Control of Heat Integrated Pressure-Swing-Distillation Process for

Article pubs.acs.org/IECR

Control of Heat Integrated Pressure-Swing-Distillation Process for Separating Azeotropic Mixture of Tetrahydrofuran and Methanol Yinglong Wang,* Zhen Zhang, Huan Zhang, and Qing Zhang College of Chemical Engineering, Qingdao University of Science and Technology, Qingdao 266042, China S Supporting Information *

ABSTRACT: Dynamic control of a pressure-swing-distillation process for separation of azeotropic mixture of tetrahydrofuran and methanol is explored. The pressure-swing-distillation processes involved with no, partial, and full heat integration are simulated using Aspen Plus Dynamics. The influences of the selection of the sensitive temperature stage in the low-pressure column on the dynamic responses in the pressure-swing distillation with different heat integration were investigated. The results indicate that a suitable temperature control stage in the low-pressure column is crucial to achieve efficient control of the process. In addition, more time is needed to reach the quality specifications under feed disturbances for both components when heat integration is added in the distillation sequence.

1. INTRODUCTION The separation of azeotropic mixtures is an interesting and important topic for academic research and industrial application.1−3 The binary mixture of tetrahydrofuran (THF) and methanol exhibits a minimum boiling azeotrope because of molecular repulsion between the two chemical components, thus leading to difficulty in obtaining high-purity products. In our previous work, the separation of this mixture using pressure-swing distillation (PSD) was reported using steadystate simulation.4 However, dynamic control of the PSD was not involved in that study. It is of great importance to study the dynamic control process, which can help maintain the purity of the products in the face of disturbances in feed flow rate and feed composition. Dynamic control schemes for separating azeotropic mixtures have been studied in recent years.5−18 Skogestad5 proposed a systemic and model-based method for the control structure design. Luyben13,15 studied the separation of azeotropic mixtures in detail from steady-stage design to dynamic control. Since the controller design that based on heuristic methods draws the attention of most researchers, influences of control variables on dynamic performances of control schemes are studied widely.19−28 For example, Wang22 investigated the effect of entrainer loss on plantwide control of an isopropanol dehydration process. Jones27,28 focused on the selection of the primary and secondary control variables in the plantwide control system design. More importantly, all the studies on control structure and the selection of control variables can promote the applications of dynamic control in the chemical processes.17,29,30 For instance, Luyben29 showed the control schemes for the multiunit heterogeneous azeotropic distillation for separating the mixture of ethanol and water with benzene as the entrainer. Wei17 gave an efficient control structure of the PSD for separating the mixture of dimethyl carbonate and methanol with the large feed flow and feed composition disturbances. In addition, the dynamic control of the divided wall column attracts great attention due to its investment and © 2015 American Chemical Society

energy savings compared with traditional distillation schemes.31−35 Design, modeling, and optimization of the PSD processes have been studied recently.36,37 Some researchers also concentrated on the dynamic controllability of the PSD process.13,17,38,39 Luyben13 demonstrated the basic control structure of the PSD process for separating the acetone/ methanol azeotropic mixture in detail. The possibilities of the heat integration for the PSD process, thermally coupled distillation sequences, and dividing wall distillation make them more competitive than traditional distillation.40−44 Among them, in addition to energy and capital saving, the PSD process has a prominent merit that no third component is introduced in the distillation system for separating azeotropic mixtures. Thus, the dynamic control of the PSD with heat integration for separating the azeotrope deserves further study.37−39,45 Li37 gave a composition/temperature cascade control structure for the new PSD with partial heat integration, which can handle the feed flow and feed composition disturbances perfectly. Yu39 presented various control structures for the PSD with full heat integration when the mixture of methylal and methanol was used. Also, the comparison between the PSD and extractive distillation is reported according to the steady-stage and/or dynamic control performance.36,46 The temperature controller detects the signal from the control stage of a column and then transmits it to the manipulated variable. The selection of the temperature control stage based on the slope criterion was advised by Luyben.15 The slope criterion is feasible for distillation column with a large change in temperature from stage to stage. As for the distillation column that the temperature profile shows no obvious temperature change from stage to stage, Luyben suggests choosing the stage according to the sensitivity Received: Revised: Accepted: Published: 1646

November 6, 2014 January 9, 2015 January 19, 2015 January 19, 2015 DOI: 10.1021/ie505024q Ind. Eng. Chem. Res. 2015, 54, 1646−1655

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Industrial & Engineering Chemistry Research

Figure 1. Temperature profiles and temperature slope value plots of two columns.

Figure 2. Basic temperature control structure for the PSD process without heat integration.

criterion.15 It is known that temperatures on all stages can be quickly affected by manipulating heat input. Luyben46 investigated the effect of the two alternative control stages in extractive column on the tight control of an extractive distillation system for separating CO2 and ethane in the enhanced oil recovery processes, suggesting that composition control is required.

In this work, the PSD processes containing a low-pressure column (LPC) and a high-pressure column (HPC) are used for separation of azeotropic mixture of tetrahydrofuran and methanol. Two alternative control stages are selected in the LPC, respectively, while one stage is chosen as the control stage in the HPC. We develop effective control structures of the PSD without and with heat integration for investigating how the two 1647

DOI: 10.1021/ie505024q Ind. Eng. Chem. Res. 2015, 54, 1646−1655

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Industrial & Engineering Chemistry Research Table 1. Controller Tuning Parameters for Different PSD Processes no integration controlled variable manipulated variable transmitter range controller output range ultimate gain ultimate period gain Kc integral time τI controlled variable manipulated variable transmitter range controller output range ultimate gain ultimate period gain Kc integral time τI

TC1 T1,33 QR1 376.2−476.2 K 0−10.951 GJ/h 2.27 5.40 min 0.71 11.88 min TC2 T2,18 QR2 286.2−386.2 K 0−4.489 GJ/h 32.01 3.00 min 10.01 6.60 min

partial integration

full integration

T1,33 QR1 376.2−476.2 K 0−10.951 GJ/h 2.76 4.80 min 0.86 10.56 min

T1,33 QR1 376.2−476.2 K 0−10.951 GJ/h 3.72 4.80 min 1.16 10.56 min

T2,18 RR2 286.2−386.2 K 0−5.980 58.36 21.00 min 18.24 46.20 min

T2,31 RR2 290.2−390.2 K 0−6.817 21.84 10.20 min 6.82 22.44 min

(1) Feed is flow-controlled (reverse acting). (2) The pressure in each column is controlled by manipulating the heat removal of condenser (reverse acting). (3) Reflux drum levels in both columns are held by manipulating distillate flow (direct acting). (4) Sump levels in both columns are held by manipulating bottom flow (direct acting). (5) The temperature on stage 33 in the HPC is controlled by manipulating the reboiler heat input in the HPC (reverse acting). (6) The temperature on stage 18 in the LPC is controlled by manipulating the reboiler heat input in the LPC (reverse acting). In the control scheme, reflux ratios are 1.79 for the HPC and 2.99 for the LPC, which are held constantly in each column. Relay-feedback tests are run on the two temperature controllers to determine ultimate gains and periods, and the Tyreus− Luyben tuning rule is used in two temperature controllers. Table 1 lists the temperature transmitter ranges, controller output ranges, and tuning parameters. To further model the other dynamics in the control system, a deadtime of 1 min is set for each temperature controller. All level controllers are proportional with gain (Kc) of 2. Figure 3 shows the dynamic responses for the basic control structure of the PSD without heat integration. It is observed that when new stable regulatory controls are achieved for positive 20% (solid lines) and negative 20% (dashed lines) flow rate disturbances at 1 h in feed stream, the purities of THF in bottom stream of HPC (B1) return to 99.87 and 99.92 mol % (Figure 3a), respectively. However, the purity of methanol in the bottom stream of the LPC (B2) has a large deviation from the desired value of 99.90 mol % when the control structure encounters positive 20% feed flow rate disturbance. The responses as feed composition disturbances from 25 to 30 mol % methanol (solid lines) and from 25 to 20 mol % methanol (dashed lines) are also simulated and similarly the obtained results are for the purity of THF and methanol as 20% step changes in feed composition disturbance at 1 h (Figure 3b). Composition/Temperature Cascade Control Structure. Temperature control has the advantage of being fast, but it

control stages in the LPC affect the dynamic performance of the PSD process for separation of THF and methanol.

