Potential for Significant Energy-Saving via Hybrid Extraction

Article pubs.acs.org/IECR

Potential for Significant Energy-Saving via Hybrid Extraction− Distillation System: Design and Control of Separation Process for n‑Propanol Dehydration Wei-Lun Chang and I-Lung Chien* Department of Chemical Engineering, National Taiwan University, Taipei 10617, Taiwan S Supporting Information *

ABSTRACT: The hybrid extraction−distillation separation system is a process combining liquid−liquid extraction and also distillation. This hybrid system can be viewed as a derivative of the heterogeneous azeotropic distillation method. In this work, the potential for significant energy-saving via this hybrid process is demonstrated with both conceptual illustration and a case study of n-propanol dehydration. Diisopropyl ether (DIPE) is selected as the extraction solvent considering its favorable properties of density, heat of vaporization, and lower toxicity. Since the solvent flow rate is the most important design degree-of-freedom in this process that influences overall energy consumption, an iterative optimization procedure is conducted to minimize total annual cost. Furthermore, a novel control strategy is proposed on the basis of closedloop and open-loop sensitivity tests. Here, an adjustable solvent flow rate during dynamic control allows steady-state operation at the optimal condition. Dynamic simulation results show that both n-propanol and water products can still be maintained at high purities despite large variations in feed flow rate and feed composition disturbances.

1. INTRODUCTION The separation of azeotropic mixtures is an important task for both academic researchers and industrial companies. Various techniques for azeotropic separation, such as azeotropic distillation, extractive distillation, pressure-swing distillation, pervaporation, adsorption, extraction, etc., have been studied and implemented. Mahdi et al.1 reviewed the published literature about azeotropic separation including distillation processes, membrane processes, and process intensification techniques. Besides the technique of dividing-wall columns, hybrid separation processes are also a promising technique for the sake of energy-saving. The hybrid separation process is a strategy of combining two or more techniques to separate the azeotropic mixtures. In this work, a hybrid extraction− distillation system is the focus since extraction is a separation method that requires no energy usage. It is expected to remove one component in the original mixtures via extraction, and then high-purity products can be obtained via further distillation. The solvent added plays a key role on how the hybrid design configurations will become. Errico et al.2 proposed a hybrid extraction−distillation system for biobutanol purification. Also, there are many experimental results, reviewed by Pereiro et al.,3 showing that ionic liquid as solvent can be applied to liquid− liquid extraction. However, in those systems, all their solvent species are the heaviest component. This means some of the fresh feed should be vaporized to become the distillate products, representing a necessary heat input to the system. Also, the heaviest solvent species are not allowed to form any © XXXX American Chemical Society

new azeotropes with other components; otherwise, the further distillation for purification may not be feasible. On the other hand, Kürüm et al.,4 Lucia et al.,5 and Chen et al.6 have successfully utilized low-boiling solvent species, like methyl-tertbutyl ether and ethyl acetate for acetic acid dehydration, or npropyl formate and diisopropyl ether (DIPE) for pyridine dehydration via hybrid extraction−distillation. These species can form a new minimum boiling heterogeneous azeotrope in the system. The resulting design configurations can be viewed as a derivative from a heterogeneous azeotropic distillation technique. This argument will be explained in this work. Most importantly, the potential for significant energy-saving via a hybrid extraction−distillation system with low-boiling solvent species will be demonstrated in this paper with conceptual illustration and with a new case study of n-propanol dehydration. n-Propanol (NPA) forms a low-boiling azeotrope with water, making the separation unable to carry out with regular distillation. An et al.7 proposed an extractive distillation system with N-methyl-2-pyrrolidone (NMP) as entrainer, whereas PlaFranco et al.8 selected ethylene glycol as entrainer. NMP enhances more in the relative volatility than ethylene glycol does. Except for the extractive distillation, there are only some Received: June 8, 2016 Revised: September 15, 2016 Accepted: October 13, 2016

A

DOI: 10.1021/acs.iecr.6b02210 Ind. Eng. Chem. Res. XXXX, XXX, XXX−XXX

Article

Industrial & Engineering Chemistry Research

Figure 1. (a) Conceptual design flowsheet of heterogeneous azeotropic distillation. (b) Material balance lines of heterogeneous azeotropic distillation.

literature studies9−13 about the vapor−liquid−liquid equilibrium or vapor−liquid equilibrium experimental results with regard to this system of NPA/water. Thus, the extractive distillation system in An et al.7 will serve as the original design for the purpose of comparison. They proposed a three-column conventional extractive distillation system using NMP as entrainer. The preconcentration column and the extractive distillation column were further combined as a two-column extractive distillation system with a side-reboiler. However, because the enhancement in relative volatility is not that great, an entrainer-to-feed ratio of about 2.2 and total column stages over 92 stages were used in these designs. This work is contributed to implement a hybrid extraction−distillation system for NPA dehydration, which is expected to significantly reduce energy consumption. Chen et al.6 proposed a control strategy for this kind of hybrid system for pyridine/water separation. However, with a fixed solvent-to-feed ratio during dynamic control, a trade-off needs to be made between optimal steady-state design and

dynamic controllability to reject the composition disturbance from a fresh feed. A conservative increased solvent-to-feed ratio needs to be operated rather than at the optimal ratio, which represents a 9.21% increase in total annual cost. The other contribution of this work is to devise a novel control strategy which allows the steady-state operation right at the optimal condition but is still able to fulfill dynamic controllability.

