Journal Pre-proofs Efficient extractive distillation design for separating binary azeotrope via thermodynamic and dynamic analyses Zhaoyou Zhu, Xiaopeng Yu, Yixin Ma, Jingwei Yang, Yinglong Wang, Peizhe Cui, Xin Li PII: DOI: Reference:
S1383-5866(19)34746-X https://doi.org/10.1016/j.seppur.2019.116425 SEPPUR 116425
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Separation and Purification Technology
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Please cite this article as: Z. Zhu, X. Yu, Y. Ma, J. Yang, Y. Wang, P. Cui, X. Li, Efficient extractive distillation design for separating binary azeotrope via thermodynamic and dynamic analyses, Separation and Purification Technology (2019), doi: https://doi.org/10.1016/j.seppur.2019.116425
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Efficient extractive distillation design for separating binary azeotrope via thermodynamic and dynamic analyses
Zhaoyou Zhua,c, Xiaopeng Yua, Yixin Mab, Jingwei Yanga, Yinglong Wanga,c, Peizhe Cuia, Xin Lia,*
College of Chemical Engineering, Qingdao University of Science and Technology,
Qingdao 266042, China b College of Chemical and Environmental Engineering, Shandong University of Science
and Technology, Qingdao, 266590, China c
Shandong Collaborative Innovation Center of Eco-Chemical Engineering, Qingdao
University of Science and Technology, Qingdao 266042, China
Corresponding Author *E-mail: [email protected]
Abstract Efficient and stable separation of ethyl acetate/cyclohexane mixture in an economic way is significant for the pesticide industries. The economy of the extractive distillation (ED) process performs and the pressure-swing distillation (PSD) process was compared based on total annual cost (TAC). Thermodynamic analysis was used to further modify the ED process. Two-stage preheating was applied to improve the energy use efficiency and further reduce the costs. The results show that the TAC of the ED process is lower than that of PSD process by 58.2% and the cost of the twostage preheating extractive distillation process is 9.3% lower than traditional ED process. To separate the mixture steadily, an improved multi-proportion control structure was developed. The proportion elements change their output signals based on the initial ratios when the input signals change. The improved control structure can separate the components to 99.9 mol%. Both the economy and controllability of the modified process perform well.
Keywords: Extractive distillation; Pressure-swing distillation; Dynamic control; Thermodynamic analysis; Cost saving.
1. Introduction Ethyl acetate/cyclohexane mixture is widely used as organic solvent in the pesticide industries and chemical engineering industries. It is important to recycle the components in terms of economy and environment. Ethyl acetate and cyclohexane form minimum azeotrope at atmospheric pressure. Special distillation processes have been developed to separate such azeotropes in the past decades. The special distillation processes include pressure-swing distillation (PSD) [1-3] and extractive distillation (ED) [4-6]. PSD process uses the changes of azeotropic compositions under different operating pressures to realize the separation [7, 8]. Compared with ED, no third components are required. Munoz et al.  separated the isobutyl alcohol/isobutyl acetate by PSD and ED. The results show that total annual cost (TAC) of PSD is lower than ED. Similar results can be obtained from the separation of di-n-propyl ether/n-propyl alcohol , isopropyl alcohol/diisopropyl ether  and tetrahydrofuran/ethanol . Modla and Lang  compared the two processes based on the batch distillation. Acetone and methanol were separated and the processes were evaluated by energy consumption and recoveries. The performance of pressure-swing batch distillation (PSBD) with heat integration (HI) is better than batch extractive distillation (BED). Same results can be obtained when a varied-diameter column is applied in the process . ED process is usually operated under atmospheric pressure and it can be used to separate pressure-insensitive mixtures. Luyben investigated the separation of acetone/methanol  and acetone/chloroform . The economy was taken into 3
