Dynamic control analysis of an eco-efficient side-stream extractive distillation configuration

Dynamic control analysis of an eco-efficient side-stream extractive distillation configuration

Journal Pre-proofs Dynamic Control Analysis of an Eco-Efficient Side-stream Extractive Distillation Configuration Qingjun Zhang, Pengyuan Shi, Wei Hou...

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Journal Pre-proofs Dynamic Control Analysis of an Eco-Efficient Side-stream Extractive Distillation Configuration Qingjun Zhang, Pengyuan Shi, Wei Hou, Shunjin Yang, Aiwu Zeng, Youguang Ma, Xigang Yuan PII: DOI: Reference:

S1383-5866(19)30762-2 https://doi.org/10.1016/j.seppur.2020.116525 SEPPUR 116525

To appear in:

Separation and Purification Technology

Received Date: Revised Date: Accepted Date:

25 February 2019 20 November 2019 4 January 2020

Please cite this article as: Q. Zhang, P. Shi, W. Hou, S. Yang, A. Zeng, Y. Ma, X. Yuan, Dynamic Control Analysis of an Eco-Efficient Side-stream Extractive Distillation Configuration, Separation and Purification Technology (2020), doi: https://doi.org/10.1016/j.seppur.2020.116525

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Dynamic Control Analysis of an Eco-Efficient Side-stream Extractive Distillation Configuration Qingjun Zhang 1, Pengyuan Shi 1, Wei Hou 1, Shunjin Yang 1, Aiwu Zeng *1, 2, Youguang Ma 1, 2, Xigang Yuan 1, 2 1 State Key Laboratory of Chemical Engineering, School of Chemical Engineering and Technology, Tianjin University, 300350, PR China 2 Chemical Engineering Research Center, Collaborative Innovative Center of Chemical Science and Engineering (Tianjin), PR China

Corresponding Author: Aiwu Zeng * Tel.: +86-022-27404732. Fax: +86-022-27404496. E-mail: [email protected] (Aiwu Zeng) Abstract: A novel eco-efficient side-stream extractive distillation alternative was proposed to separate the acetone and methanol azeotrope using water as the separating agent, which evolved from thermally coupling extractive distillation configuration with a side rectifier and the classical two-way vapor and liquid communications was converted to a one-way liquid-only transfer with adding one distillation sections and a corresponding reboiler. This arrangement was much more efficient and attractive in terms of the economics, energy-saving and environmental properties in comparison with the conventional and thermally coupling configurations. Dynamic control behavior analysis of this highly integrated non-ideal arrangement with and without heat integration is explored in this paper. The stable and robust regulatory control is achieved with the dual-product composition held rigorously closed to their initial steady-state design specifications by the construction of a newly effective and robust control structure when handling the large throughput and feed composition disturbances where the specially key control loops are controlling the temperatures of Stages 76 and 67 by the manipulation of the side-stream flowrate and reboiler heat duty for extractive column C1 with a feedforward action of reboiler heat duty-to-feed flowrate. Keywords: Side-stream extractive distillation; Heat integration; Process control; Acetone/methanol 1. Introduction Extractive distillation is a widely used alternative method for separating the azeotrope and close-boiling mixtures in chemical and petrochemical industries, while its intrinsic obstacle is the high energy consumption rates caused by the lower thermodynamic efficiency (Zhang et al. 2019a). As is well-known, the exergy losses or energy inefficiency for conventional extractive distillation process are attributed to the remixing effect of intermediate component for ternary mixture separation. The current urgent is how to reduce the energy consumption rates and improve the thermal efficiency. Some enhanced strategies, process integration and intensification, had been adopted to alleviate the loss works and improve the performance of the distillation systems such as heat integration, thermally coupling and dividing-wall column (Luyben, 2016; Tututi-Avila et al., 2014, 2017a; Xia, et al., 2012; Zhao, et al., 2018). Currently, the exploration of extractive dividing-wall column (EDWC) has received some attentions in terms of the steady-state economics and dynamic control behaviors considering its economics superiority. However, there are some issues that hinder its development and application such as how to efficiently and simply adjust the splitting vapor flowrate (vapor splitting ratio) and there is always a tradeoff between the steady-state economics and dynamic response performance for the conventional and intensified arrangements. The dynamic control

behavior of an economically efficient EDWC option for ethanol dehydrating process was explored by Tututi-Avila et al. (2014). The large throughput and feed composition disturbances were efficiently handled by the constructed overall control scheme with the intensified variable vapor split strategy. The corresponding control performance comparisons of the conventional and intensified configurations were conducted as well and indicated that there was a conflict (tradeoff) between the steady-state economics and dynamic controllability in terms of peak deviation transients and steady-state offsets (especially for feed composition interferences). The effectiveness and performance of alternative extractive distillation configurations were investigated by Li et al. (2017) for separating a 2-methoxyethanol and toluene mixture. The EDWC sequence was economically superior to others, and it was also robustly controllable when taking the variable vapor split option than the counterpart at handling the large step changes in throughput and feed composition, while the dynamic control behavior comparison for different alternatives was not explored in that study. Zhang et al. (2013) discussed the steady-state economics and dynamic control performance of the energy-saving EDWC arrangement for separating the minimum-boiling azeotrope of ethyl-acetate and isopropanol. The proposed EDWC alternative was prevailed over the conventional process in terms of energy consumption rates and total annual cost, and the dynamic control performance of control scheme with variable vapor split option was much better than that of constant vapor split option when handling the large step changes in throughput and feed composition. The large transient product deviation was efficiently weakened, and all of the disturbances were well rejected with the small product offsets taking the vapor split ratio at the bottom edge of dividing-wall as the one of key adjusting variables at facing feed disturbances (Xia et al., 2012). As is well-known, the vapor split ratio is constant in the steady-state EDWC design stage by the cross-sectional areas and the pressure drops on both sides of the partition. An expensive and complex device is required to adjust this variable. Therefore, the new economically equivalent and more operable alternative configuration should be developed. A recent interesting and insightful paper by Tututi-Avila et al. (2017b) developed an economically efficient extractive distillation arrangement and the effectiveness was tested and ranked by taking the separation of acetone /methanol and ethanol/water azeotropes and close-boiling toluene/aniline mixture as the examples. They proposed a novel extractive distillation configuration, which evolved from a partially thermally coupling extractive distillation process with the side rectifier (thermodynamically equivalent EDWC sequence). It was specially modified by substituting the interconnected vapor-liquid streams using a direct liquid side-stream as the feed of previous side rectifier. This equivalent configuration was much more efficient and flexible than the corresponding conventional, thermally coupled systems (including EDWC sequence) in terms of the economics, energy-saving and environmental properties. This new structure was proposed by Tututi-Avila et al. (2017b) through extending the design concept of Agrawal (2000, 2014) wherein this classical two-way vapor and liquid communications was converted to a one-way liquid-only transfer with adding one or more distillation sections and a corresponding reboiler and/or a condenser with each added section. Therefore, the more operable and controllable configuration was then established since it is easier to adjust the flow of a liquid stream than to regulate the corresponding the flow of a vapor stream. Its performance was further boosted by taking the integrated condenser/reboiler exchanger(s), which was achieved by adjusting the operation pressure(s) of column(s), and/or the simple feed-effluent heat exchanger. This intensified option was also economically efficient than the corresponding EDWC configuration in terms of the energy-saving and total annual cost (Wu et al., 2013; Tututi-Avila et al.,

2017b). This efficient option was extended to a multiazeotropic acetonitrile/methanol/benzene mixture using chlorobenzene as the separating agent by Wang et al. (2018). The performance and effectiveness of the side-stream extractive distillation (SEDC) arrangement was economically superior to the corresponding conventional and the intensified EDWC configuration, and the less 6.32% and 14.39% in total annual cost for the EDWC and SEDC arrangements were achieved when compared to the conventional case. Besides, it was extended to the pressure-swing distillation configuration by Wang et al. (2018). This arrangement was much more efficient and attractive in terms of the economics, energy-saving and environmental properties in comparison with the conventional and heat-integrated configurations with the aid of the mass and heat transfers through this intermediate connection stream. The exploration of the dynamic control behavior analysis for the economically efficient system is an indispensable for ensuring optimal and safe operation in the face of the inevitable interferences that the process will encounter as taught by the terminology of “simultaneous design” (Luyben, 2017). Dynamic control behavior analysis of the intensified economic-optimal SEDC arrangement was explored by Wang et al. (2018). All the products can be held closely to their steady-state design specifications using composition control loops with the acceptable deviations and short settling time after inserting the small throughput and feed composition disturbances. The further inexpensively and effectively indirect temperature control scheme was then developed by Yang et al. (2019) through the detailed analysis of the effect of design variables on the controlled variables, and the corresponding effectiveness was evaluated by introducing plus and minus 10% and 15% step changes in production rate and feed composition. However, there is an issue which is some unrealistic plumbing arrangements in the investigation of dynamic control processes. The select of control valves to control the two solvent flowrates to the two columns is unworkable. The solvent comes from the base of each column and the external makeup stream, and all branches have control valves. The solvent flowrate to the column C1 was controlled by supposing a valve (call it “V1”) on the blended solvent streamline before the total stream splitting into two streams. Another valve (call it “V2”) in the line to column C2 after the split is supposed to control the flowrate of solvent to column C2. This plumbing cannot work. The discharged pressure of control valve V2 are the pressure in column C2 (solvent feed locations). So, the available pressure drop over valve V2 is quite small, which is probably insufficient to overcome the hydraulic head in the feed line to the top of column C2. The control valves on the total blended streamline and the bottom stream of column C1 cannot exist simultaneously. The valve in the line from the blended stream should be removed. And the control valves from the base and the splitting stream to column C1 should be taken to regulate the solvent flowrate. Ma et al. (2019) investigated the dynamic control response performance of the energy-saving side-stream extractive distillation process by inserting different disturbances (10% and 20%) of throughput and feed composition with the separation of acetone and methanol using water as a separating agent as the example. The special feature control loops of their constructed efficient control structure were that product methanol composition was controlled by manipulating the side-stream withdrawn flowrate. The effective product quality control was achieved at handling the plus and minus 10% step changes in throughput and feed composition and the settling time was short to achieve the steady-state designs, while for 20% step changes, the system was not robustly controlled owing to the quite large product transient deviation, longer settle time and oscillated fluctuation of product methanol purity. There is also a major issue which is some unrealistic plumbing arrangements in the investigation of dynamic control processes. An

inappropriate control valve (named V7), regulating solvent flowrate and its discharged pressure, was located in the recycling solvent stream between the downstream of the blender and the upstream of the cooler since all branches have control valves. The system was feasible when handling plus 20% step changes in throughput, while it is of failure after 2 h of added interference owing to the side-stream flowrate dropping to zero when inserting minus 20% step change in throughput at 1 h in our repeated simulations. The system is also failed to achieve the stable regulatory control for this control structure owing to product methanol composition oscillated at near 18 h without converging to the initially steady-state design value. Moreover, the system was also not worked after 2 h of added disturbances owing to the side-stream flowrate dropping to zero when inserting 20% increase feed acetone composition disturbance at 1 h. Furthermore, the corresponding dynamic control behavior analysis for this economic-efficiently novel side-stream extractive distillation arrangement has not again received attention. Therefore, the efficient and robust control scheme is necessarily needed to develop to effectively improve the anti-interference ability and response robustness for this energy-saving alternative with a practical and workable plant-wide pluming arrangement. Besides, the investigation of the dynamic control for the intensified heat-integrated side-stream extractive distillation configuration was not found in the open literatures. This complex and highly integrated nonlinear configuration makes the robust control difficult and challenge, and the most typical effect is the reduced ability to resist interferences or even uncontrollable. Therefore, the dynamic control behavior analysis for this arrangement with and without heat integration is an indispensable and extremely urgent. That is the purpose of this paper. The contribution of this paper is to develop an efficient and robust control scheme to resist the large step changes in throughput and feed composition for this highly integrated non-ideal side-stream extractive distillation arrangement with and without heat integration. The robust control scheme with feedforward strategy is proposed, and the stable regulatory control is achieved for facing these feed interferences. 2. Conventional Side-stream Extractive Distillation Configuration 2.1 Process description Fig. 1 shows the flowsheet of an energy-efficiently novel side-stream extractive distillation arrangement. The Aspen UNIQUAC physical properties package is used. The fresh feed is fed into the extractive column C1 (operated at 1 atm) in the Stage 55 of an 81-stage column. And the relevant feed flowrate and inlet temperature are 540 kmol·h-1 and 320 K, respectively, with the equimolar composition of acetone and methanol. The solvent water (99.99 mol%) is fed on the Stage 33 at 1069.20 kmol·h-1 and 320 K. The light acetone product (99.40 mol%) is taken out in the distillate stream D1 of extractive column C1. The liquid side-stream, extracted at the Stage 73, is as the feed of column C2 (operated at 1 atm) at the Stage 29 of a 37-stage column. Product methanol (99.50 mol%) is removed out in the distillate stream D2 of column C2. The bottoms for these dual-columns are the entrainer (water) with the composition of 99.99 mol%. The small makeup solvent stream is added to compensate the corresponding entrainer losses in these two distillate streams along with the combination of bottom-end streams in these two columns to recycle back to extractive column C1 after cooling down to 320 K. According to the column sections analysis, the stripping section in column C2 has the same function as the section (started from the side-stream withdrawn location and extended up to the bottoms end in extractive column C1), which achieve the recovery of solvent water. The detailed information is demonstrated in Fig. 1, and the column configurations (total

stages, feed location and reflux ratio, etc.) are taken directly from Tututi-Avila et al. paper (2017b). For extractive column C1, it is a complex configuration with a side stream. The degree of freedom is increased by one due to the existence of liquid side withdrawn stream. The paired designed and manipulated variables for product specifications of these two columns are shown below. 1. Product acetone purity in the distillate stream (D1) is set at 99.40 mol% by manipulating the distillate flowrate of column C1; 2. The entrainer water purity in the bottoms (B1) is set at 99.99 mol% by manipulating the liquid side-stream flowrate of column C1; 3. The acetone impurity in the side stream (FS) is set at 0.131 mol% by manipulating the reflux ratio (RR1) of column C1; 4. Product methanol purity in the distillate stream (D2) is set at 99.50 mol% by manipulating the distillate flowrate of column C2; 5. The entrainer water purity in the bottoms (B2) is set at 99.99 mol% by manipulating the reflux ratio (RR2) of column C2. Herein, of the particular importance is that the acetone impurity in the side stream should be regulated in the composition specification steps. Once the acetone gets into the column C2, it will go overhead as the distillate impurity for the product methanol, which has a detrimental effect on the product methanol purity. Therefore, the acetone content should be regulated in the side stream. The corresponding heat requirements in these two reboilers are 9.164 MW and 5.336 MW for achieving the specifying separations, respectively, for columns C1 and C2. And the other parameters (reflux ratios, condenser heat removals, and column diameters, etc) are illustrated in Fig. 1.

Fig. 1. Flowsheet of a conventional energy efficient side-stream extractive distillation process for separating acetone and methanol mixture.