2. HYBRID EXTRACTION−DISTILLATION SYSTEM 2.1. Derivative from Heterogeneous Azeotropic Distillation. With NPA dehydration taken as an example, a traditional conceptual design flowsheet of heterogeneous azeotropic distillation (HAD) and its material balance lines are shown as Figure 1a,b. The two-column design is composed of an azeotropic column (C-1) with a top decanter, and another column (C-2) served as the preconcentration column for fresh feed and also served as the recovery column for an aqueous outlet stream. High-purity NPA and water can be obtained at the bottoms of the C-1 and C-2 columns, respectively. B


Article


Figure 2. (a) Conceptual design flowsheet of hybrid extraction−distillation. (b) Material balance lines of hybrid extraction−distillation.

A hybrid extraction−distillation process can be viewed as a derivative from heterogeneous azeotropic distillation, because they both require adding an appropriate third component to form a new minimum boiling heterogeneous azeotrope and require a large liquid−liquid envelope, as shown in Figures 1b and 2b with the same third component, DIPE, for illustration. The difference is that the requirement of a preconcentration column in HAD process is replaced in the hybrid process with an extraction column. In the HAD process, composition of the overall feed into the azeotropic column is taken to a distillation region where NPA is at the stable node by mixing organic reflux and preconcentrated feed. However, this task is accomplished

Figure 2 shows the conceptual hybrid extraction−distillation process flowsheet and the material balance lines. The fresh feed is sent into an extraction column with a solvent-rich stream from the decanter. Then, the extract phase is sent into an azeotropic column (C-1) to obtain a bottom product of NPA. The heterogeneous overhead vapor is condensed and performs liquid−liquid separation in a decanter, where the organic stream is served as the solvent stream recycled into the extraction column. Besides, the raffinate phase from the extraction column and the aqueous stream are sent into a recovery column (C-2) to obtain high-purity water as a bottom product. C


Article


Figure 3. Important material flows in (a) heterogeneous azeotropic distillation; (b) hybrid extraction−distillation.

process, fresh NPA is sent into the preconcentration column, vaporized, and condensed out of the top of this column as a D2 stream. Then, the D2 stream is sent into the azeotropic C-1 column to obtain the bottom product of NPA. Thus, phase change through this pathway represents significant heat input in the HAD process. On the other hand, in the hybrid process, fresh NPA is extracted in the extraction column and finally becomes the bottom product of the C-1 column. It always remains as the liquid phase, which means only that sensible heat is needed in the hybrid process. As an improved design of the HAD process, most of the heat input to vaporize the D2 stream in the HAD process could be reused in the C-1 column by replacing the C-2 column with a stripper. However, the drawback is that there is no ability to obtain a more favorable distillate composition of this C-2 column. Chang et al.14 and Wu et al.15 had demonstrated this concept for isopropyl alcohol and pyridine dehydration, respectively. Thus, the pathway of fresh NPA is not considered one of the main factors to contribute significant energy-saving for the hybrid process. Second, the blue arrows show how the water, accompanying NPA into the azeotropic C-1 column, is dealt with to become an aqueous outlet from the decanter. In the HAD process, because of the binary NPA/water azeotrope, it is inevitable to vaporize some water at the same time when NPA is vaporized

in the hybrid process through extracting fresh feed with the solvent stream. Note that the solvent-rich stream into the extraction column comes from the organic reflux of the heterogeneous overhead of a C-1 column, while the organic reflux is directly sent back into the C-1 column in the HAD process. Most importantly, the solvent flow rate is an additional design degree-of-freedom (DOF) that can significantly affect the energy consumption in the hybrid extraction−distillation process, while this DOF is not available for HAD process. 2.2. Potential for Energy-Saving. In the above section, the difference in design configurations between HAD and hybrid processes is discussed. Here, in this section, it is illustrated how the difference in design configurations leads to energy-saving by focusing on the phase change through these processes. This is because reboiler duty is provided for the enthalpy change of top vapor overhead plus bottom product in the form of latent heat, sensible heat, or heat of mixing according to energy balance of the reboiler plus the column. In the above contributions of reboiler duty, the required latent heat is especially different between these two processes. Figure 3a,b shows the three most important material flows that are quite different in HAD and hybrid processes. First, the red arrows point out how NPA in the fresh feed becomes the bottom product of the C-1 column. In the HAD D


Article

Industrial & Engineering Chemistry Research Table 1. Properties of Several Pure Solvent Candidates solvent candidate

chloroform

TAME

benzene

trichloroethylene

DIPE

isopropyl acetate

density (kg/cum) heat of vaporization (kJ/mol) LD50 [rat, oral] (mg/kg) IARC category