consideration to compare the ED process and PSD process. The results show that ED process is more attractive. Thermodynamic analysis is usually used to modify the process and then improve the energy use efficiency . Heat integration technology is often used in the PSD process because of the large temperature difference between the bottom of low pressure column (LPC) and the top of high pressure column (HPC). As for the ED process, the HI technology can also be applied if the temperature difference is sufficient. Ghuge et al.  investigated the separation of tetrahydrofuran/water and compared the two processes with and without HI. The pressure of the solvent recovery distillation column was set at negative pressure so the HI can be implemented. The economic performance of ED process is better than PSD process. The previous works show that both the PSD and ED can separate the mixtures well. Pan et al. designed an extractive distillation unit using low eutectic solvent to dehydrate ethanol. The developed MPC control scheme has good control and energy efficiency . Tripodi et al. compared the efficiency and the cost of pressure-swing distillation and extractive distillation for the recovery of acetonitrile by ammonia oxidation of ethanol. The results showed that extractive distillation had higher efficiency and better economy than pressure swing distillation . Gao et al. used pressure
acetate/methanol/water mixture. The results show that the MVR separation of methyl acetate/methanol/water ternary mixtures can achieve low energy consumption and low total annual cost . In practical production, the operations of the industrial equipment are complex and 4
unstable so the controllability of the processes is another important issue . Different control structures have been explored to keep the separated products at high-purity level when disturbances occurred [23-25]. Temperature controller is used to control the stage temperature to maintain the high-purity components indirectly [26-28]. However, the controllability is not efficient when the disturbances cause a large temperature change in the column [29, 30]. Composition controller is usually cascaded with other controllers to control the composition of product streams directly by manipulating the feed flowrate, the reflux ratio, or the stage temperature [31-33]. For simple processes, the control structures perform well. While for the complex processes such as ED process which a third component is added, the control structures should be further explored. In this paper, we focus on providing an idea for the optimal separation method of ethyl acetate/cyclohexane by comparing the economy of PSD process and ED process based on the TAC. Through the thermodynamic analysis, the ED process was improved and a two-stage preheated ED (TSPED) was developed. An improved multi-proportion control structure was explored to improve the controllability of the TSPED process, which provided insight into stable production. 2. Design basis 2.1 Separation feasibility and distillation sequence of PSD process The processes were calculated with non-random two liquid (NRTL) property method. The simulation results are in good agreement with the experimental results . The binary interaction parameters are provided in Table 1. 5
Table 1. Binary interaction parameters. Ethyl acetate(i) +
Ethyl acetate(i) +
Figure S1 shows the T-x-y curves for the binary azeotrope at 1 atm and 10 atm. The azeotropic composition varies from 53 mol% ethyl acetate (1atm) to 96 mol% ethyl acetate (10 atm). The feed stream is 100 kmol/h with the compositions of 50 mol% ethyl acetate and 50 mol% cyclohexane. The Txy curves show that the azeotrope cannot be separated with the separated sequence of LPC-HPC so the feed stream should be sent to HPC at first. In HPC, high-purity cyclohexane (B1) is obtained at the bottom and the minimum azeotrope (D1) is sent to the LPC. The high-purity ethyl acetate (B2) is collected at the bottom of LPC and the minimum azeotrope (REC) would be sent back to the HPC (Figure 1a). Furthermore, the process with heat integration is designed as shown in Figure 1b.
Figure 1 Flowsheet of pressure-swing distillation process (a) without heat integration; (b) with heat integration 2.2 Entrainer selection and design of ED process In the ED process, the selection of entrainer is significant for the separation efficiency and energy saving. The relative volatility is an important criterion for selecting the entrainer. The larger relative volatility means that the difficulty of separation is lower. Four different solvents, which are dimethyl sulfoxide (DMSO), cyclohexanol, N,N-dimethylacetamide (DMAC), N-Methyl pyrrolidone (NMP) and N,N-dimethylformamide (DMF), were selected as the candidate entrainers. Figure 2 shows the effect of different solvents on vapor-liquid equilibrium of ethyl 7
acetate/cyclohexane. The DMSO performs better than the other solvents so it is selected as the entrainer in this study .
Figure 2 Relative volatilities of the ethyl acetate/cyclohexane The DMSO stream and the raw materials are fed to the first column (Figure 3). The cyclohexane is withdrawn at the top of the first column while the entrainer and ethyl acetate are sent to the second column. The high-purity ethyl acetate is obtained at the top while the DMSO is sent back to the first column. There is a make-up stream to add fresh entrainer and mix with the recycled stream.
