at 9.53−10.65 atm (the fluctuation range of the column pressure in the HPC when the feed flow rate has ±20% changes) were investigated. The bubble point temperature of the mixture as a function of pressure is given, and a corresponding line is fitted with a slope value of 5.18. The temperature signal is adjusted by the following equation: TPC = T2,34 − (P − 10.13)5.18. Both the temperature on stage 34 and the column pressure of the HPC are

used to calculate TPC, whose value is sent to the temperature controller TC2. The “flowsheet equations” should be adjusted to implement the heat integration, and the pressure-compensated temperature control structure is shown in Figure S5b. The dynamic responses of the pressure-compensated temperature control structure to the disturbances are shown as Figure S8. According to the comparison between Figures S6 and S8, it is found that the purity of THF in the HPC can be effectively controlled by the pressure-compensated temperature control structure. However, the purity of ethanol has not obviously improved when a new steady state arrived under the disturbances of feed flow and feed composition. 2.2.3. Improved Pressure-Compensated Temperature Control Structure. On the basis of the pressure-compensated H


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Figure 8. Control structure of PHIPSD for separating the mixture containing 10 mol % ethanol and 90 mol % THF.

disturbances in the feed flow rate and feed composition, and the purities exceed the initial specification (99.90 mol %). Luyben44 presented a control structure that estimates the composition by measuring both temperature and pressure of the control tray for the system in which pressures change significantly because of the mode of operation. This control structure makes the control of the PHIPSD process concise in some case studies. Figure S9 shows the control structure (CS1) suggested by Luyben. Flowsheet equations for estimating the composition on the control trays are shown in Figure S10. Dynamic performance comparisons between the CS1 and the control structure given in this work (CS2) can be found in Figure S11. The CS1 cannot make the purity of ethanol return to its specification when the disturbance of increase of ethanol in feed stream occurs. The CS2 can maintain the purities of two products under two kinds of disturbances. It is also found that the fluctuation ranges of product purities controlled by the CS2 are narrower than those of product purities controlled by the CS1. Thus, the improved pressure-compensated temperature control structure (i.e., the CS2) is selected in the PHIPSD process. 2.3. Comparisons between Two Processes. Both distillation processes can separate the binary azeotrope of ethanol and THF effectively. However, PHIPSD demonstrates the advantages, when compared with ED, of a simple process, less TAC, no third solvent introduced, and good controllability. Further comparison of dynamic control performances of the two flowsheets can be seen in Figure 7. From Figure 7a, the purities of two products produced from two alternative distillation processes quickly return to the acceptable value when the disturbances of the feed flow rate are introduced. The purity of ethanol

temperature control structure, we introduce a composition controller that detects the purity of ethanol in the bottom stream of the LPC to form the composition−temperature cascade control structure (Figure 5). If the pressure-compensated temperature control structure for PHIPSD for separating the regular binary mixture can maintain the purities of two products, the composition controller is not needed. Note that TC1 is on “cascade” and the composition controller CC is the primary controller. The parameters of the CC controller of KC = 17.69 and τI = 40.92 min are obtained after the relay−feedback test and the Tyreus−Luyben tuning rule are employed. This control structure would be effective because the composition−temperature cascade control and the pressure-compensated temperature control are used to control the purities of two products. Figure 6 shows the control performances of the improved pressure-compensated temperature control structure. As can be seen from Figures S8 and 6, the improved pressure-compensated temperature control structure can deal with the disturbances of positive changes of the feed flow rate and the content of ethanol in feed stream to maintain the purity of ethanol produced from the LPC compared with the pressure-compensated temperature control structure. Apart from the pressure-compensated temperature control in the HPC, the composition−temperature cascade control can supply sufficient heat for the LPC by the TC1 controller according to the purity of ethanol. As shown in Figure 6, the purity of ethanol has large transient deviations at 1−1.5 h (for the feed flow rate disturbances) and 1.5 h (for the feed composition disturbances) and then returns to its initial value gently at about 10 h. In addition, the purity of THF in the bottom of the HPC reaches a new steady state quickly for the I


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Table 3. Tuning Parameters of the Controllers for PHIPSD with the Mixture of 10 mol % Ethanol and 90 mol % Ethanol 10 mol % ethanol

TC1

TC2

CC

controlled variable manipulated variable transmitter range controller output range ultimate gain ultimate period gain, KC integral time, τI 90 mol % ethanol controlled variable manipulated variable transmitter range controller output range ultimate gain ultimate period gain, KC integral time, τI

T1,25 QR1/F 378.1−478.1 K 0−0.115 GJ/kmol 17.50 3.60 min 5.47 7.92 min

T2,15 QR2/F 299.4−399.4 K 0−0.015 GJ/kmol 13.06 4.20 min 4.08 9.24 min

XB1(THF) R/F 0−1.998 0−1.799 92.35 89.40 min 28.86 196.68 min

T1,10 Qaux 258.5−358.5 K 0−5.49 GJ/h 13.38 4.80 min 4.18 10.56 min

Tpc,24 QR2/F 381.2−481.2 K 0−0.019 GJ/kmol 30.40 6.00 min 9.50 13.20 min

XB1(THF) R/F 0−1.998 0−2.216 192.70 44.40 min 60.30 97.68 min

Figure 9. Control structure of PHIPSD for separating the mixture containing 90 mol % ethanol and 10 mol % THF.

control performances of PHIPSD compared with the ED process are concluded in the former section. It will be interesting to study the control structures of PHIPSD for different feed composition. 3.1. Mixture Containing 10 mol % Ethanol and 90 mol % THF. The optimal PHIPSD scheme for the separation of 10 mol % ethanol and 90 mol %THF is HPC-LPC. Figure 8 gives the final dynamic control structure for the PHIPSD process separating this mixture. An auxiliary condenser is used to remove redundant heat in the condenser of the HPC. In the control loop, the reflux flow rate is in proportion to the feed flow rate for the HPC to improve the dynamic response of the control structure in

produced from PHIPSD achieves a much more satisfying performance with ED when the negative 20% change in feed flow rate occurs. For the disturbances of feed flow composition, control structure for PHIPSD shows superiority because its dynamic responses have narrower fluctuation ranges and shorter recovery times than those of the ED process.

3. PHIPSD WITH DIFFERENT FEED COMPOSITION In our previous paper,6 the optimal PSD processes for separating the mixtures containing 10 mol % ethanol−90 mol %THF and 90 mol % ethanol−10 mol % THF are given. More satisfying J


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Figure 10. Comparisons of dynamic responses between the PHIPSD processes with different feed composition: (a) feed flow rate disturbances and (b) feed composition disturbances.

the face of large reflux ratio of the HPC. This ratio of reflux flow rate/feed flow rate (R/F) is controlled by a composition controller, CC, that detects the purity of THF in the bottom stream of the HPC. The heat duties of two columns are also in proportion to the feed flow rate. The ratios are controlled by corresponding temperature controllers to hold the temperatures in columns. After running the relay−feedback tests and using the Tyreus−Luyben tuning rule, the parameters of the controllers are listed in Table 3. 3.2. Mixture Containing 90 mol % Ethanol and 10 mol % THF. For the separation of the mixture of 90 mol % ethanol and 10 mol % THF, the best sequence of PHIPSD is arranged as the LPC−HPC scheme. An auxiliary reboiler is added to replenish insufficient heat for the reboiler of the LPC. The final control structure of the PSD scheme is provided as Figure 9. The temperature control stages for the LPC (with 45 stages) and the

HPC (with 35 stages) are located at stage 10 and 24, respectively. A pressure-compensated temperature control structure is used in the HPC to handle the effect of changing pressure on the sensitive stage in the HPC. In addition, a composition controller, CC, that detects the purity of ethanol in the bottom stream of the LPC is used to control the R/F ratio in the LPC to produce qualified ethanol. Table 3 lists the tuning parameters using the same calculated methods as the former for the temperature and composition controllers. 3.3. Comparisons between PHIPSD for Different Feed Composition. Figure 10 gives the responses of the control structures of the PHIPSD processes for separating the mixtures with different composition. It is found that the fluctuation ranges of two products for separating the 50 mol % ethanol and 50 mol % THF mixture are smaller than those of the other two feed composition. It is because a composition−temperature K


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composition controller is essential to the effective control performances.

cascade control for the LPC and a pressure-compensated temperature control for the HPC are assembled in the PHIPSD process for separating the mixture of equal composition. A multiplier block, Qtot/F, for the LPC can also improve the sensitivity of the control structure to the feed disturbances. The purity of ethanol produced form the 10 mol % ethanol and 90 mol % THF mixture is poorer than that produced from the other two composition mixtures. For the mixtures containing 50 mol % and 90 mol % ethanol, the corresponding control structure brings the purity of ethanol back to its initial value quickly with the help of a composition controller. It is indicated that a composition controller in the LPC in the PHIPSD process for separating the 50 mol % ethanol and 50 mol % THF and 90 mol % ethanol and 10 mol % THF mixtures is essential to effective controls. As can be seen from Figure 10b, the purity of THF needs more time to return to its specification for the separation of 90 mol % ethanol and 10 mol % THF mixture when the composition of ethanol is changed from 90 mol % to 95 mol %. The reason is many stages in the LPC delay the reaction from the reflux stream to the bottom product purity. The optimal sequences for separating the mixtures containing 10 mol %, 50 mol %, and 90 mol % ethanol are HPC−LPC, LPC−HPC, and LPC−HPC, respectively. An auxiliary condenser for PHIPSD for the mixture containing 10 mol % ethanol and auxiliary reboilers for PHIPSD for the mixture containing 50 mol % and 90 mol % ethanol are used to compose partial heat integration. Feed steam composition has a great influence on the control structure of PHIPSD because of the different combinations of auxiliary heat exchanger.

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ASSOCIATED CONTENT

S Supporting Information *

The Supporting Information is available free of charge on the ACS Publications website at DOI: 10.1021/acs.iecr.5b01642. Dynamic simulation details and additional figures (PDF)

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AUTHOR INFORMATION

Corresponding Author

*E-mail: [email protected]. Notes

The authors declare no competing financial interest.

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ACKNOWLEDGMENTS Comments and suggestions from two anonymous reviewers and Professor William L. Luyben are gratefully acknowledged. Financial support 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.

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4. CONCLUSION Control structures of ED and PHIPSD for separating the mixture of 50 mol % ethanol and 50 mol % THF and the comparisons between the two processes are given in this work. Dynamic controls of PHIPSD for separating other two streams with different composition are also investigated in view of the superiorities of PHIPSD compared with ED. Although the introduction of a composition controller in the control structure of ED caused long lags, the control structure can maintain the purities of two products within appropriate times. The proposed control structure shows its effectiveness in the ED process in which the light component is captured by the entrainer in the ED column and the temperature difference between the top and bottom of the entrainer recovery column is not quite large. The control structures of PHIPSD are effective in handling the feed disturbances for the binary azeotropic mixture of ethanol and THF whose T−x−y diagram shows special features under different pressures, although long lags are produced by a composition controller and the control structure cannot deal with the large disturbances of feed composition for the mixtures of 10 mol % ethanol and 10 mol % THF. The control structures are suitable for the binary azeotropic mixture whose temperature differences between dew points and bubble points show no obvious changes with compositions near the azeotropic composition. According to the comparisons between the ED and PHIPSD processes, it is found that PHIPSD has the advantages of energy savings and good controllability. Furthermore, comparisons between PHIPSD for three different composition mixtures were also given based on the investigations of their dynamic controls. It is concluded that the feed composition mainly impacts the dynamic control structures of PHIPSD through the different combination of auxiliary heat exchanger and that the

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NOTATION CC = composition controller ED = extractive distillation EG = ethylene glycol HPC = high pressure column KC = gain of controller LPC = low-pressure column PHIPSD = partially heat-integrated pressure-swing distillation PSD = pressure-swing distillation R/F = reflux flow rate/feed flow rate S/F = solvent flow rate/feed flow rate TC1, TC2 = temperature controller of column 1, 2 TAC = total annual cost THF = tetrahydrofuran τI = integral time of controller REFERENCES

(1) Genduso, G.; Amelio, A.; Luis, P.; Van der Bruggen, B.; Vreysen, S. Separation of Methanol-Tetrahydrofuran Mixtures by Heteroazeotropic Distillation and Pervaporation. AIChE J. 2014, 60, 2584−2595. (2) Luyben, W. L. Methanol/Trimethoxysilane Azeotrope Separation Using Pressure-Swing Distillation. Ind. Eng. Chem. Res. 2014, 53, 5590− 5597. (3) Mahdi, T.; Ahmad, A.; Nasef, M. M.; Ripin, A. State-of-the-Art Technologies for Separation of Azeotropic Mixtures. Sep. Purif. Rev. 2015, 44, 308−330. (4) Wang, Y.; Cui, P.; Ma, Y.; Zhang, Z. Extractive Distillation and Pressure-Swing Distillation for THF/Ethanol Separation. J. Chem. Technol. Biotechnol. 2015, 90, 1463. (5) Skogestad, S. Control Structure Design for Complete Chemical Plants. Comput. Chem. Eng. 2004, 28, 219−234. (6) Wang, Q.; Yu, B.; Xu, C. Design and Control of Distillation System for Methylal/Methanol Separation. Part 1: Extractive Distillation Using DMF as an Entrainer. Ind. Eng. Chem. Res. 2012, 51, 1281−1292. (7) Wang, S.-J.; Yu, C.-C.; Huang, H.-P. Plant-Wide Design and Control of DMC Synthesis Process via Reactive Distillation and Thermally Coupled Extractive Distillation. Comput. Chem. Eng. 2010, 34, 361−373. (8) Luyben, W. Plantwide Dynamic Simulators in Chemical Processing and Control; CRC Press: Boca Raton, FL, 2002.

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Industrial & Engineering Chemistry Research (9) Gorak, A.; Schoenmakers, H. Distillation: Operation and Applications; Academic Press: London, 2014. (10) Luyben, W. L. Control of the Heterogeneous Azeotropic NButanol/Water Distillation System. Energy Fuels 2008, 22, 4249−4258. (11) Luyben, W. L. Control of a Column/Pervaporation Process for Separating the Ethanol/Water Azeotrope. Ind. Eng. Chem. Res. 2009, 48, 3484−3495. (12) Li, G.; Bai, P. New Operation Strategy for Separation of Ethanol− Water by Extractive Distillation. Ind. Eng. Chem. Res. 2012, 51, 2723− 2729. (13) Wei, H.-M.; Wang, F.; Zhang, J.-L.; Liao, B.; Zhao, N.; Xiao, F.-k.; Wei, W.; Sun, Y.-H. Design and Control of Dimethyl Carbonate− Methanol Separation via Pressure-Swing Distillation. Ind. Eng. Chem. Res. 2013, 52, 11463−11478. (14) Cao, W.; Mujtaba, I. M. Simulation of Vacuum Membrane Distillation Process for Desalination with Aspen Plus. Ind. Eng. Chem. Res. 2015, 54, 672−680. (15) Xu, H.; Ye, Q.; Zhang, H.; Qin, J.; Li, N. Design and Control of Reactive Distillation−Recovery Distillation Flowsheet with a Decanter for Synthesis of N-propyl Propionate. Chem. Eng. Process. 2014, 85, 38− 47. (16) Hameed, S.; Das, B.; Pant, V. A Self-Tuning Fuzzy PI Controller for TCSC to Improve Power System Stability. Electr. Power Syst. Res. 2008, 78, 1726−1735. (17) Pavkovic, D.; Polak, S.; Zorc, D. PID Controller Auto-Tuning Based on Process Step Response and Damping Optimum Criterion. ISA Trans. 2014, 53, 85−96. (18) Barreto, A. A.; Rodriguez-Donis, I.; Gerbaud, V.; Joulia, X. Optimization of Heterogeneous Batch Extractive Distillation. Ind. Eng. Chem. Res. 2011, 50, 5204−5217. (19) Suppino, R. S.; Cobo, A. J. G. Influence of Solvent Nature on Extractive Distillation of the Benzene Hydrogenation Products. Ind. Eng. Chem. Res. 2014, 53, 16397−16405. (20) You, X.; Rodriguez-Donis, I.; Gerbaud, V. Improved Design and Efficiency of the Extractive Distillation Process for Acetone−Methanol with Water. Ind. Eng. Chem. Res. 2015, 54, 491−501. (21) Henley, E. J.; Seader, J. D.; Roper, D. K. Separation Process Principles; Wiley: New York, 2011. (22) Zhang, Z.-g.; Huang, D.-h.; Lv, M.; Jia, P.; Sun, D.-z.; Li, W.-x. Entrainer Selection for Separating Tetrahydrofuran/Water Azeotropic Mixture by Extractive Distillation. Sep. Purif. Technol. 2014, 122, 73−77. (23) Arifin, S.; Chien, I.-L. Design and Control of an Isopropyl Alcohol Dehydration Process via Extractive Distillation using Dimethyl Sulfoxide as an Entrainer. Ind. Eng. Chem. Res. 2008, 47, 790−803. (24) Gil, I. D.; Gómez, J. M.; Rodríguez, G. Control of an Extractive Distillation Process to Dehydrate Ethanol using Glycerol as Entrainer. Comput. Chem. Eng. 2012, 39, 129−142. (25) Qin, J.; Ye, Q.; Xiong, X.; Li, N. Control of Benzene− Cyclohexane Separation System via Extractive Distillation Using Sulfolane as Entrainer. Ind. Eng. Chem. Res. 2013, 52, 10754−10766. (26) Torres-Ortega, C. E.; Segovia-Hernández, J. G.; Gómez-Castro, F. I.; Hernández, S.; Bonilla-Petriciolet, A.; Rong, B.-G.; Errico, M. Design, Optimization and Controllability of an Alternative Process Based on Extractive Distillation for an Ethane−Carbon Dioxide Mixture. Chem. Eng. Process. 2013, 74, 55−68. (27) 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. (28) Wang, S.-J.; Wong, D. S.; Yu, S.-W. Effect of Entrainer Loss on Plant-Wide Design and Control of an Isopropanol Dehydration Process. Ind. Eng. Chem. Res. 2008, 47, 6672−6684. (29) Luyben, W. L.; Chien, I.-L. Design and Control of Distillation Systems for Separating Azeotropes; Wiley: New York, 2011. (30) Luyben, W. L. New Control Structure for Feed-Effluent Heat Exchanger/Reactor Systems. Ind. Eng. Chem. Res. 2012, 51, 8566−8574. (31) Luyben, W. L. Pressure-Swing Distillation for Minimum- and Maximum-Boiling Homogeneous Azeotropes. Ind. Eng. Chem. Res. 2012, 51, 10881−10886.

(32) Wang, Y.; Cui, P.; Zhang, Z. Heat-Integrated Pressure-SwingDistillation Process for Separation of Tetrahydrofuran/Methanol with Different Feed Compositions. Ind. Eng. Chem. Res. 2014, 53, 7186− 7194. (33) Zhu, Z.; Wang, L.; Ma, Y.; Wang, W.; Wang, Y. Separating Azeotropic Mixture of Toluene and Ethanol via Heat-Integration Pressure Swing Distillation. Comput. Chem. Eng. 2015, 76, 137−149. (34) Luis, P.; Amelio, A.; Vreysen, S.; Calabro, V.; Van der Bruggen, B. Simulation and Environmental Evaluation of Process Design: Distillation vs. Hybrid Distillation−Pervaporation for Methanol/ Tetrahydrofuran Separation. Appl. Energy 2014, 113, 565−575. (35) Mulia-Soto, J. F.; Flores-Tlacuahuac, A. Modeling, Simulation and Control of an Internally Heat Integrated Pressure-Swing Distillation Process for Bioethanol Separation. Comput. Chem. Eng. 2011, 35, 1532− 1546. (36) Kiran, B.; Jana, A. K. A Hybrid Heat Integration Scheme for Bioethanol Separation through Pressure-Swing Distillation Route. Sep. Purif. Technol. 2015, 142, 307−315. (37) Luyben, W. L. Design and Control of a Fully Heat-Integrated Pressure-Swing Azeotropic Distillation System. Ind. Eng. Chem. Res. 2008, 47, 2681−2695. (38) 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. (39) Wang, Y.; Zhang, Z.; Zhang, H.; Zhang, Q. Control of Heat Integrated Pressure-Swing-Distillation Process for Separating Azeotropic Mixture of Tetrahydrofuran and Methanol. Ind. Eng. Chem. Res. 2015, 54, 1646−1655. (40) 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. (41) Luyben, W. L. Comparison of Extractive Distillation and Pressure-Swing Distillation for Acetone-Methanol Separation. Ind. Eng. Chem. Res. 2008, 47, 2696−2707. (42) Modla, G.; Lang, P. Removal and Recovery of Organic Solvents from Aqueous Waste Mixtures by Extractive and Pressure Swing Distillation. Ind. Eng. Chem. Res. 2012, 51, 11473−11481. (43) Hosgor, E.; Kucuk, T.; Oksal, I. N.; Kaymak, D. B. Design and Control of Distillation Processes for Methanol−Chloroform Separation. Comput. Chem. Eng. 2014, 67, 166−177. (44) Luyben, W. L. Distillation Design and Control Using Aspen Simulation; Wiley: New York, 2013. (45) Grassi, V. G., II. Process Design and Control of Extractive Distillation; Springer: New York, 1993. (46) Luyben, W. L. Plantwide Control of an Isopropyl Alcohol Dehydration Process. AIChE J. 2006, 52, 2290−2296.

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