1441.602 30.166 300 2B

744.895 33.131 1602−2100

850.729 32.656 930 1

1420.671 33.400 4920 2A

697.608 30.699 8470

843.592 35.635 6160−7380

ternary diagram into at least two distillation regions, where pure water and NPA are at the stable nodes, respectively. These above criteria are similar in the HAD process. However, in hybrid process, there are two additional necessities to make it feasible. One is that the solvent should have stronger affinity to NPA than water, so that NPA can be extracted from fresh feed in the extraction column. That is, the distribution factor of NPA must be greater than 1. This can be observed in a ternary diagram with LLE tie-lines tilted more to the water corner. The other is that the liquid−liquid envelope is large enough for the extract phase composition to be located in the distillation region where pure NPA is at the stable node, so that highpurity NPA can be separated from the extract phase by further distillation. Several favorable solvent candidates are taken from Kürüm et al.4 and Gmehling and Möllmann.16 The ternary diagrams formed with all of these candidates satisfy the above requirements. In Table 1, some additional properties for these solvent candidates are listed, including density, heat of vaporization, lethal dose 50% (LD50), and IARC category. First, a good solvent should provide a large density difference between two liquid phases for an easier liquid−liquid separation. Pure densities of NPA and water are 780.271 and 985.859 kg/cum, respectively. Thus, benzene and isopropyl acetate are eliminated, since their pure densities are too close to NPA and water and that may lead to insufficient density difference for liquid−liquid extraction. Second, because the solvent stream is vaporized in the azeotropic column, lower heat of vaporization is favorable for energy-saving. These candidates are similar in the heat of vaporization, so none of it is eliminated. Next, because of the awareness of safety and health concerns, toxicity is an important factor for any industrial process. Here, LD50 and IARC categories are taken into account, where these data are acquired from the website of the Occupational Safety and Health Administration in Taiwan. The IARC category is proposed by International Agency for Research on Cancer based on the strength of evidence for human carcinogenicity. For chloroform and trichloroethylene, they are categorized into groups 2B and 2A, which means that they are probably carcinogenic to humans. For benzene which is categorized into group 1, there is strong evidence that it can cause cancer in humans. Thus, these three candidates are not suitable. The remaining DIPE has greater LD50 (less toxicity) than TAME, so it is finally selected as solvent for the next simulations of hybrid design. 3.2. Overall Design Flowsheets (Fixed Column Stages). The conceptual design flowsheet is mentioned in Section 2.1 and Figure 2a. The fresh feed is at 300 kmol/h (20% NPA and 80% water) as is the same in An et al.7 It is fed into a five-stage equilibrium and adiabatic extraction column. Afterward, the extract phase is sent into a 10-stage stripper (C1), while the raffinate phase is sent into another 10-stage stripper (C-2). Top vapor overhead from each stripper is condensed and subcooled to 320 K so that cooling water can be used as utility here. Purity of NPA product is set at 99.9 mol %

out of the C-2 column. For example, if NPA is preconcentrated to 40%, that means the other 60% content, mostly water, should be vaporized as well. This water sent into the C-1 column should be vaporized again and then becomes the aqueous stream out of the decanter. Even with a C-2 as stripper, water still needs to be vaporized and condensed at least once per cycle through the HAD process. On the other hand, in the hybrid process, most of the water is separated into the raffinate phase in the extractor. There is only a small amount of water undesirably extracted and sent into the C-1 column, which is then vaporized and becomes an aqueous stream (AQ) out of the decanter. The difference in the amount of this water could be considered as one of the major factors contributing to energy-saving. Thus, this hybrid process will be quite competitive for the separation of diluted solution and mixtures that could not be preconcentrated well. Finally, the green arrows represent the pathways of organic reflux. In the HAD process, as mentioned above, the organic reflux flow rate is not a degree-of-freedom. The purpose of this organic reflux is to carry all of the water fed into this column to the top. Thus, the flow rate depends on the azeotropic composition and the amount of water fed in (or extent of preconcentration). In Figure 1b, DIPE is selected as the entrainer in HAD process just for a clear comparison with the hybrid process. In fact, it is not an appropriate entrainer for the HAD process because the heterogeneous azeotropic composition contains too little water, which will lead to a large amount of organic reflux in the system. On the other hand, organic reflux flow rate is another degree-of-freedom for the hybrid process that can be optimized. Thus, organic reflux could also be considered as one of the major factors in energy-saving because it is vaporized and condensed through these processes. The extent of energy-saving depends especially on the entrainer or solvent selection. As for the difference in energy consumption between extractive distillation and the hybrid extraction−distillation process, it is hard to illustrate because the designs are based on totally different concepts. There are two major disadvantages for extractive distillation. First, all components in the fresh feed need to be vaporized and condensed as products in distillate streams. Also, it requires high-grade steam as the utility in the reboilers or else vacuum systems are required because of the entrainer as the heaviest component of the system.