Make-up 323.15 K 0.051 kmol/h 1.0 DMSO
0.4 atm 325 K 566.2 kW
0.7 atm 342 K 680.6 kW
C1 RR=1.52 ID=0.999 m
Feed 303.15 K 100 kmol/h 0.5 ethyl acetate 0.5 cyclohexane
D1 50 kmol/h 0.001 ethyl acetate 0.999 cyclohexane
RR=0.2 ID=0.811 m
374.4 K 1117.7 kW
D2 50 kmol/h 0.999 ethyl acetate 0.001 cyclohexane
439 K 765.1 kW
B1 180 kmol/h 0.2776 ethyl acetate 0.0003 cyclohexane 0.7221 DMSO
B2 130 kmol/h 0.0 ethyl acetate 0.0 cyclohexane 1.0 DMSO
Figure 3 Flowsheet of extractive distillation process 3. Determination of the preferred process 3.1 Comparisons of PSD process and ED process According to our previous work , the minimum TAC is used as the objective function to optimize the processes, and the optimization is based on the sequential iterative optimization procedure (Figure S2). The basis of the economics is shown in Table S1. In the PSD process, the optimized operating pressure of LPC is 0.5 atm. As for the pressure of HPC, the TAC decreases with the pressure increasing but high pressure increases the risk so the pressure was set at 10 atm. Figure S3 shows the changes of TAC with the variables of PSD process. The optimal stage numbers of HPC and LPC are 69 and 36 by calculating TAC under different stages. The feed locations are 40 for HPC and 15 for LPC. The recycle stream was fed to stage 34 of the HPC. The flowrate and composition of recycle stream also have impact on TAC. When the 9
recycle stream is 39 kmol/h with composition of 51.1 mol% ethyl acetate and 48.9 mol% cyclohexane, the minimum TAC is 1672516 $/y. The TAC would be reduced to 1497377 $/y when the heat integration was implemented. In the ED process, the C1 column and C2 column are operated at 0.7 atm and 0.4 atm. It is noteworthy that operating at vacuumed conditions may cause air leakage into the column. Usually, the pressure of the column with leak prevention measures can reduce to lower than 0.4 atm. The low operating pressures for the columns are in the range of 0.4 atm~0.7 atm in this study. The pressures were determined according to the use of cooling water and high pressure steam to realize the minimal TAC. The optimal stage numbers of C1 and C2 are 48 and 18 (Figure S4). The feed stages of the two columns are 36 and 14. The feed location of the recycle stream is 14 and the flowrate of recycle stream is 130 kmol/h. The minimum TAC of ED process is 699639 $/y which is much lower than the PSD process by 58.2%. 3.2 ED process with Two-stage preheating To improve the energy utilization and further reduce the TAC, the thermodynamic analysis profiles of exergy loss and enthalpy deficit are investigated and shown in Figure 4. Higher exergy loss reflects the lower thermodynamic efficiency so the energy consumption increases. The rate of exergy loss on the column stage can be expressed as Eq.1. Eloss L j 1 E jL1 V j 1 E Vj 1 Fj E jF L j E jL V j E Vj S j E jS EQ , j
T EQ , j Q j 1 o T
Figure 4 Thermodynamic profiles of ED: (a)(b) exergy loss; (c)(d) enthalpy deficit. The enthalpy deficit in Figure 4c and 4d is the cumulative heating and cooling requirements for the column stages and the exergy losses are caused by the system driving force or entropy production due to the losses of momentum (pressure drop), heat (temperature driving force/mixing) and chemical potential (mass transfer driving force/mixing). Figure 4a and 4b show that the largest exergy losses are located around the feed stages. The temperature of the feed stream is lower than the stage temperature so the mixing of the different state streams has negative impact on energy utilization. The feed streams of C1 and C2 can be preheated sequentially by the bottom stream of C2 to the feed stages temperature, which are 350 K for C1 and 390 K for C2 (Figure 11
5). The temperature of the bottom stream is decreased from 439 K to 379 K. Taking the comprehensive utilization of the energy into consideration, the feed stream temperatures 350 K and 390 K were selected since they are close to the bubble point. The heat duty of the first preheater is 239 kW and that for the second preheater is 119 kW. Less enthalpy deficit and exergy loss prove the role of two-stage preheating in saving energy (Figure 6). The exergy loss reduces from 27 kW to 15 kW in the 36 stage of C1. It also reduces in C2. The minimum TAC of TSPED process is decreased to 634429 $/y though two preheaters were added in the process. The TAC is further reduced by 9.3%. Make-up 323.15 K 0.05 kmol/h 1.0 DMSO
0.7 atm 342 K 680.6 kW
0.4 atm 325 K 601.1 kW D1 50 kmol/h 0.001 ethyl acetate 0.999 cyclohexane
RR=1.52 ID=0.867 m
374.4 K 886.7 kW
Feed 303.15 K 100 kmol/h 0.5 ethyl acetate 0.5 cyclohexane
B1 180 kmol/h 0.2776 ethyl acetate 0.0003 cyclohexane 0.7221 DMSO
D2 50 kmol/h 0.999 ethyl acetate 0.001 cyclohexane
RR=0.274 ID=0.827 m
439 K 681.3 kW
B2 130 0.0 0.0 1.0
kmol/h ethyl acetate cyclohexane DMSO
Figure 5 Flowsheet of two-stage preheated extractive distillation process.