2.2 Dynamic control The dynamic control behavior analysis for this energy-efficient alternative is an indispensable for ensuring the optimal and safe operations as taught by the terminology of “simultaneous design” (Luyben, 2017; Zhang et al. 2020a). The preference of the regulatory control loops in distillation column is not the direct composition

measurement but the temperature inferential control strategy owing to its characteristics of convenience, inexpensiveness and simplicity (Luyben, 2013). However, the inefficiency or invalidation is sometimes observed for these inferential control schemes at facing throughput or feed composition disturbances, therefore, the robust and efficient control structure(s) should be redevised. The significant issue of finding a robust regulatory control loop for inferential temperature control strategy is to select tray location(s) whose temperature should be controlled. In this work, the sensitivity, singular value decomposition (SVD) and relative gain array (RGA) criteria are, respectively, employed to determine the sensitive temperature tray location(s) and their corresponding sensor-valve pairs between the controlled and manipulated variables (Luyben, 2013; Smith, 2012). For sensitivity criterion, the temperature changes of all stages are obtained with the small variation (± 0.1% of the design value) of the corresponding manipulated variables (such as rebolier heat duty (QR), side-stream withdrawn flowrate (FS) and reflux ratio (RR)) (Luyben, 2013; Zhang, et al. 2020b). The open-loop steady-state gains between the temperature on that tray and each manipulated variable are obtained by dividing the change in the tray temperature at holding the one of manipulating variables constant with another variation. And the tray with largest temperature change is the most sensitive that should be controlled owing to the relevant composition change from tray to tray being also large holding operation pressure constant. As obtained from Fig. 2, the temperatures of Stages 67 and 33 are sensitive to change in reflux ratio and reboiler heat input for columns C1 and C2, respectively. Besides, the side-stream flowrate has a significant effect on the Stage 76 temperature for extractive column C1 in comparison with the other two operation variables (reflux ratio and reboiler heat duty). The conventional distillation control wisdom illustrates that the tighter control can be achieved by manipulating the vapor flowrate than the liquid reflux flowrate (Luyben, 2018). Herein, for this system, the controlled temperature locations are all near to the bottoms end of columns, so it is dynamically feasible and efficient for pairing the reboiler heat inputs with these controlled sensitive temperatures for columns C1 and C2, respectively. If the manipulated variable of reflux ratio is used, the response lag time is quite large due to the inherent liquid hydraulic lags with typically 3~6 s per tray (Luyben, 2013; Zhang, et al. 2020b). For example, controlling the temperatures of Stages 67 and 33 with liquid reflux flowrate have the 201~402 s and 99~198 s lags in the temperature control loops, respectively, for columns C1 and C2, which indicates the inefficient operation. When the controlled locations are obtained, another issue is the sensor-valve pairing. Fig. 2(b) clearly gives the corresponding pairs of these controlled and manipulated variables for column C2. The temperature of Stage 33 is controlled by manipulating the corresponding reboiler heat input, while it is complex for extractive column C1 since there are the multiple extreme points including Stages 67 and 76, as observed from Fig. 2(a). These two temperatures are sensitive to the change of the reboiler heat duty and side-stream flowrate. Of the particular importance is that the successful operation of the columns C1 and C2 is to keep acetone from dropping down to the tray where the liquid side stream is withdrawn. When the acetone gets into the column C2, it will go overhead as the impurity for the product methanol, causing a detrimental effect on product methanol purity. The objective of column section from Stage 55 to Stage 73 is to keep acetone going down to the bottoms, therefore, the largest tray temperature change(s) (Stage 67) should be controlled to alleviate the effect of acetone component on product methanol purity of column C2. And the Stage 76, the large temperature change location for the column section between Stage 73 and Stage 81, should be also controlled to suppress the methanol impurity from falling into the bottom-end of extractive column C1 resulting in interfering with the solvent purity. In other

words, the dual-point temperature control scheme is used in the column C1. As observed from the above analysis, there are two possible options for the selection of operating variables to pair the corresponding controlled variables. For example, the temperature of Stage 67 can be controlled by manipulating either the side-stream withdrawn flowrate or reboiler heat duty. To further clearly determine the correspondingly manipulated variables for the temperatures of Stages 67 and 76 on column C1, the evaluation SVD and RGA methods are used. The detailed implications for the SVD criterion and the corresponding decomposed matrix are illustrated in Luyben (2013) and Zhang et al. (2018). Table 1 gives the SVD analysis results of sensor-valve pairing. The principal component of the first U vector indicates the Stage 67 temperature is the most responsive. The principal component of the first row in the VT matrix indicates that reboiler heat duty QR is the strongest manipulated variable. The pairing suggested by this analysis is the reboiler heat duty (QR) with stage 67 temperature (T67) and side-stream flowrate (FS) with stage 76 temperature (T76). Besides, to further determine the interactions of these control loops for a given control structure, the RGA method is also used. The detailed implication of the RGA concept is illustrated in the works of Smith (2015), Hansen et al. (1998) and Juan et al. (2019). Table 1 also enumerates the results of RGA analysis. Similarly, the resulting sensor-valve pairing looks good due to the lighter interactions for these two control loops. Therefore, the different pairs of sensitive temperature control tray locations for extractive columns C1 is verified and compared in the following rigorous dynamic simulations to further rank the effectiveness and performance of different alternative control pairing options.

Fig. 2. Open-loop sensitivity analysis of the columns C1 (a) and C2 (b). Table 1. The results for facing ± 0.1% variation of design values for SVD and RGA analysis criteria. SVD Analysis + 0.1% changes

K

U

VT



-0.1% changes

FS

QR

FS

QR

T67

43.0167

1294.9310

T76

189.4640

208.9484

T67

-28.3076

-921.388

T76

-190.935

-214.445

T67

-0.9860

T76

-0.1669

-0.1669

T67

-0.9705

-0.2409

0.9860

T76

-0.2409

0.9705

FS

QR

FS

QR

-0.0564

0.9984

0.0774

-0.9970

-0.9984

-0.0564

0.9970

0.0774

1.3137

948.7618

0.1799 CN

179.0279

7.3024

5.2995 RGA Analysis

FS ∧

QR

FS

QR

T67

-0.038

1.038

T67

-0.036

1.036

T76

1.038

-0.038

T76

1.036

-0.036

Besides, the slope criterion, the earliest and most frequently used method, was taken to choose the sensitive tray temperature location(s) in the previous work (Ma et al., 2019). When temperature profiles are obtained at steady-state design conditions, the position of sensitive temperature tray(s) can be acquired by the slope of curve with the maximum value. As obtained from Fig. 3, the slope of Stage 32 is the largest for column C2. However, the conditions of extractive column C1 are complicated, which has multiple extreme points including Stages 76, 57 and 32. And Stages 57 and 32 are the false sensitive locations generated by the cold inlet feed streams. It is important to note that an important result is not obtained that the Stage 67 is also a sensitive temperature should be controlled when compared with the previous sensitivity criterion. Because the objective of the section from the Stage 55 to Stage 73 is to keep acetone going down to the bottoms, therefore, the large tray temperature changes (Stage 67) should be controlled to alleviate the effect of acetone component in the side-stream on the product methanol purity of column C2.

Fig. 3. Temperature difference (a) and temperature (b) profiles for Columns C1 and C2.

The conventional proportional-only controller with the gain (Kc) of 2 and integral time (τI) of 9999 min is employed to all level control loops. The proportional-integral temperature and composition controllers are used to the regulatory control loops, and their ultimate periods and frequencies are determined by the relay-feedback testing procedures and Tyreus-Luyben tuning rules. The dead-time blocks of 1-min and 3-min are inserted to the temperature and composition control loops. And the proportional-integral controllers are also used to control the throughput and operation pressures with the controller parameters of Kc = 0.5/τI = 0.3 min and Kc = 20/τI = 12 min, respectively. In the initial control structure CS1-1, shown in Fig. 4, the detailed control loops with their paired controlled and manipulated variables are enumerated below. 1. The feed is flow controlled, which is inserted to the throughput disturbances (reverse action).

2. The condenser heat removals are manipulated to control the pressures (reverse action). 3. The manipulation of distillate flowrates is to control the reflux drum levels (direct action). 4. The manipulation of bottoms flowrate in column C2 is to control the base level (direct action). 5. The base level of column C1 is controlled by adjusting the solvent makeup flowrate (reverse action). 6. The temperatures of Stages 67 and 33 are controlled by manipulating the feedforward ratio of reboiler heat input to feed flowrate of columns C1 and C2, respectively, and the Stage 76 temperature is controlled by manipulating the liquid side stream flowrate. 7. The reflux ratios are fixed for both columns. 8. The set-point of the bottoms recycling flow controller of extractive column C1 is manipulated by the ratio of solvent to feed flowrate (reverse action). 9. The cooler heat removal is manipulated to control the temperature of recycle entrainer stream to 320 K (reverse action). Herein, several control loops should be underlined in this control structure. The first is that the feedforward action (reboiler heat duty to feed flowrate ratio) is used to alleviate the large product transient deviation and improve the response performance; The second is the control loops of the Stage 76 temperature pairing with the side-stream flowrate, and it sends the signal to side-stream flowrate controller operated on “cascade” to vary with disturbance. Dynamic response results for facing feed flowrate disturbances are illustrated in Fig. 5(a). The solid and dashed lines are, respectively, for 20% increases and decreases. The effective product quality control is achieved with the dual-product composition held closed to their initial steady-state design specifications. All temperatures are well controlled to quickly return to their design values. The settling time for this system is around 7 h at inserting disturbances at 1 h. There is a product offset for distillate methanol purity, and the ultimate stable composition is around 99.30 mol% for facing plus 20% step change in throughput, while another product composition is rigorously close to the initially steady-state design value for handling plus and minus 20% step changes in production rate. Fig. 5(b) gives the dynamic response results for facing feed composition disturbances. The solid lines are for feed acetone composition increase from 50 mol% to 60 mol% with the corresponding reduction of methanol composition in fresh feed stream, and the dashed lines are for 20% step decreases with the relevant augment of methanol composition in fresh feed stream, which is for the feed acetone composition decrease from 50 to 40 mol%. Good product composition control is achieved with the dual-product composition held closed to their initially steady-state design specifications, especially for product acetone in extractive column C1. All temperatures are well controlled to quickly go back to their initially steady-state design values. The settling time for this system is around 7 h at inserting disturbances at 1 h. There is also a product offset in methanol composition deviated from the corresponding steady-state designs. And the final stable value is around 99.30 mol% methanol.

Fig. 4. Basic control structure (CS1-1) for conventional side-stream extractive distillation process.

Another alternative sensor-valve pairing control scheme (denoted as CS1-2) is shown in Fig. S1 of Supporting Information. The special control loops for this control scheme are that the temperatures of Stages 67 and 76 are controlled by manipulating the side-stream flowrate and reboiler heat duty, respectively, which is the inverse configuration of initial control structure CS1-1. Dynamic response results for handling plus and minus 20% step changes in throughput and feed composition are given in Fig. S2 (a) and (b) of Supporting Information. The stable regulatory control is also achieved, while there are large products transient maximum peaks, especially for product methanol of column C2, in comparison with the initial control structure CS1-1. The dynamic control behavior comparisons of different control strategies with alternative sensor-valve pairing options are demonstrated by taking the integral absolute error (IAE) (Pan, et al., 2019; Zhang, et al., 2020a, b) as the evaluation indicator, which represents the difference of numerical integration (area) between product composition response values and the corresponding design specifications after adding interferences, and the results are enumerated in Table 2. And the direct comparisons of initial control structures CS1-1 and CS1-2 are shown in Fig. 6 with handling the same disturbances of throughput and feed composition. The performance of initial control scheme CS1-1 is superior to that of another alternative strategy owing to the less transient product composition deviation and smaller settling time with the same product offsets. Besides, the dynamic response performance of this initial control structure CS1-1 is also superior to that of the optimal control strategy devised by Ma et al. (2019) in terms of settling time, transient deviations and system stability. Of the particular importance is that it is difficult for the control structure developed by Ma et al. (2019) to robustly control especially for facing 20% decrease feed composition disturbance since the product methanol is fluctuated at near 15 h without stable around the steady-state design values. Therefore, our developed control structure is robust and efficient for facing 20% feed disturbances. As obtained from Fig. 5, there is a product offset for distillate methanol composition in column C2, which can be eliminated or attenuated by the application of the direct composition control loops. And the corresponding intensified control scheme CS2 is developed and demonstrated in Fig. 7. A feature control loop of this control structure CS2 is that product methanol composition is controlled by manipulating the reflux ratio of column C2.

Besides, the remaining control loops are the same as the initial control structure CS1-1. Fig. 8 shows the dynamic response results for this improved control structure CS2. The implication of the solid and dashed lines is the same as the above control schemes. Fig. 8(a) gives the results of this control structure CS2 for handling plus and minus 20% step changes in throughput. These interferences are well handled with the stable regulatory control achieved. Dual-product purity is returned to their initially steady-state specifications. And all controlled temperatures are also going back to their initially steady-state values. Therefore, this control scheme is efficient and robust for rejecting the large throughput disturbance. Responses of feed acetone composition changes from 50 to 60 mol% and 50 to 40 mol% are demonstrated in Fig. 8(b). The efficient product quality control is also achieved with the dual-distillate composition held quite closed to their initially steady-state designs. And all three temperatures are also well controlled to quickly come back to their initial designs. Therefore, this improved control structure CS2 is efficient and robust at handling the large feed composition perturbations. The detailed tuning parameters about the regulatory controllers for control structure CS2 are enumerated in Table S1 of Supporting Information.

Fig. 5. Dynamic response results of basic control structure CS1-1: (a) ± 20% throughput disturbances; (b) ± 20% feed composition disturbances.

Table 2. Dynamic response performance comparisons for different alternative control structures. 20% Throughput disturbance Items

CS1-1

CS1-2

CS1-2 (+20%)

CS2

0.00524

0.00762

0.00350

0.00532

IAExD2,MeOH

0.05271

0.06462

0.00682

0.00690

0.04014

IAESum

0.05795

0.07223

0.01032

0.01222

0.54530

IAExD1,

CS3 (Ma et al.) 0.50516

Acetone

20% Feed composition disturbance Items

CS1-1

CS1-2

CS2 (+20%)

CS2

0.00856

0.00913

0.00668

0.00853

IAExD2, MeOH

0.04374

0.06462

0.00325

0.00333

0.04145

IAESum

0.05230

0.07375

0.00993

0.01186

0.97146

IAExD1,

CS3 (Ma et al.) 0.93001

Acetone

Fig. 6. Dynamic response comparisons for alternative sensor-valve pairing options: (a) 20% throughput disturbances; (b) 20% feed composition disturbances.

Fig. 7. Improved control structure (CS2) for conventional energy efficient side-stream extractive distillation process.

Fig. 8. Dynamic response results of improved control structure CS2: (a) ± 20% throughput disturbances; (b) ± 20% feed composition disturbances.

2.3 Dynamic control behavior comparisons for different alternative control schemes The dynamic control behavior comparisons for different control alternatives CS2 and control structure devised by Ma et al. (denoted as CS3) are illustrated in this section, as shown in Fig. 9. Besides, the directly intuitive and quantitative analysis method IAE is also used to examine these dynamic response performances, and the results are shown in Table 2. What the conclusion is that the effectiveness and robustness of control structure CS2 is greatly superior to that of control structure CS3 in terms of settling time, transient deviation and product offsets and system stability. Dynamic response performance comparisons for facing 20% increase of throughput disturbance are illustrated in Fig. 9(a). The solid lines are for the dynamic response results of control scheme CS2, and the dashed lines are for the response performance of control structure CS3. What the conclusion is that the dynamic response performance of control structure CS3 is poor in comparison with the control structure CS2. As obtained from Fig. 9(a), the system is failed to achieve the stable regulatory control for control structure CS3 owing to product methanol composition oscillated at near 17 h without converging to the initially steady-state design value. Moreover, the system is not work after 2 h of added interference for control structure CS3 owing to the side-stream flowrate dropping to zero when inserting minus 20% step change in throughput at 1 h. In other words, the corresponding results obtained from Ma et al. are not achieved in our repeated simulations. Dynamic response performance comparisons for handling minus 20% step change in feed acetone composition are illustrated in Fig. 9(b). Dynamic response performance of control structure CS3 is also poor in comparison with control structure CS2. As obtained from Fig. 9(b), the system is also failed to achieve the stable regulatory control for control structure CS3 owing to product methanol composition oscillated at near 18 h without converging to the initially steady-state design value. Moreover, the system is also not work after 2 h of added disturbances for control structure CS3 owing to the side-stream flowrate dropping to zero when inserting 20% increase feed acetone composition disturbance at 1 h. The corresponding results obtained from Ma et al. are also not achieved in our repeated simulations.

Fig. 9. Comparisons of dynamic response results for different control strategies: (a) +20% throughput disturbance; (b) -20% feed composition disturbance.

3. Energy-efficient Heat-integrated Side-stream Extractive Distillation Configuration The performance of the conventional side-stream extractive distillation configuration is enhanced by the application of the heat integration technology. In that study (Tututi-Avila et al. 2017b), the steady-state heat-integrated side-stream extractive distillation configuration was also studied, and it can achieve the reductions of 26.98% in total annual cost, 29.34% in total energy consumption rates and 29.34% in carbon footprints in comparison with the conventional side-stream extractive distillation process, while the corresponding dynamic controllability for this economic attractive alternative is also not explored in that study. And the exploration of the dynamic control behavior analysis for this intensified option has not found in the open literature. That’s also the purpose of this work. 3.1 Process description The flowsheet of the energy efficient heat-integrated side-stream extractive distillation configuration is shown in Fig. 10. The operation pressure of column C2 is increased from 1 atm to 8 atm to maintain the feasible heat- transferring temperature differential driving force between the reboiler of column C1 and the condenser of column C2 with that of the extractive column C1 changed from 1 atm to 0.8 atm under the premise of using cooling water as cold utility for condenser. The column configuration parameters (total stages and feed locations) are the same as the conventional process. It should be noted that the bottoms water composition is not obtained to 99.99 mol% for column C2. The corresponding bottoms composition is 99.75 mol% water by the sensitivity analysis when the composition of recycle solvent stream is set at 99.90 mol% water holding bottoms composition constant of extractive column C1. The reflux ratio is increased from 0.974 to 1.75 for column C2 and the reboiler heat duty is increased from 5.336 MW to 7.777 MW to achieve this specifying separation in column C2. These results are the same as the previous work (Tututi-Avila et al. 2017b) with the exception of bottoms water composition. The temperature-enthalpy (T-H) diagram for this efficient configuration is shown in Fig. 10, which is obtained by the assumption of isothermal operation of condensers and reboilers (Smith, 2005). The “box” represents the column, which is not rectangles since the feed is not the saturated liquid or the heat loss to ambient across the wall. The interval among the adjacent lines of two boxes indicates the temperature difference between the condenser of column C2 and the reboiler of extractive column C1. The efficient operation is conducted when the interval is larger than 20~30 K (Zhang et al., 2019; Luyben, 2013). The shaded zone illustrates the degree of heat integration. As clearly illustrated from T-H diagram in Fig. 10, the heat-integrated operation is feasible. A 67.28% portion of the heat requirement in reboiler (6.253 MW) in extractive column C1 is provided by the condenser heat removal in column C2, and the auxiliary steam-driven reboiler (3.041 MW) is used to compensate the deficiency of energy requirement in extractive column C1. The detailed information for the heat-integrated side-stream extractive distillation process is shown in Fig. 10.

Fig. 10. Flowsheet of an energy efficient heat-integrated side-stream extractive distillation process and the corresponding T-H diagram.

3.2 Dynamic control As taught by the terminology of “simultaneous design”, the exploration of the dynamic controllability for this economic-attractive complex configuration is also an indispensable for ensuring the optimal and safe operation when taking into consideration the efficient operations obtained at the steady-state level. The control strategy is as the similar of conventional side-stream configuration, and it was also used in this complex alternative. And the detailed basic control scheme CS1 is shown in Fig. 11. The feedforward action of reboiler heat duty to feed flowrate ratio is also applied to this complex configuration to improve the control performance, such as reduce transient deviation(s) for product compositions. The constructed procedures of feedforward action are the same as the conventional configuration for column C2, while it is not trivial for extractive column C1. The total heat requirement for achieving the specifying separation for extractive column C1 is calculated by the summation of the steam flowrate in the auxiliary reboiler and the overhead vapor flowrate from column C2 multiplied by their appropriate latent heats of vaporization. This signal is fed into a “total Q” controller that adjusts the steam valve on the auxiliary reboiler to keep the total heat input at its set-point. The set-point of the total Q controller is the output signal from the Stage 67 temperature controller. The detailed procedures are shown in Fig. 11, and the Aspen Dynamics implementation procedures are shown in Fig. 12. Dynamic response results of this control structure CS1 for facing throughput disturbances are shown in Fig. 13(a). The solid and dashed lines are, respectively, for 10% increases and decreases. The effective product quality control is achieved with the dual-distillate composition held closed to their initial steady-state design specifications. All temperatures are also well controlled to quickly return to their initial design values. The settling time for this system is around 6 h at inserting disturbances at 1 h. There is also a product offset in methanol composition deviated from the steady-state designs. And the final stable value is around 99.43 mol% and 99.57 mol% methanol, respectively, for facing positive and negative disturbances. Dynamic response results for facing feed composition disturbances are illustrated in Fig. 13(b). The solid lines are for feed acetone composition increase from 50 mol% to 55 mol% with the corresponding reduction of methanol composition in fresh feed

stream, and the dashed lines are for feed acetone composition reduction from 50 mol% to 45 mol% with the corresponding increase of methanol composition in fresh feed stream. The stable regulatory control is also achieved with the dual-distillate composition held closed to their initial steady-state design specifications. All the temperatures are well controlled to quickly go back to their initial designs. The settling time for this system is around 6 h at inserting disturbances at 1 h. There is also a product offset in methanol composition deviated from the steady-state designs. And the final stable value is around 99.54 mol% and 99.46 mol% methanol, respectively, for facing positive and negative disturbances. These deviations are smaller in comparison with the throughput disturbances. This product offset can be eliminated or attenuated by the application of the direct composition control loops. And the corresponding intensified control scheme CS2 is developed and demonstrated in Fig. 14. A feature control loop of this control structure is that product methanol purity is controlled by manipulating the reflux ratio for column C2. Besides, the rest control loops are the same as control structure CS1. Dynamic performance response results for improved control structure CS2 are demonstrated in Fig. 15. The solid lines represent the responses for positive increment of disturbances, while the dashed lines show the response for negative changes. Responses of throughput from 540 to 594 kmol·h-1 and 540 to 486 kmol·h-1 are shown in Fig. 15(a). These interferences are well handled with the stable regulatory control achieved. Both product purities are closely returned to their steady-state design specifications. All controlled temperatures are also quickly going back to their initially steady-state values. The settling time for this system is around 6 h at inserting disturbances at 1 h. Therefore, this control scheme is efficient and robust for rejecting large throughput disturbance. Responses of feed acetone composition from 50 to 55 mol% and 50 to 45 mol% are demonstrated in Fig. 15(b). The solid and dashed lines are for 10% increases and decreases, respectively. The efficient product quality control is achieved with the dual-product composition held quite closed to their initially steady-state design specifications. All three temperatures are also well controlled to quickly come back to their initial designs. The settling time for this system is around 6 h at inserting disturbances at 1 h. Therefore, this improved control structure CS2 is efficient and robust at facing large feed composition perturbations. The detailed tuning parameters about the regulatory control loops for intensified control structure CS2 are enumerated in Table S2 in Supporting Information.

Fig. 11. Basic control structure (CS1) for the economic-attractive heat-integrated side-stream extractive distillation process.

Fig. 12. Aspen Dynamics implementation procedures for the feedforward action by the manipulation of the auxiliary reboiler heat duty QAUX.

Fig. 13. Dynamic response results of basic control structure CS1: (a) ± 10% throughput disturbances; (b) ± 10% feed composition disturbances.

Fig. 14. Improved control structure (CS2) for the economic-attractive heat-integrated side-stream extractive distillation process.

Fig. 15. Dynamic response results of improved control structure CS2: (a) ± 10% throughput disturbances; (b) ± 10% feed composition disturbances.

4

Conclusions A new control structure was proposed for the eco-efficient side-stream extractive distillation configuration

with and without heat integration taking the separation of acetone and methanol azeotrope using water as solvent as the example and was assessed by introducing the large throughput and feed composition disturbances. The effectiveness of dynamic controllability was enhanced by the application of the direct composition control loops in column C2 with dual-temperature control in extractive column C1. Good product-quality control was achieved with the key control loops of controlling the temperatures of Stages 76 and 67 by manipulating the side-stream flowrate and reboiler heat duty of extractive column C1. Furthermore, the dynamic response performance was superior to the control structure devised by Ma et al.

Acknowledgment We are grateful to the comments and suggestions from the editor (Prof. de Haan) and the anonymous reviewers. These will have a great positive effect on my future research works by studying your comments carefully. Special thanks again to you for your good comments.

Nomenclature EDWC

Extractive dividing-wall column

SEDC

Side-stream extractive distillation

TAC

Total annual cost

RR

reflux ratio

QR

reboiler heat duty

QR/F

Reboiler heat duty-to-feed flowrate

Kc

Controller gains

τI

Integral time

TC

Temperature controller

CC

Composition controller

PC

Pressure controller

FC

Flow controller

LC

Level controller

IAE

Integral absolute error

SVD

Singular value decomposition

RGA

Relative gain array

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Graphical Abstract:

Highlights: 1. Dynamic control analysis of eco-efficient side-stream extractive distillation process is studied. 2. A newly robust control scheme is developed for this complex configuration. 3. Dynamic control performance is improved by using direct composition control loops. 4. Dynamic response performance is superior to that given in a published paper.

Conflict of Interest Form The authors declared that there are no conflicts of interests. We declare that we do not have any commercial or associative interest that represents a conflict of interest in connection with the work submitted.

Author Statement All authors have read and approved to submit it to your journal. There is no conflict of interest of any authors in relation to the submission. This paper has not been submitted elsewhere for consideration of publication.