3. STEADY-STATE SIMULATIONS OF HYBRID PROCESS 3.1. Selection on Solvent Species. A nonrandom twoliquid (NRTL) thermodynamic model is selected to describe the nonideal phase behavior. The detailed information about the binary parameters is available in the Supporting Information. Solvent selection is one of the most important parts in the design of hybrid process. By adding a third component, a heterogeneous minimum boiling azeotrope should be formed. Also, the internal distillation boundaries need to split the E


Article


Figure 4. (a) Reboiler duty of each column. (b) Total reboiler duty varying with solvent flow rate of both configurations.

should be different for different solvent species, too. It is unfair to compare these differences under the same solvent flow rate since their minimum and optimal flow rates are not the same. Also, it will be quite time-consuming to obtain the optimal operation and also the difference in energy consumption with different solvent species, if one simulates again and again under different solvent flow rates and with different solvent species. In a comparison of Figures 5b and 6b, the optimal material balance lines of the C-1 column and the results are almost the same regardless of different design configurations. The difference in their minimum solvent flow rates does not come from the C-1 column but from the V2 stream. This represents that the optimal operation occurs when the extract phase is at that composition regardless of configurations, which is favorable for further distillation. At this extract phase composition, the lowest amount of overhead needs to be vaporized from the C-1 column to obtain roughly the same production rate of high-purity NPA as NPA in the fresh feed, resulting in a minimum recirculated solvent flow rate and also minimum energy consumption. Moreover, the optimal extract phase composition is observed to be right near the intersection of liquid−liquid envelope and isovolatility curve of water/DIPE as shown in Figure 7a. This generalization of the optimal extract composition was also observed for other solvent species in Table 1 with many simulations to be at the intersection of liquid−liquid envelope and isovolatility curve of water/solvent. Thus, as long as the extract phase composition is kept at this intersection (varying with solvent species), it is fair and easier to compare different solvent species with regard to optimal energy consumption. However, for some other ternary systems, there may be a limitation from extraction which prevents the extract phase composition from trending toward this intersection. Take pyridine dehydration in Chen et al.6 as an example. As shown in Figure 7b, the optimal extract phase is far from this intersection. Instead, it is bounded because of poorer extraction perform-

by varying the reboiler duty in C-1 (QR1), and purity of water product is also set at 99.9 mol % by varying the reboiler duty in C-2 (QR2). In Chen et al.,6 they sent V2 stream back into the decanter because the composition is within liquid−liquid envelope, while V2 is sent into the extraction column as in Figure 2a in this paper because it has a similar composition as a fresh feed. This slight variation of design configurations will be simulated to check which one is better. Although Chen et al.6 optimized the total stages of three columns and also optimized the solvent flow rate, their results show that those total stages really have a small effect on total annual cost as compared to the effect of solvent flow rate. Thus, total stages of three columns are fixed in this section, while the solvent flow rate is optimized under an objective function of total reboiler duty. Figure 4a,b shows how solvent flow rate influences the reboiler duty. In addition, Figure 5a shows the overall design flowsheet at an optimal solvent flow rate if V2 is sent into the extraction column, while Figure 6a shows the overall design flowsheet at optimal solvent flow rate if V2 is sent into the decanter. Material balance lines on ternary diagrams for both configurations are shown in Figures 5b and 6b, respectively. Both design configurations hit a minimum solvent flow rate that leads to minimum total reboiler duty. The overall process cannot be converged with any other solvent flow rate below these minimums. As shown in Figures 4 and 6, since the solvent stream becomes more impure when V2 is sent into the decanter, the raffinate phase becomes more impure too, leading to higher QR2. Considering the total reboiler duty, it is apparently better to send the V2 stream back into the extraction column. Thus, the steady-state flowsheet shown in Figure 5a is served as the initial steady-state for the following dynamic simulations. 3.3. Generalization on Optimal Operation. As the results shown in Figure 4, the minimum optimal solvent flow rate is quite different for different design configurations, and it F


Article


Figure 5. (a) Optimal design flowsheet with V2 stream back into the extraction column. (b) Material balance lines of this configuration.

3.4. Economic Evaluation (Optimized Column Stages). Since total annual cost (TAC) is the most convincing index to evaluate process economics, economic evaluation and optimization are provided in this section. Detailed formula for the economics and equipment sizing are based on Seider et al.17 and Luyben,18 which are available in the Supporting Information. The iterative optimization procedure is provided in Figure 8. The solvent flow rate is the most important design variable so it is searched at all combinations of total stages, while column stages are also optimized to check their influence on TAC.

ance; that is, the extract phase is on a tie-line, whose extrapolation will cross the feed composition (10% pyridine plus 90% water). This tie-line will overlap the operating line of extraction, preventing the extract phase composition from getting closer to the isovolatility curve. The system will hit a minimum solvent flow rate and minimum energy usage at this limited extract composition before reaching the intersection. Thus, if the system may be limited by poorer extraction performance, the optimal extract phase composition should be located on the tie-line whose extrapolation crosses feed composition, rather than the intersection discussed above. G


Article


Figure 6. (a) Optimal design flowsheet with V2 stream back into the decanter. (b) Material balance lines of this configuration.

the solvent flow rate is the most important design variable for this hybrid process that should be carefully optimized.