Figure 6 Thermodynamic profiles of TSPED and EG: (a)(b) exergy loss; (c)(d) enthalpy deficit. 4. Control structure design of TSPED process 4.1 Basic control structure design and performance (CS1) Control structure design of TSPED process is significant for the stable separation. The basic control structure of TSPED process to separate the binary azeotrope is shown as Figure S5. There are eleven controllers in the basic control structure . The stage 20 of C1 and the stage 15 of C2 were selected as the temperature control stages based on the temperature sensitive tray method (Figure 7) . The controller Frec cascaded with S/F can increase or decrease the flowrate of the entrainer stream by manipulating the V21. The parameters of the level controllers, the flow controllers, and the pressure controllers were set at KC=2 (τI=9999 min), KC=0.5 (τI=0.3 min) and KC=20 (τI=12 13
min). The gains and the integral times of the temperature controllers with deadtime elements are obtianed by the Tyreus-Luyben tuning rule. The parameters of the temperature controllers are shown in Table S2.
Figure 7 Slope criterion for TSPED process. Figure S6 shows the dynamic responses of the basic control structure when the ±10% flowrate disturbances and the ±10% composition disturbances are introduced. The results show that the control of ethyl acetate product purity is not satisfactory. The main impurity in the stream is different (Figure S7), so the basic control structure needs to be improved to handle the disturbances efficiently. 4.2 Control structure with heat duty of reboiler/feed flow rate (Q/F) ratio (CS2) In order to improve the anti-disturbance of the system, the Q/F control structure 14
was studied. The CS2, which has Q/F control structure, is shown in Figure 8. The principle of Q/F control structure：when the flow rate into the C2 column changes, the reboiler heat duty changes in time. In terms of liquid distribution in the distillation column, the Q/F control structure makes the system more stable and enhances the antidisturbance capability. The dynamic responses to ±10% feed composition and feed flow rate are shown in Figure S8a and Figure S8b. The transient deviation of water is less compared with basic control structure when deal with -10% feed flow rate disturbances (the cyclohexane mole fraction changes from 50mol% to 45mol%) in Figure S7b. The dynamic performance is not obviously better than the basic control structure when deal with ±10% feed composition disturbance. The Q/F control structure cannot control the composition disturbance well, because the Q/F control structure is insensitive to the composition disturbance and cannot implement corresponding adjustments in time. The improved control strategies should be investigated.
Figure 8 The control structure CS2 of the reboiler heat duty to the feed flow rate (Q/F). 4.3 Improved multi-proportion control structure (CS3) To further improve the controllability of the TSPED process, a multi-proportion control structure was explored (Figure 9). Composition controllers (CC1 and CC2) were cascaded with temperature controllers (TC1 and TC2). The deadtimes of 3 min were inserted into the control loop for the composition controllers. The output signals of the TC1 and TC2 were connected with reflux ratio elements (Table S3). For the ethyl acetate product stream (D2), the main impurity is cyclohexane so high purity ethyl acetate can be obtained if the cyclohexane is withdrawn at the top of the column C1. A QR/XCH element was set to control the heat input of the reboiler in the column C1 based on the cyclohexane composition in the feed stream F2. If the cyclohexane increases, the heat input of the reboiler would increase so more cyclohexane can be vaporized. For the cyclohexane product stream (D1), the main impurity is ethyl acetate. The ethyl acetate component comes from the feed stream F1 and the recycle stream REC. If the component can be obtained in D2 as more as possible, the purity of the cyclohexane would be purer. The flowrate of F1 and REC was added firstly (SUM element) and then used to control the heat input of the reboiler in column C2 (Q/SUM element). In the column C1, the separation of cyclohexane and ethyl acetate is depended on the flowrate of entrainer. The SUM2 element was applied to add together the ethyl acetate flowrate in stream F1 and stream REC. Then the flowrate was used to determine the amount of entrainer DMSO (S/F element). The output signal was connected with the flow 16
controller Frec. A high/low value selector was inserted in the S/F control loop. The function of the selector was set as high value selection, so the higher entrain flowrate between the set value and the calculated value would be selected as the input signal for the controller Frec. The initial entrainer flowrate with 129.995 kmol/h was used as the set point to ensure the sufficient entrainer amount when the disturbances occur. Figure 10 shows the dynamic responses of the improved control structure by introducing flowrate disturbances and composition disturbances. Both the components of the binary azeotrope can be separated to 99.9 mol% so the improved control structure can control the process well.