2. PROCESS STUDY AND SELECTION OF TEMPERATURE CONTROL STAGE In our previous work, the optimized flowsheet processes for separation of azeotropic mixture containing 75 mol % THF and 25 mol % methanol were available via the steady-state PSD with and without heat integration.4 The pressure difference between two adjacent stages is assumed to 0.0068 atm in the LPC and HPC. Furthermore, an auxiliary condenser in the PSD process is also integrated into the process for the purpose of partial heat integration. It is of great importance to select the temperature control stage in both columns. Figure 1 shows temperature profiles and corresponding slope values as functions of column stages for the HPC and LPC. It can be seen that the HPC could be effectively controlled because its temperature profile exhibits a fairly sharp temperature break; however, the shape of the temperature profile has no breaks for the LPC and the slope values of the temperature profile stay almost the same in each stage. According to the “slope criterion” suggested by Luyben,15 stage 33 in the HPC is selected as a control stage to keep its temperature constant, while stage 18 or 31 in the LPC can be selected as a control stage in the temperature control loop structure. 3. CONTROL STRUCTURES FOR THE PSD PROCESSES The volumes of reflux drums and sumps are specified to provide holdup for 5 min when the vessels are filled with ∼50% using the heuristic method.15 The ratios of height to diameter of reflux drums and sumps are both set as 2. The tray-sizing option in Aspen Plus is selected to calculate the diameters of two columns. Pumps and valves are sized to give adequate pressure drops to handle changes in flow rates. 3.1. Control of the PSD without Heat Integration. Basic Temperature Control Structure. Figure 2 shows the basic temperature control structure for separating binary mixture of THF and methanol without heat integration. It is noticed that nine controllers in this scheme are included as follows. 1648


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First, stage 18 in the LPC is selected as a temperature control stage as the same as that in the basic control structure (Figure 4). Table 2 lists the tuning parameters of the composition Table 2. Tuning Parameters of the Composition Controller with Different Deadtimes for the PSD Process without Heat Integration deadtime ultimate gain ultimate period gain Kc integral time τI

2 min

3 min

4 min

889.80 6.00 min 278.06 13.2 min

345.98 11.40 min 108.12 25.08 min

70.18 22.20 min 21.93 48.84 min

controller with different deadtimes for the PSD process. It can be seen that the integral time is positively correlated with the deadtime of the composition controller. The shorter the deadtime of the controller, the less time the control scheme needs to reach a new stead state from the disturbed state. In reality, however, the detectors have specific measurement deadtimes, and the measurement of the composition controller has larger deadtime than the temperature controller. The deadtime of the composition controller is set at 3 min, which is consistent with most published papers.10,46,47 Notice that the controller “TC2” is on “cascade” with its set point from the temperature of stage 18 in the LPC, and the composition controller “CC” is tuned with TC2 on cascade. The output signal from the controller “TC1” is the heat duty of reboiler in the bottom of the LPC, and the output signal from the controller TC2 is the ratio of heat input of the reboiler in the distillation of the LPC to the feed flow rate “QR2/F”. Figure 5 gives the dynamic response results for the composition/ temperature cascade control structure for the PSD without heat integration. The solid lines stand for the increases in feed flow rate, and the methanol content in the feed is increased from 25

Figure 3. Dynamic responses for the basic control structure for the PSD without heat integration: feed flow rate and feed composition disturbances.

may not keep the product purity constant. Composition control is slow, but it will drive product purity to the desired value. To integrate the advantageous features of both, a cascade combination of composition and temperature control structure is developed according to the performance of the basic control structure. As illustrated in section 2, both stage 18 and 31 in the LPC are possible temperature control stages, where the temperature can be controlled by manipulating the heat duty of reboiler in the LPC. In this section, we will discuss the influence of the temperature control stage in the LPC on cascade control dynamic performance.

Figure 4. Composition/temperature cascade control structure for the PSD process without heat integration. 1649


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to 30 mol % methanol. The dashed lines represent the decrease in feed flow rate and the methanol content in the feed decreased from 25 to 20 mol %. It is noticed that decreasing the methanol in the feed produces less distillates (D1) and more bottoms (B1) in the LPC, and temperature decreases and reboiler duty increases in the LPC. Both of the two columns are well-controlled. Then, Figure S1 shows the cascade control structure for the PSD process with stage 31 in the LPC as the temperature control stage. The temperature controller TC1 and the composition controller CC are retuned. Figure S2 gives the responses of the control structure encountering disturbances in feed flow rate and feed composition as the above. Both product purities are held close to their specifications. It is noticed that the temperature of stage 18 in the LPC is more closely controlled than that on stage 31. The reason for this is that the transmitted temperature signal from stage 18 is quicker than that from 31 because the stream is fed at stage 20 in the LPC. Thus, stage 18 in the LPC is selected as the control stage in the cascade control scheme of the PSD without heat integration. 3.2. Control of the PSD with Partial Heat Integration. Basic Temperature Control Structure. As for the PSD process with partial heat integration, the heat input of the LPC is completely provided by the HPC with combined reboiler/ condenser. The excessive heat of the HPC is removed by an auxiliary condenser added in the process. The area of 16.67 m2 for the combined reboiler/condenser is calculated according to the steady-state simulation. The heat duties of HPC condenser and the LPC reboiler are 979.37 and 586.09 kW, respectively. The heat removal rate of the auxiliary condenser is equal to 393.28 kW, and the negative sign of this value is specified as the initial output signal of the “PC1” controller. The overall heat transfer coefficiency is specified as 0.002 044 8 GJ/(h m2 K). Figure S3 gives the basic temperature control structure and the controller faceplates. In this control structure, the flow controller has the normal setting of 0.50 for Kc and 0.30 min

Figure 5. Dynamic responses for the composition/temperature cascade control structure for the PSD without heat integration: feed flow rate and feed composition disturbances.

Figure 6. Improved composition/temperature cascade control structure for the PSD process with partial heat integration. 1650


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Industrial & Engineering Chemistry Research for integral time (τI) and the level controllers have a Kc value of 2. After running relay-feedback tests on the two temperature controllers, the ultimate gains and the periods on two temperature controllers are determined and listed in Table 1. The “flowsheet equations” function is used to achieve the partial heat integration (Figure S4a). The first equation is used to calculate the reboiler duty of LPC, while the heat removal rate of HPC is calculated by the second equation. Figure S5 gives the dynamic responses of the basic control structure for the PSD with partial heat integration. The control structure is tested by making fresh feed flow rate disturbances (Figure S5a) and feed composition disturbances at 1 h (Figure S5b). It can be seen that the THF in B1 stream can keep almost the desired purity when the control loop arrives at a new steady state after encountering ±20% changes in feed flow rate and feed methanol. The purity of methanol in B2 stream has a large deviation from the initial value when the feed flow rate varies from 100 to 120 kmol/h, and the methanol content in the feed is changed from 25 to 30 mol %. Improved Composition/Temperature Cascade Control Structure. The basic temperature control structure could not achieve effective disturbance rejection for feed flow rate and feed composition disturbances. The reason for this is that the almost constant heat input of the LPC could not provide enough energy to produce sufficient vapor up to the LPC when the disturbances occur. Thus, the low purity of methanol in B1 was discharged. To solve this problem, the multiplier QR2/F and the composition controller are added to provide the heat for better dynamic performance based on the basic temperature control structure (Figure S5). The influence of temperature control stage in the LPC on the control structure of the PSD with partial heat integration will be investigated in this section. First, Figure 6 shows the improved cascade control structure with the detective temperature signal of the LPC from stage 18. Notice that only one equation is used in the “flowsheet equation” (Figure S4b) because the ratio of QR2/F is controlled by TC2. Figure 7 gives the performance of the cascade control structure when positive (solid lines) and negative (dashed lines) 20% disturbances in the feed flow rate and feed methanol are encountered at 1 h. It can be seen that the purity of THF in B1 stream is brought back to 99.92 and 99.87 mol % when the methanol content in the feed stream varies from 25 to 30 mol % and from 25 to 20 mol %, respectively, while the methanol in B2 stream can be kept constant after 12 h using the cascade control structure for the feed flow and feed composition disturbances. Second, Figure S6 shows the cascade control structure for the PSD process with partial heat integration when stage 31 is selected as the temperature control stage in the LPC. Figure S7 gives results for feed flow and feed composition disturbances. It is noticed that the purity of methanol in B2 stream can return to the desired value of 99.90 mol % after transient deviations when the feed flow and feed composition disturbances occurred at 1 h. However, the THF in B1 has a fairly large deviation after 4 h when the feed methanol varies from 25 to 20 mol %. It can be concluded that the composition/temperature cascade control structure with stage 18 or 31 in the LPC as the temperature control stage can handle the ±20% disturbances in the feed flow rate and feed composition very well. The difference between the two cascade control structures is the way of variation of the methanol in the B2 stream. So stages 18 and 31 of the LPC can be selected as control stages in the control loop of the PSD with partial heat integration.

Figure 7. Dynamic responses for the improved composition/ temperature cascade control structure for the PSD with partial heat integration: feed flow rate and feed composition disturbances.

3.3. Control of the PSD with Full Heat Integration. Basic Control Structure. The large temperature difference between the condenser of the HPC and the reboiler of the LPC makes the full heat integration possible for the PSD process. As for the fully heat-integrated process, the only heat input that is required is from the heat duty QR1 of the HPC reboiler because the heat duty of the LPC reboiler is provided by the combined condenser/reboiler. The result of overall heat transfer coefficiency (0.002 044 8 GJ/(h m2 K)) multiplied by the heat transfer area and the temperature difference between the condenser of the HPC and the reboiler of the LPC equal to the heat duty of combined condenser/reboiler. The flowsheet equations are employed to achieve the full heat integration in Aspen Plus Dynamics. In Figure S4c, the first equation was used to calculate the heat duty input of the LPC, and the second equation was used to make the equivalence of the heat input rate of the LPC reboiler and heat removal rate of the HPC condenser. The basic control structure is given in Figure S8. After running the relay-feedback tests on two temperature controllers, the calculated gains and periods by Tyreus−Luyben tuning are 1.06 and 10.56 min, respectively. Figure S9 gives the effects of feed flow rate and feed composition disturbances on the control structure. Both of the temperatures on stage 33 in the HPC and on stage 18 in the LPC are controlled very well by manipulating the reboiler duty of HPC and reflex ratio (RR2) of LPC. The product purities in two bottom streams are 1651


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Figure 8. Improved pressure−compensated temperature control structure for the PSD with full heat integration when stage 31 in the LPC is selected as the control stage.

parameters are listed in Table 1. The control structure and control panel of all controllers are shown in Figure S13. Feed flow rate and feed composition disturbances are added to test the performances of the control scheme (Figure S14). This pressure-compensated temperature control structure with stage 31 in the LPC as the temperature control stage shows the same variation tendencies for the purities of two products and temperature when stage 18 is selected as the control stage in the LPC. Improved Pressure-Compensated Temperature Control. The improved pressure-compensated temperature control structure is established based on the above pressurecompensated temperature control structure (Figure S11). A composition controller that detects the methanol in the bottom stream of the LPC and a “deadtime” element with deadtime of 3 min are added in the control loop. Figures S15 and 8 give the improved pressure−composition control structure with the detected temperatures of stages 18 and 31 in the LPC, respectively. The performances of the control structure are tested by introducing the feed flow rate and feed composition disturbances at 1 h (Figures S16 and 9). It can be seen that the temperatures on two control stages of two columns deviate from the initial values. The purity of THF in the bottom stream of the HPC returns to close proximity of its setting point. Methanol in the bottom stream of the LPC bounces back to its steady-state value. It can be found from Figure 9 that the improved pressure-compensated temperature control structure can deal with the feed flow rate and composition disturbances very well when stage 31 in the LPC is selected as the control stage. The reboiler duties of HPC changed directly after disturbances occurred, while the temperature on stage 31 of the LPC controlled by manipulating the reflex ratio (RR2) arrives at a new steady-state value gradually. It is also observed that the THF in B1 stream in the control loop with stage 31 in the LPC as the control stage returns to the constant values in ∼3 h, which is faster than the control loop when stage 18 in the LPC is selected as the control stage. The purity of methanol in the

not maintained at their specifications (99.90 mol %) when the positive disturbances of feed flow rate and feed composition of methanol are encountered. An increase of the feed flow rate results in a temperature drop in the stripping section of the LPC, leading to more THF escaping from the bottom. Pressure-Compensated Temperature Control. Pressurecompensated temperature control has been described by Buckley.48 Before setting up pressure-compensated temperature control for the PSD with full heat integration, we investigate the bubble point temperatures of the liquids on stage 33 in the HPC at the pressure of 9−11 atm. Figure S10 shows the bubble point temperature of the mixture as a function of pressure, and a corresponding line is fitted with a slope value of 5.016. The pressure-compensated temperature is calculated by the following equation, TPC = T1,33 − (P − 10.1325), where the temperature is measured from stage 33 and the pressure is obtained from the top of the HPC. The TPC from deadtime element “dead1” is sent to the temperature controller TC1. The flowsheet equation is applied to carry out the pressurecompensated temperature calculation (Figure S4d).39 In the following text, we will investigate the effect of the control stage in the LPC on the pressure-compensated temperature control structure. First, stage 18 in the LPC is selected as the control stage (Figure S11) in the pressure-compensated temperature structure, and corresponding dynamic responses of this control structure on the ±20% changes in feed flow rate and feed composition are shown in Figure S12. It can be seen that the pressure-compensated temperature control structure could not handle the feed flow rate and feed composition disturbances because the methanol in B2 stream changes from 99.90 to 93.50 mol % when the feed flow rate varies from 100 to 120 mol % at 1 h. Second, the temperature of stage 31 in LPC is measured and controlled. The controller is tuned by running relay-feedback tests and using Tyreus−Luyben tuning rules. The temperature transmitter ranges, controller output ranges, and tuning 1652


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fairly sharp temperature break; however, the shape of the temperature profile has no breaks for the LPC and the slope values of the temperature profile stay almost the same in each stage. Therefore, the selection of the temperature control stage in the LPC is crucial to efficiently achieve dynamic control of the process. The influences of alternative temperature control stages 18 and 31 in the LPC on the control structures are investigated when stage 33 is used as the control stage in the HPC. As for the PSD process without heat integration, the dynamic control loop with stage 18 used as the control stage in the LPC can handle the disturbances of feed flow rate and feed composition more effectively. To the PSD with partial heat integration, the selection of control stage in LPC has little influence on the final efficient control structure from the controllability viewpoint. For the PSD with full heat integration, it shows a better control performance with stage 31 of the LPC as the control stage the feed flow and feed composition disturbances. On the basis of the steady-state simulation, the PSD with heat integrated sequence can save more energy than the conventional PSD process. In order to handle the disturbances effectively, the heat integrated distillation process should be added with the complicated control structure due to its degrees of freedom. This work will facilitate the selection of temperature control stages in PSD processes without and with heat integration for separation of other azeotropic systems, especially for the systems whose temperature profiles have two large temperature changes from stage to stage.

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ASSOCIATED CONTENT

S Supporting Information *

Dynamic simulation details and additional figures (Figures S1− S16). This material is available free of charge via the Internet at http://pubs.acs.org.

Figure 9. Dynamic responses for the improved pressure−compensated temperature control structure for the PSD with full heat integration when stage 31 in the LPC is selected as the control stage: feed flow rate and feed composition disturbances.

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B2 stream changes gently (Figure 9). Therefore, stage 31 as the temperature control stage in the LPC performs better than stage 18 in the PSD with full heat integration. 3.4. Discussion of Alternative Temperature Control Stages in LPC. Stages 18 and 31 in the LPC can both act as temperature control stages according to the slope criterion, and it is dynamically feasible for them pairing with heat input. In the LPC, stage 18 is located in the rectifying section, while stage 31 lies in the stripping section. On one hand, the temperature signal from stage 18 to the temperature controller shows faster response than that from stage 31 when the disturbances in flow rate and composition occurred, because the feed stream is fed at stage 20 in the LPC. On the other hand, the temperature of stage 31 would vary quicker than that of stage 18 by manipulating the reboiler duty. To combine the advantages of these two alternative temperature control stages in the LPC, the selection of the temperature control stages should be considered in the PSD with different heat integration.

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 National Natural Science Foundation of China (Project 21306093) is gratefully acknowledged.

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4. CONCLUSIONS In this work, dynamic controls of PSD processes without and with heat integration for separation of an azeotropic mixture of 75 mol % THF and 25 mol % methanol are simulated using Aspen Plus Dynamics. The results show that the HPC can be effectively controlled because its temperature profile exhibits a 1653

NOTATION HPC = high pressure distillation column Kc = gain of controller LPC = low pressure distillation column PSD = pressure swing distillation PC1, PC2 = pressure controller of the HPC, LPC QR = reboiler duty QR/F = reboiler duty/molar flow rate of feed RR = reflex ratio TC1, TC2 = temperature controller of the LPC, HPC THF = tetrahydrofuran τI = integral time of controller DOI: 10.1021/ie505024q Ind. Eng. Chem. Res. 2015, 54, 1646−1655

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Industrial & Engineering Chemistry Research process for bioethanol separation. Comput. Chem. Eng. 2011, 35, 1532−1546. (46) Luyben, W. L. Control of an extractive distillation system for the separation of CO2 and ethane in enhanced oil recovery processes. Ind. Eng. Chem. Res. 2013, 52, 10780−10787. (47) 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. (48) Buckley, P. S.; Luyben, W. L.; Shunta, J. P. Design of Distillation Column Control Systems; Instrument Society of America: Research Triangle Park, NC, 1985.

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