Detailed TAC plots are available in the Supporting Information. The optimization results show the optimized column stages are 3, 6, and 12 for extractor, C-2, and C-1 column, respectively, while it is fixed at 5, 10, and 10, respectively, in Figure 5a. Table 2 summarizes the itemized TAC terms of two-column extractive distillation by An et al.7 and the hybrid processes studied in this work. Results show that a significant savings on TAC (62.19%) is realized via this hybrid process in Figure 5a as compared to two-column extractive distillation, while further optimizing the column stages only saves another 0.73% of the original TAC. It is concluded that

4. DYNAMIC CONTROL STRATEGY For this hybrid process, Chen et al.6 proposed a control strategy where two-stage temperature control and a fixed solvent-to-feed ratio were used. However, the optimal ratio cannot reject unmeasured feed composition disturbances. Thus, they proposed a trade-off to conservatively increase solvent-to-feed ratio for dynamic control. In this work, a novel control strategy is developed based on open-loop and closed-loop sensitivity tests with Aspen Plus steady-state simulations. The main H


Article


Figure 7. Optimal extract phase composition (a) near the intersection of isovolatility curve and liquid−liquid envelope; (b) on the tie-line whose extrapolation crosses fresh feed composition.

liquid phases. The holdup time of 40 min with 50% total liquid level is used. Then, the steady-state simulation of Figure 5a is exported to the dynamic simulation of Aspen Plus Dynamics. Since the extractor unit is not supported with pressure-driven simulations, flow-driven simulations in Aspen Plus Dynamics are selected with the top stage pressures of every column set at atmospheric pressure. Pressure drops inside the columns will be automatically calculated by Aspen Plus Dynamics. Although flow-driven simulations make an assumption of perfect flow control, which is usually accurate for liquid streams, it is still acceptable for a hybrid process, where there are only two vapor streams. 4.1. Inventory Control Loops. The intuitive regulatory control and inventory control loops are designed as follows. Fresh feed flow rate is flow-controlled with a PI controller of Kc = 0.5 and τI = 0.3 min. Both column base levels are controlled

objective is to allow the steady-state operation right at the optimal condition but still be able to fulfill dynamic controllability instead of a trade-off design. No online composition measurement will be used in the quality control loops for wider industrial applications. This control strategy is expected to hold the product purities at high-purity specifications despite any feed composition or throughput disturbances. The tray sizing option in Aspen Plus is utilized to calculate the column diameters with the assumed tray spacing of 0.6096 m and weir heights of 0.0508 m. The resulting column diameters are 1.091 m for the C-1 column and 0.409 m for the C-2 column. Other equipment sizing follows the recommendations by Luyben and Chien.19 Volumes of the column bases are sized to provide 10 min holdup with 50% liquid level. As for the decanter, it is sized bigger to allow the separation of two I


Article


the cooler duty. These cooling temperature PI controller parameters are set at Kc = 1 and τI = 20 min. After establishing the above regulatory control and inventory control loops, there are three remaining degrees-of-freedoms in this system, including reboiler duty of two columns and the solvent flow rate. Note that the flow rates of extract phase and raffinate phase cannot be specified or manipulated because there is no liquid level of extractors in Aspen Plus Dynamics. Next, in Section 4.2, it will be illustrated how the additional three control loops are determined for manipulating these three remaining degrees-of-freedom. 4.2. Quality Control Loops. A control strategy is desirable if it is able to hold the product purities at specifications despite any disturbances. Closed-loop sensitivity tests are conducted for the purpose of determination on controlled variables. It is a test which can be done with Aspen Plus steady-state simulations, assuming that the overall process will finally achieve perfect composition control after the anticipated disturbances are introduced to the system. Those process variables or ratios of two process variables that remain almost unchanged before and after the introduced disturbance are suitable to be selected as controlled variables. The process is tested under composition disturbance of ±25% NPA changes in the fresh feed, that is, increasing NPA from 20 mol % to 25 mol % or decreasing NPA from 20 mol % to 15 mol %, where water is decreased or increased to keep the total at 100 mol %. As mentioned in Section 4.1, there are three remaining degrees-of-freedom. However, the numbers of product purities of interest are just two, which means one of the degrees-offreedom should be treated before conducting closed-loop sensitivity tests. A new idea on treating this degree-of-freedom is conducted in this study, so a novel control strategy will be developed. As discussed in Section 3.3, there exists an optimal extract phase composition that leads to minimum solvent flow rate. The core concept is to hold the same optimal material balance line in the C-1 column regardless of any disturbance. That is, when NPA in fresh feed increases, the bottom NPA product, overhead vapor, and solvent flow rate need to be scaled up as well, at the same ratio of how much NPA increases. Thus, if NPA increases from 20 mol % to 25 mol % in fresh feed, the closed-loop sensitivity test is carried out at the solvent flow rate equal to 131.25 kmol/h (105 × 1.25). On the other

Figure 8. Optimization procedure (optimizing column stages).

by manipulating bottom flow rates separately. In the decanter, the aqueous phase level is controlled by manipulating aqueous outlet flow, while the organic phase level is controlled by manipulating solvent makeup flow. P-only controllers are used in level control loops. To speed up closed-loop control behavior in the recycle loop, the settings of Kc = 10 for the decanter level loops are used. Otherwise, Kc = 2 is used in the other level control loops. The pressure at the top of two strippers is controlled by manipulating their overhead vapor flow separately with PI controllers of default Kc = 20 and τI = 12 min for the purpose of tight pressure control. In addition, there are two temperature controllers controlling temperature of the condensed overhead stream at 320 K by manipulating Table 2. Itemized TAC Terms in Various Systems

two-column extractive distillation column capital cost reboiler capital cost condenser capital cost decanter capital cost extraction column cost side-reboiler capital cost cooler capital cost LP steam cost HP steam cost cooling water cost total reboiler duty total operating cost total capital cost TAC

×103 ×103 ×103 ×103 ×103 ×103 ×103 ×103 ×103 ×103 kW ×103 ×103 ×103

$ $ $ $ $ $ $ $/year $/year $/year $/year $ $/year

hybrid process (fixed column stages)

hybrid process (optimized column stages)

C1

C2

C1

C2

C1

C2

526.123 116.259 87.849

229.642 108.435 64.692

87.750 82.929 355.895

30.857 39.960 87.146

104.843 82.176 351.111

19.284 39.854 90.700

0 0 24.319 120.402 49.070 1459.853 54.046 5349.500 1562.969 1277.720 1988.876 (+0%) J

140.437 93.630 0 0 430.387 0 15.336 1920.820 445.723 918.603 751.924 (−62.19%)

138.108 64.647 0 0 425.403 0 15.151 1898.576 440.553 890.724 737.461 (−62.92%) DOI: 10.1021/acs.iecr.6b02210 Ind. Eng. Chem. Res. XXXX, XXX, XXX−XXX

Article

Industrial & Engineering Chemistry Research hand, if NPA decreases from 20 mol % to 15 mol % in fresh feed, another closed-loop sensitivity test is carried out at the solvent flow rate equal to 78.75 kmol/h (105 × 0.75). However, it is still necessary to identify three controlled variables from these tests, one of which is used to adjust the solvent flow rate during dynamic control, because of the unmeasured feed composition. Besides, open-loop sensitivity tests are performed in which ±0.1% change of reboiler duty is made. An appropriate temperature control point should have large sensitivity (large temperature deviation in open-loop sensitivity test) to the manipulated variables. Also, this control point should be almost unchanged (small deviation in closed-loop sensitivity test) for the sake of perfect composition control. In a comparison of the closed-loop and open-loop sensitivity tests on stage temperatures in the C-1 column, shown in Figure 9a,b, temperature on the sixth stage (T6) is the best temperature control point. Thus, the reboiler duty QR1 will be used to control the T6 temperature in the C-1 column.

Figure 10. Closed-loop sensitivity test on temperature profile in the C2 column.

Table 3. Brief Results in Process Variables from ClosedLoop Sensitivity Testsa original value (kmol/h) (kW)

SOL QR2/(AQ + RAF) SOL/QR1

deviation under perfect composition control

20%NPA + 80% water

25%NPA + 75%water

15%NPA + 85%water

105 1.788 0.07158

25.00% −0.12% −0.53%

−25.00% 0.05% 0.55%

a

Full results in process variables from closed-loop sensitivity tests are available in the Supporting Information. The following abbreviations apply: SOL, solvent flow rate; QR2, reboiler duty of C-2; AQ, aqueous outlet flow rate of decanter; RAF, raffinate phase flow rate; QR1, reboiler duty of C-1.

simultaneously when the temperature control manipulates the reboiler duty. Overall control strategy is shown in Figure 11. PI control is used in the temperature control loop containing 1 min of dead time. Note that, for better dynamic performance, an inner-loop ratio multiplier (QR1/FF) will adjust the reboiler duty as soon as the feed flow rate changes, and the set-point of this ratio is determined by the temperature control. Originally, the temperature controller is tuned by relay feedback test provided in Aspen Plus Dynamics with Tyreus−Luyben tuning rules.20 However, transient deviations in product purities are quite large at these tuning parameters. The PI control parameters are retuned to Kc = 3.26 and τI = 27.32 min with reverse action for achieving smaller transient deviations. 4.3. Closed-Loop Dynamic Responses. Two types of disturbances will be used to test the proposed control strategy. The tough disturbance of unmeasured fresh feed composition changes is considered first. Large variations include ±25% NPA changes in the fresh feed, that is, increasing NPA from 20 mol % to 25 mol % or decreasing NPA from 20 mol % to 15 mol %, where water is decreased or increased to keep the total at 100 mol %. Figure 12 shows the closed-loop dynamic responses to feed composition changes. It is observed that the temperature on sixth stage in C-1 column is quickly controlled back to its set point value. The solvent flow rate successfully increased or decreased according to the unmeasured feed composition changes. Note that the phenomenon of a small oscillation is needed to prevent a large transient deviation in NPA product purity. Although the NPA transient deviation is quite large, it only deviates within the first hour of the transient response.

Figure 9. (a) Closed-loop sensitivity test on temperature profile in the C-1 column. (b) Open-loop sensitivity test with reboiler duty (QR1) on temperature profile in the C-1 column.

Although all of the stage temperatures in the C-2 column do not deviate much in closed-loop sensitivity tests as shown in Figure 10, the ratio of reboiler duty in the C-2 column to raffinate plus aqueous flow rates (QR2/(RAF + AQ)) is also a good candidate to manipulate QR2 as shown in Table 3. Detailed information on all process variables from closed-loop sensitivity tests can be found in the Supporting Information. This ratio is selected as the second controlled variable so that no temperature control is required in the C-2 column. Also, shown in Table 3, the last controlled variable is the ratio of solvent flow rate to reboiler duty in the C-1 column (SOL/ QR1), which is applied to manipulate the solvent flow rate K


Article


Figure 11. Overall proposed control structure for hybrid extraction−distillation.

Figure 12. Closed-loop dynamic responses to feed composition changes.

Most importantly, both NPA and water product purities are maintained near their specifications in the long run despite these unmeasured disturbances.

The other tested disturbance is the 20% throughput change (feed ± 20%). Although throughput changes can be considered as a known disturbance, there is no need to adjust any set point L


Article


Figure 13. Closed-loop dynamic responses to ±20% throughput changes.

the distillation region where pure NPA is at the stable node. DIPE is selected as extraction solvent for NPA dehydration considering its favorable properties of density, heat of vaporization, and less toxicity. Since the solvent flow rate is an important degree-of-freedom that affects total energy consumption, it is optimized to obtain an optimal flowsheet. Also, results show that it is more appropriate to send V2 stream back into the extraction column. Compared to the two-column extractive distillation by An et al.,7 significant savings on total annual cost (over 62%) is realized via this hybrid process. Moreover, a novel control strategy is also devised on the basis of closed-loop and open-loop sensitivity tests, which can allow the operation at optimal design conditions. The core concept of this strategy is to hold the material balance lines in the azeotropic column by adjusting the solvent flow rate. Temperature on the stage sixth in the azeotropic column is controlled by manipulating the reboiler duty. The solvent flow rate is adjusted by a fixed ratio of solvent flow rate to reboiler duty in the azeotropic column. Another ratio control is applied to manipulate the reboiler duty in the water recovery column based on the total feed flow rate into this column. Results show that, with the aid of this control strategy, all of the product purities can still be maintained near their specifications despite any feed disturbances.

in this control strategy. Figure 13 displays the dynamic responses of this disturbance. With the aid of inner-loop ratio multiplier, QR1/FF, the temperature only deviates within 2 K and is quickly controlled back to its set point value. The solvent flow rate is also successfully adjusted. Both product purities are also maintained near their specifications although there are very small steady-state deviations in water purity. If necessary, a more complicated feedfoward action can be added into the QR2 loop of proposed control structure so that the ratio can be properly adjusted during “known” throughput changes.

5. CONCLUSIONS In this work, the potential for significant energy-saving via a hybrid extraction−distillation system is demonstrated with a conceptual illustration and a case study of n-propanol dehydration. This hybrid process can be viewed as a derivative from the heterogeneous azeotropic distillation process. In the HAD process, composition of the overall feed into the azeotropic column is taken to a distillation region where NPA is at the stable node by mixing organic reflux and preconcentrated feed. However, this task is accomplished in the hybrid process through extracting fresh feed with the solvent stream. The potential for significant energy-saving for the hybrid process comes from a huge decrease in vaporizing water and organic stream compared to the HAD process. As a feasible solvent for the hybrid process, it should form a new heterogeneous and minimum boiling azeotrope in the ternary system and should be able to extract NPA from the fresh feed. In addition, the liquid−liquid envelope needs to be large enough for the extract phase composition to be located in

■

ASSOCIATED CONTENT

S Supporting Information *

The Supporting Information is available free of charge on the ACS Publications website at DOI: 10.1021/acs.iecr.6b02210. M


Article


■

Azeotropic Distillation Systems. Ind. Eng. Chem. Res. 2014, 53, 1537−1552. (16) Gmehling, J.; Möllmann, C. Synthesis of Distillation Processes Using Thermodynamic Models and the Dortmund Data Bank. Ind. Eng. Chem. Res. 1998, 37, 3112−3123. (17) Seider, W. D.; Seader, J. D.; Lewin, D. R.; Widagdo, S. Product and Process Design Principles Synthesis, Analysis, and Evaluation; Wiley: Hoboken, NJ, 2010. (18) Luyben, W. L. Comparison of extractive distillation and pressure-swing distillation for acetone/chloroform separation. Comput. Chem. Eng. 2013, 50, 1−7. (19) Luyben, W. L.; Chien, I. L. Design and Control of Distillation Systems for Separating Azeotropes; Wiley: Hoboken, NJ, 2010. (20) Tyreus, B. D.; Luyben, W. L. Tuning PI Controllers for Integrator/Dead Time Processes. Ind. Eng. Chem. Res. 1992, 31, 2625−2628.

Details regarding model parameters, process optimization, and closed-loop sensitivity tests (PDF)

AUTHOR INFORMATION

Corresponding Author

*Tel.: +886-3-3366-3063. Fax: +886-2-2362-3040. E-mail: [email protected]. Notes

The authors declare no competing financial interest.

■

ACKNOWLEDGMENTS The research funding from the Ministry of Science and Technology of R.O.C. under Grant MOST 105-2218-E-002003 is greatly appreciated.

■

REFERENCES

(1) Mahdi, T.; Ahmad, A.; Nasef, M. M.; Ripin, A. State-of-the-Art Technologies for Separation of Azeotropic Mixtures. Sep. Purif. Rev. 2015, 44, 308−330. (2) Errico, M.; Sanchez-Ramirez, E.; Quiroz-Ramirez, J. J.; SegoviaHernandez, J. G.; Rong, B. G. Synthesis and design of new hybrid configurations for biobutanol purification. Comput. Chem. Eng. 2016, 84, 482−492. (3) Pereiro, A. B.; Araújo, J. M. M.; Esperanca, J. M. S. S.; Marrucho, I. M.; Rebelo, L. P. N. Ionic Liquids in Separations of Azeotropic Systems − A Review. J. Chem. Thermodyn. 2012, 46, 2−28. (4) Kürüm, S.; Fonyo, Z.; Kut, Ö . M. Design Strategy for Acetic Acid Recovery. Chem. Eng. Commun. 1995, 136, 161−176. (5) Lucia, A.; Amale, A.; Taylor, R. Energy Efficient Hybrid Separation Processes. Ind. Eng. Chem. Res. 2006, 45, 8319−8328. (6) Chen, Y. C.; Li, K. L.; Chen, C. L.; Chien, I. L. Design and Control of a Hybrid Extraction−Distillation System for the Separation of Pyridine and Water. Ind. Eng. Chem. Res. 2015, 54, 7715−7727. (7) An, Y.; Li, W.; Li, Y.; Huang, S.; Ma, J.; Shen, C.; Xu, C. Design/ optimization of energy-saving extractive distillation process by combining preconcentration column and extractive distillation column. Chem. Eng. Sci. 2015, 135, 166−178. (8) Pla-Franco, J.; Lladosa, E.; Loras, S.; Montón, J. B. Approach to the 1-propanol dehydration using an extractive distillation process with ethylene glycol. Chem. Eng. Process. 2015, 91, 121−129. (9) Pla-Franco, J.; Lladosa, E.; Loras, S.; Montón, J. B. Isobaric Vapor−Liquid−Liquid Equilibria for the Ternary Systems Ethanol + Water + Propyl acetate and 1-Propanol + Water + Propyl acetate. J. Chem. Eng. Data 2014, 59, 2054−2064. (10) Pla-Franco, J.; Lladosa, E.; Montón, J. B.; Loras, S. Evaluation of the 2-Methoxyethanol as Entrainer in Ethanol−Water and 1Propanol−Water Mixtures. J. Chem. Eng. Data 2013, 58, 3504−3512. (11) Lee, L. S.; Shen, H. C. Azeotropic Behavior of a Water + nPropanol + Cyclohexane Mixture Using Cyclohexane as an Entrainer for Separating the Water + n-Propanol Mixture at 760 mmHg. Ind. Eng. Chem. Res. 2003, 42, 5905−5914. (12) Pienaar, C.; Schwarz, C. E.; Knoetze, J. H.; Burger, A. J. Vapor− Liquid−Liquid Equilibria Measurements for the Dehydration of Ethanol, Isopropanol, and n-Propanol via Azeotropic Distillation Using DIPE and Isooctane as Entrainers. J. Chem. Eng. Data 2013, 58, 537−550. (13) Orchillés, A. V.; Miguel, P. J.; Vercher, E.; Martinez-Andreu, A. Isobaric Vapor−Liquid Equilibria for 1-Propanol + Water + 1-Ethyl-3methylimidazolium Trifluoromethanesulfonate at 100 kPa. J. Chem. Eng. Data 2008, 53, 2426−2431. (14) Chang, W. T.; Huang, C. T.; Cheng, S. H. Design and Control of a Complete Azeotropic Distillation System Incorporating Stripping Columns for Isopropyl Alcohol Dehydration. Ind. Eng. Chem. Res. 2012, 51, 2997−3006. (15) Wu, Y. C.; Lee, H. Y.; Huang, H. P.; Chien, I. L. Energy-Saving Dividing-Wall Column Design and Control for Heterogeneous N


Potential for Significant Energy-Saving via Hybrid Extraction

Recommend Documents