Figure 9 Improved multi-proportion control structure.
Figure 10 Dynamic responses of the improved control structure: (a) flowrate disturbances; (b) feed composition disturbances. 18
5. Conclusions In this work, the favorable separation performance of ED process for ethyl acetate/cyclohexane mixture was studied based on the minimum TAC. Thermodynamic analysis of two columns in ED process shows that the mixture streams with different states reduces the energy efficiency. The heat of the recycle stream is used to preheat the feed steams of the two columns to the corresponding feed stage temperatures. The two-stage preheated ED process improves the energy efficiency. The minimum TAC is further reduced by 9.3%. The traditional control structures cannot keep the purity of the product streams stably when the feed flowrate disturbances and the feed composition disturbances occur. In order to improve the anti-disturbance ability of the TSPED process, an improved multi-proportion control structure was explored. The control structure can adjust the parameters immediately by monitoring the changes of the streams based on the multiproportion models. The high/low value selector enables the amount of the entrainer so the purities of the separated components can be kept to over 99.9 mol%, which shows good controllability.
Corresponding Author E-mail: [email protected]
Notes The authors declare no competing financial interest.
Acknowledgments This work is supported by the National Natural Science Foundation of China (No. 21676152) and National Natural Science Foundation of China (No. 21776145).
Nomenclature AD = azeotropic distillation BED = batch extractive distillation DMAC = N,N-dimethylacetamide DMF = N,N-dimethylformamide DMSO = dimethyl sulfoxide ED = extractive distillation FC = flow controller HI = heat integration HPC = high pressure column τI = integral time of the controller (min) KC = gain of the controller LPC = low pressure column NMP = N-Methyl pyrrolidone NRTL = non-random two liquid PSBD = pressure-swing batch distillation PSD = pressure-swing distillation TAC= total annual cost TSPED = two-stage preheated ED
Supplementary material Table S1. Basis of economics. Table S2. Parameters of basic control structure. Table S3. Parameters of improved control structure with multi proportion. Figure S1 Txy diagram for ethyl acetate/cyclohexane. Figure S2 Sequential iterative optimization procedure Figure S3 Total annual costs of PSD under different parameters: (a) stage number of HPC; (b) stage number of LPC; (c) feed stage of HPC; (d) feed stage of LPC; (e) flowrate of recycle stream; (f) ethyl acetate mole fraction of recycle stream; (g) feed stage of recycle stream; (h) pressure of LPC. Figure S4 Total annual costs of ED under different parameters: (a) stage number of C1; (b) stage number of C2; (c) feed stage of C1; (d) feed stage of C2; (e) feed stage of recycle stream; (f) flowrate of recycle stream. Figure S5 Basic control structure. Figure S6 Dynamic responses of the basic control structure: (a) flowrate disturbances; (b) feed composition disturbances. Figure S7 Compositions of impurities in the ethyl acetate product stream. Figure S8 Dynamic responses to ±10% disturbance of the Q/F control structure: (a) feed flow rate; (b) feed composition
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Supplementary material Efficient extractive distillation design for separating binary azeotrope via thermodynamic and dynamic analyses
Zhaoyou Zhua c, Xiaopeng Yua, Yixin Mab, Jingwei Yanga, Yinglong Wanga c, Peizhe Cuia, Xin Lia *
of Chemical Engineering, Qingdao University of Science and Technology,
Qingdao 266042, China bCollege
of Chemical and Environmental Engineering, Shandong University of
Science and Technology, Qingdao, 266590, China cShandong
Collaborative Innovation Center of Eco-Chemical Engineering, Qingdao
University of Science and Technology, Qingdao 266042, China
Corresponding Author *E-mail: [email protected]
List of Table captions and Figure captions Table captions Table S1. Basis of economics. Table S2. Parameters of basic control structure. Table S3. Parameters of improved control structure with multi proportion.
Figure captions Figure S1 Txy diagram for ethyl acetate/cyclohexane. Figure S2 Sequential iterative optimization procedure Figure S3 Total annual costs of PSD under different parameters: (a) stage number of HPC; (b) stage number of LPC; (c) feed stage of HPC; (d) feed stage of LPC; (e) flowrate of recycle stream; (f) ethyl acetate mole fraction of recycle stream; (g) feed stage of recycle stream; (h) pressure of LPC. Figure S4 Total annual costs of ED under different parameters: (a) stage number of C1; (b) stage number of C2; (c) feed stage of C1; (d) feed stage of C2; (e) feed stage of recycle stream; (f) flowrate of recycle stream. Figure S5 Basic control structure. Figure S6 Dynamic responses of the basic control structure: (a) flowrate disturbances; (b) feed composition disturbances. Figure S7 Compositions of impurities in the ethyl acetate product stream. Figure S8 Dynamic responses to ±10% disturbance of the Q/F control structure: (a) feed flow rate; (b) feed composition
Table S1. Basis of economics. Parameters
Heat transfer Coefficient = 0.852 kW/K∙m2 Capital cost = 2427.08 × (A, m2)0.65×(2.29+FC) Q
Heat transfer area: A = (K × ∆t) Reboilers
Heat transfer Coefficient = 0.568 kW/K∙m2 Capital cost = 2427.08 × (A, m2)0.65×(2.29+FC) Q
Heat transfer area: A = (K × ∆t)
FC=(Fd+Fp) ×Fm, Fm=3.75, Fd=1.35, Fd=0.8, Fp=0 Q is the duty of heat exchanger (kW); ∆t is temperature difference (K); k is heat transfer coefficients Capital Cost = 4794.33 × (D, m)1.066 × (H, m)0.802×(2.18+FC) Length: H = 1.2 × 0.61 × (NT - 2) Diameter: D is calculated by Aspen tray sizing FC=Fm×Fp, Fm=3.67, Fp =1.10 when pressure=10 atm, Fp =1.00 when pressure is pressure≤3.4 atm. LP stream (433 K): $7.78 per GJ MP stream (457 K): $8.22 per GJ HP stream (527 K): $9.88 per GJ Cool water (310 K): $0.354 per GJ
Table S2. Parameters of basic control structure. Controller
Transmitter range (°C)
Integral time τI (min)
Controller action Controlled variable
Table S3. Parameters of improved control structure with multi proportion. Controller
X(CH, Mole fraction)
X(EA, Mole fraction)
Controller action Controlled variable Manipulated variable Transmitter range (°C) Gain KC Integral time τI (min)
Figure S1 Txy diagram for ethyl acetate/cyclohexane.
Figure S2 Sequential iterative optimization procedure
Figure S3 Total annual costs of PSD under different parameters: (a) stage number of HPC; (b) stage number of LPC; (c) feed stage of HPC; (d) feed stage of LPC; (e) flowrate of recycle stream; (f) ethyl acetate mole fraction of recycle stream; (g) feed stage of recycle stream; (h) pressure of LPC. 36
Figure S4 Total annual costs of ED under different parameters: (a) stage number of C1; (b) stage number of C2; (c) feed stage of C1; (d) feed stage of C2; (e) feed stage of recycle stream; (f) flowrate of recycle stream.
Figure S5 Basic control structure.
Figure S6 Dynamic responses of the basic control structure: (a) flowrate disturbances; (b) feed composition disturbances. 42
Figure S7 Compositions of impurities in the ethyl acetate product stream.
Figure S8 Dynamic responses to ±10% disturbance of the Q/F control structure: (a) feed flow rate; (b) feed composition
Highlights 1. ED process is modified based on thermodynamic analysis and TSPED is developed. 2. The TAC of TSPED process can be saved by 9.3% compared with ED process. 3. Multi proportion control structure is explored to perform effective control.
Author Statement No conflict of interest exits in the submission of this manuscript, and the manuscript is approved by all authors for publication. I would like to declare on behalf of my co-authors that the work described was original research that has not been published previously, and not under consideration for publication elsewhere, in whole or in part. All the authors listed have approved the manuscript that is enclosed.
Declaration of interests ☒ The authors declare that they have no known competing financial interests or personal relationships that could have appeared to influence the work reported in this paper.
☐The authors declare the following financial interests/personal relationships which may be considered as potential competing interests: