Synthesis, Design and Control of an Azeotropic Distillation System for

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Synthesis, Design and Control of an Azeotropic Distillation System for Methanol-Isopropyl Acetate Separation Kalp Mishra, and Nitin Kaistha Ind. Eng. Chem. Res., Just Accepted Manuscript • DOI: 10.1021/acs.iecr.8b01779 • Publication Date (Web): 23 Dec 2018 Downloaded from http://pubs.acs.org on December 24, 2018

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Industrial & Engineering Chemistry Research

Synthesis, Design and Control of an Azeotropic Distillation System for Methanol-Isopropyl Acetate Separation Kalp Mishra and Nitin Kaistha* Department of Chemical Engineering Indian Institute of Technology Kanpur, Kanpur 208016 (INDIA) Abstract The existence of a minimum boiling homogenous azeotrope precludes the separation of a methanol (MeOH) - isopropyl acetate (IPAc) binary mixture into its constituent pure components using simple distillation. In this work, alternative separation processes for recovering pure MeOH and pure IPAc are synthesized. Three promising flowsheet alternatives, FS3a-c, are further analyzed for energy efficient operation. The analyzed flowsheets use water as an entrainer for MeOH and consist of a liquid-liquid extractor (LLX) and three columns. The flowsheets differ in where the fresh feed is fed. The dominant design variables for the process are the entrainer (water) recirculation rate to the LLX and the number of extractive trays. A design steady state is obtained by adjusting these dominant variables for (near) minimum total reboiler duty. For the nominal equimolar binary fresh feed, FS3a is found to be significantly more energy efficient than the other two variants. A decentralized plantwide control system is synthesized for its recommended process design and shown to reject large throughput and fresh feed composition changes. It is also shown that the flowsheet recommendation can change for extreme fresh feed compositions. Thus for a MeOH lean feed, FS3b is clearly superior while for a MeOH rich feed, FS3c is superior. The recommended flowsheet, FS3a, is more energy efficient over a large fresh feed composition range.

Keywords:

Conceptual process design, azeotropic distillation, extractive distillation flowsheet synthesis, plantwide control

To whom all correspondence should be addressed. Email: [email protected]; Phone: +91-512-2597513; Fax: +91512-2590104. *

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Introduction Distillation remains the most preferred separation technique in the process industry, primarily due to its simplicity and ability to accomplish two clean separation tasks in a single unit, whereby the top and bottom products are nearly free of the heavy key impurity and light key impurity, respectively. The countercurrent vapor-liquid flash cascade, as realized in a tray/packed column with vapor flowing upwards and liquid flowing downwards, separates a binary ideal AB mixture into a nearly pure A (light component) distillate (top) product and a nearly pure B (heavy component) bottom product. Both high purity and recovery of each of the two components is thus achieved in a single column. The process system is very simple, usually consisting of a steam heated reboiler, a water cooled condenser and a tray section for countercurrent vapor liquid contact. No elaborate special arrangement is needed to make the two phases flow in opposite directions with the hot vapor naturally flowing upwards and the hot liquid naturally flowing downwards. Also vigorous vapor-liquid contact occurs due to boiling for good overall mass transfer coefficients. Indeed, distillation is a very clean separation technique with sloppy splits being infeasible 1,2. By a sloppy split, we mean the leakage of heavier than heavy key (lighter than light key) in the distillate (bottoms) at a given heavy key (light key) leakage cannot exceed a limit. For mixtures with highly non-ideal behaviour, phenomena such as azeotropy makes pure component recovery from a binary mixture infeasible using simple distillation. Even so, the addition of an appropriate third component (solvent or entrainer) significantly alters the vapor liquid equilibrium (VLE) to make pure component recovery feasible using an appropriate distillation sequence. As an example, we have the celebrated homogenous extractive and heterogenous azeotropic distillation sequences for breaking a binary azeotrope 3,4. Additionally, the distillation literature is replete with innovative sequences for separating a given mixture into its constituent pure components (see example references

5-9).

The

flowsheet structure depends on the specific vapor-liquid-equilibrium characteristics of the particular mixture to be separated. Over the years, the residue curve map has emerged and has been applied as a


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powerful visual tool enabling feasible distillation separation scheme synthesis to separate highly non-ideal mixtures 2,3,7-11. Even as distillation is a mature technology, there is renewed interest in the same as the manufacturing industry moves towards "green" processing to address sustainability concerns. In addition to highly selective reaction chemistry, "green" processing requires near 100% recovery of the valuable product as well as separation, recycle and reuse of raw materials such that the process output is only the finished product (near zero waste discharge). Given its ability to achieve two clean separation tasks in a single column, distillation remains the primary candidate for achieving high product recovery and purity as well as raw material recycle and reuse objectives to maximize material utilization efficiency. Thus, for example, in the semiconductor and pharmaceutical industries, where solvents are routinely used in various processing steps such as chip cleaning and active pharmaceutical ingredient (API) extraction, the key link in the transformation to "green" processing is the recovery and reuse of solvents from the generated waste mixtures. Since these mixtures often contain water, alcohols, esters and ketones, distillation-based separation is complicated by multiple azeotropes and consequent distillation boundaries. The synthesis of a feasible distillation sequence is then a major bottleneck in "green" processing. We thus have several articles on the synthesis, design and control of a distillation sequence for separating an azeotropic mixture into its constituent pure components appearing in recent years (see example references

12-16).

These reflect the renewed research interest in distillation technology for

separating non-ideal mixtures. In this work, we consider the separation of a methanol (MeOH) - isopropyl acetate (IPAc) binary mixture into its constituent pure components using distillation. Such a mixture is encountered in the pharmaceutical industry, as both MeOH and IPAc are active pharmaceutical ingredient extraction solvents 17.

The separation is complicated by the presence of a minimum boiling MeOH-IPAc binary azeotrope. The

use of water as an entrainer for breaking the azeotrope has been indicated in the literature 11,18. Even as



a conceptual flowsheet structure is presented, no further quantitative analyses on its design are presented. Given that water is the cheapest possible entrainer, it results in a very low cost for re-charging the process with fresh entrainer, say once every few months, to make-up for the slow accumulated entrainer loss in the product streams. For any other entrainer, the recharging cost can be two orders of magnitude higher. For this reason, we focus here only on water as the entrainer. The entire process design cycle consisting of (a) feasible flowsheet synthesis; (b) steady state design optimization and flowsheet screening; and (c) rigorous plantwide control for validating operability of the finally recommended flowsheet, is performed for water as the entrainer. Step (c) is important as nonlinear amplification of transients around material and energy recycle loops can compromise process operability over the envisaged disturbance space. Step (c) allows revision of the equipment size at the conceptual process design stage to ensure robust operability. The main contribution of the work is in presenting multiple flowsheet options and quantitatively evaluating three promising ones to make a considered recommendation on the best flowsheet. In the following, the residue curve map tool is applied to synthesize candidate feasible flowsheets for recovering pure MeOH and pure IPAc from a binary MeOH-IPAc mixture, using water as entrainer, if required. Using engineering common sense, a promising candidate, FS3, consisting of three columns and a liquid-liquid extractor (LLX) is screened for further analysis. FS3 exploits entrainer induced liquid-liquid phase split in the LLX to cross the distillation boundary for feasible recovery of pure MeOH and pure IPAc products. Three variants of FS3, that differ in the fresh feed location into the process, are evaluated. For the three variants, a steady state design is developed for an equimolar MeOH-IPAc binary fresh feed to minimize total reboiler duty (QTOT). In the variants, the water recirculation rate and the number of extractive trays are the dominant design variables. Steady state sensitivity analysis is used to obtain near optimum values of the dominant design variables for a near minimum QTOT design. Established heuristics are applied to obtain the other design variables. Of the three variants (FS3a-c), FS3a found to be


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significantly more energy efficient than the other two variants for the nominal feed and is therefore recommended as the preferred flowsheet. A decentralized plantwide control system with the throughput manipulator (TPM) at the fresh feed is then synthesized for FS3a. It is tested for large throughput and fresh feed composition changes using rigorous dynamic simulations. Finally, a brief analysis of the variation in energy efficiency with fresh feed composition is performed for the screened flowsheets. A summary of the main findings concludes the article.

Process Modeling Aspen Plus Version 8.8 is used for steady state and dynamic process modeling. Since the MeOHIPAc-water mixture is highly non-ideal with minimum boiling MeOH-IPAc homogenous and IPAc-water heterogenous azeotropes, the NRTL activity coefficient model is used for the liquid phase. The vapor phase is treated as ideal as the process is operated at low pressures. Experimental VLE and liquid-liquid equilibrium (LLE) data of the binary pairs is regressed using Aspen's built-in utility to fit the activity coefficient model parameters. The exact form of the NRTL model equation and the fitted model parameters are reported in Table 1. The extended Antoine parameters for modeling pure component vapor pressures are also reported in the Table. To assess the goodness of the equilibrium predictions from the model, Figure 1 compares the model predicted vapor phase mol fraction (y) in equilibrium with liquid phase mol fraction (x) at 1 bar with overlaid experimental data points for the two binary pairs of MeOH-IPAc 19 and MeOH-water 20. For the IPAc-water binary pair, experimental liquid-liquid equilibrium data 21 is fitted. The good agreement in the predicted and the experimental data points confirms the suitability of the chosen thermodynamic model.

Flowsheet Synthesis



The residue curve map (RCM) tool is applied to synthesize feasible flowsheets for separating the binary MeOH-IPAc mixture into its constituent pure components using water as an entrainer. The RCM at 1 bar along with the distillation boundaries (if any) and the liquid-liquid phase envelope is shown in Figure 2. The RCM stationary points are summarized in Table 2. The MeOH-IPAc azeotrope is an unstable node (potential distillate), the IPAc-water heterogenous azeotrope and the pure MeOH vertex are saddles (unreachable by simple distillation) and the pure water and IPAc vertices are stable nodes (potential bottoms product). These stationary points divide the composition triangle into Region I and Region II, separated by a distillation boundary . To synthesize feasible flowsheets, we first consider the possibility of using pressure swing to cross the MeOH-IPAc azeotrope. If the azeotrope composition is sensitive to pressure, then pressure swing distillatiion (PSD) may be used to obtain both pure MeOH and pure IPAc without having to add a third entrainer component. We found that as the pressure is increased, the azeotrope becomes richer in MeOH. For an equimolar fresh feed, the PSD flowsheet, FS1, is shown in Figure 3a. The high pressure (HP) column processes the fresh feed to give a MeOH rich HP azeotrope distillate with pure IPAc leaving down the bottoms. The HP azeotrope is then distilled in the low pressure (LP) column to obtain pure MeOH down the bottoms and the LP azeotrope up the top, which is recycled back to the HP column. From the VLE y vs x curve in Figure 1, notice that the MeOH rich lobe to the right of the azeotrope is very close to the 45° line implying a very difficult separation to achieve near pure MeOH. This may also be verified from the very close boiling points of the MeOH-IPAc azeotrope and MeOH in Table 2. Thus FS1 is expected to be economically unviable due to very high energy cost. The next option is to consider the use of water as an entrainer. The liquid-liquid phase split is then a means of crossing the distillation boundary and obtaining potential feeds in Region I and Region II. To synthesize feasible flowsheet alternatives, consider a binary MeOH-IPAc feed. Its composition would be on the MeOH-IPAc edge of the composition triangle. By adding enough water (entrainer), the overall


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composition can be brought sufficiently inside the liquid-liquid envelope such that decanting the same results in an aqueous layer composition in Region II and an organic layer composition in Region I. Alternatively a liquid-liquid extractor (LLX) may be used; water being a good extractive agent for MeOH. The LLX top organic draw is then nearly MeOH free and may be stripped to obtain pure IPAc (stable node) as the bottom product and a vapor distillate, which is condensed and recycled. Its composition will be near the Region I - Region II distillation boundary. Some thought is required in devising the processing scheme for the LLX aqueous draw that carries the MeOH and has a composition in Region II. If we distill it to obtain pure water (stable node) as the bottoms, then for an aqueous draw composition close to the IPAc-water edge, an IPAc free distillate with a composition close to the MeOH vertex on the MeOH-IPAc edge in Region II can be obtained. This (near) binary distillate may be further distilled to obtain the MeOH-IPAc azeotrope as the distillate, which is recycled, and nearly pure MeOH as bottom product. The separation however will be very expensive as the boiling points of MeOH and MeOH-IPAc azeotrope are very close. Thus, this proposed separation scheme, FS2, as shown in Figure 3b, although theoretically feasible, is again expected to be economically unviable. A more attractive option is to take the LLX aqueous draw and strip it such that the bottoms composition is noticeably away from the pure water vertex on the MeOH-water edge. This is allowed as residue curves closely mimic the composition profile of a column at total reflux 2. Since the Region II residue curves approach the water vertex along the MeOH-water edge (see Figure 2), it should be possible to specify the bottoms composition to lie close to the MeOH-water edge, away from the water vertex. We thus obtain an IPAc free MeOH-water binary mixture as the stripper bottoms, which can be easily processed in a simple column to obtain pure MeOH as the distillate and pure water as the bottoms. Note that the MeOH-water separation is relatively much easier with a large difference of ~35 °C in the component boiling points. The water bottoms is recycled back to the LLX post cooling. Figure 3c shows



the proposed flowsheet, FS3a, with a multistage LLX. Note that a simple decanter can be considered as a single stage extractor. Two other feasible variants with the same conceptual structure as for FS3a are obtained by altering the location of the fresh feed into the process. These are shown in Figure 3d and Figure 3e and labeled, FS3b and F3c, respectively, for convenient reference. In FS3b, the fresh feed is fed to the IPAc stripper. In FS3c, the fresh feed is fed into the column that prevents IPAc leakage down the column for an IPAc free MeOH-water bottoms. A careful examination of the RCM in Figure 2 reveals that the Region II residue curves are very similar to classic homogenous extractive distillation. The curves emanate from the minimum boiling azeotrope and move towards the entrainer (water) stable node. Figure 4 contrasts the RCM for the current system (without the liquid-liquid envelope) and classic homogenous extractive distillation for an ethanolwater-ethylene glycol mixture, where ethylene glycol is the entrainer. In Region II of the studied system, there are two sets of curves. One set reaches the water (entrainer) vertex (stable node) after approaching and then turning away from the MeOH vertex saddle. The other set reaches the water (entrainer) vertex (stable node) after approaching and then turning away from the IPAc-water azeotrope saddle. This is very similar to the extractive distillation RCM where the two sets of residue curves move towards their respective saddles and then turn away to reach the entrainer stable node. To further appreciate the similarity with classic homogenous extractive distillation, consider a standalone extractive unit consisting of the stripper, possibly an extractive section above the stripper and an LLX as shown in Figure 5. Water is known to be a good extractive agent for MeOH. The main idea is to provide sufficient water (entrainer) flow into the extractive unit and let the water (entrainer) soak in the MeOH so that negligible MeOH leaves in the top organic draw. Almost all of the feed MeOH then leaves down the bottoms along with the water entrainer. The extractive unit is the key operation in that it prevents MeOH leakage up the top and IPAc leakage down the bottoms. The top composition will be close


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to the IPAc-water base on the organic liquid-liquid equilibrium curve. Consider a simple decanter (instead of LLX) with entrainer feed near the top column stage, as in classic extractive distillation. If the decanter organic and aqueous layers are nearly MeOH free, then the vapor leaving the entrainer feed tray must be MeOH free with a composition along the IPAc-water edge. The few trays on top of the entrainer feed will cause the top vapor composition to reach the IPAc-water heterogenous azeotrope, which is a saddle in the RCM. Thus the extractive column top vapor is close to a saddle, which is analogous to the top vapor product in classic homogenous extractive distillation being close to the "not soaked" component saddle vertex. The separation has thus been aptly termed as "heterogenous extractive distillation" in the literature 11. The only difference in our suggested flowsheet is that instead of a decanter, we are using a multistage LLX and the entrainer is fed to the LLX and not the column, for better extraction of MeOH. The three FS3 variants differ from each other in terms of the feed and its location into the extractive unit (see Figure 5). In FS3a, the process fresh feed is fed directly into the LLX. In FS3c, the process fresh feed is fed into the column, so that we have an extractive section in the column. In FS3b, instead of the process fresh feed, the IPAc stripper top product is fed to the LLX. It is also possible to feed it into the column. This variant was found to result in very similar total reboiler duty for when the extractive unit feed is sent to the LLX. It is therefore not considered further and our focus shall be on FSa, FS3b and FS3c.

Steady State Flowsheet Analyses Of the candidate flowsheets synthesized, FS1 and FS2 are not analyzed any further as these require the separation of a feed stream on the MeOH rich side of the azeotrope, into pure MeOH and the MeOH-IPAc azeotrope. Since the boiling points of MeOH and the MeOH-IPAc azeotrope at 1 bar are only ~0.2 °C apart, the separation is economically unviable. This leaves the three FS3 variants as candidates meriting further steady state analysis to screen the best one. Here, instead of applying rigorous economic cost optimization, we perform simple steady state analyses to minimize the total reboiler duty (QTOT),



which correlates well with the total annualized cost (TAC). Of the three variants, the one with lowest energy consumption should be chosen. We note that each of the three variants has an extractive unit stripper that prevents IPAc from leaking down the bottoms, a MeOH recovery column, an IPAc recovery stripper and an LLX. These are referred to as respectively, the extractive unit stripper (C1), the MeOH column (C2), the IPAc column (C3) and the LLX, across all the flowsheets for easy and consistent reference. The three variants are to be designed to process 100 kmol/h of binary MeOH-IPAc fresh feed available at 37 °C. Its nominal composition is 50:50::MeOH:IPAc (mol %). The required MeOH and IPAc product purities are 99.5 mol% each. In the following, we walk the reader through engineering common sense driven steady state analyses for designing a near optimum energy efficient flowsheet variants, FS3a-c. The following sequential steps are followed for designing the flowsheet variants: 1. First, the number of trays in columns that have a "local" and not plantwide effect on the material/energy balances are fixed from a steady state analysis on standalone columns. The MeOH column, IPAc column and the extractive unit stripper trays are thus obtained. 2. Next, a steady state degree of freedom (dof) analysis is performed and specifications for production rate, product quality and LLX feed temperatures are fixed. Of the remaining three dofs, the IPAc leakage down the extractive unit stripper and the entrainer purity down the MeOH column are fixed to reasonable values using simple common sense analyses. 3. Finally, plantwide simulations are performed to fix the dominant design variables with a plantwide effect on material and energy balances to reasonable values that give near minimum total reboiler duty. These dominant design variables are the LLX entrainer rate and the number of extractive trays. Number of Trays in Columns Before performing a quantitative steady state analysis, we need to fix the number of trays in the various tray sections. The simplest is the MeOH column that performs a near ideal MeOH-water split. The


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shortcut distillation method (DSTW in Aspen Plus) is then appropriate and applied to obtain the minimum trays as ~12. The total trays is set to 40, which is about thrice the minimum trays. The MeOH column in all the three variants, FS3a-c, thus has 40 trays. Further, since the MeOH-water mixture down the extractive unit bottoms is likely to be water rich for ensuring entrainment of all the MeOH in the extractive unit feed, the MeOH column feed should be lower down the column. Here, Tray 30 is chosen as the feed tray giving 10 stripping trays. All steady state analyses are performed for this tray location. Next we consider fixing the number of trays in the IPAc stripper. To do so, a standalone stripper is simulated for a given stripper feed. The column design spec vary feature is applied that adjusts the stripper reboiler duty (vary) for an IPAc purity of 99.5 mol% (spec). The number of trays is decreased from a large number (say 20) until the stripper reboiler duty exhibits is sharp increase. This corresponds to approaching the minimum tray limit. The number of trays is set to a value sufficiently above this "sharp increase" limit, as shown in Figure 6a. In this way, the number of trays in the IPAc stripper is obtained as 7 and is applied to the three variants (FS3a-c). In the extractive unit, we have the number of LLX trays (NLLX) and number of stripping trays as well as the number of extractive trays (NE) in FS3c, which are the trays above the extractive unit stripper feed. An analysis on a standalone extractive unit is used to fix the number of stripping trays. For given extractive unit feeds from a converged plantwide simulation, a flowsheet design spec vary is applied that varies the entrainer rate (vary) to hold the LLX organic discharge MeOH mol fraction (spec) at 0.005. For a fixed number of LLX trays (5) and number of extractive trays (7) in the extractive column, the number of stripping trays (trays below extractive column feed) is reduced from a large number (say 20) till the extractive unit reboiler duty shows a sharp increase. This is illustrated in Figure 6b. The number of stripping trays in the extractive unit is then fixed at a value of 8, which is above this "sharp increase" limit. All the three flowsheets use 8 stripping trays in the extractive unit.



We are now left with fixing NLLX in all flowsheet variants and NE in FS3c. Their function is to soak in the MeOH so that the extractor unit top organic discharge is nearly MeOH free. NLLX and NE significantly affect solvent rate required for a given extractive unit organic draw MeOH impurity, and hence the plantwide balances and QTOT. It is then not possible to obtain reasonable values for the same via a standalone unit operation analysis. Plantwide steady state analyses must performed to fix them as well as any dominant steady state specifications to reasonable values for an energy efficient process.

Steady State Degrees of Freedom and Specifications All columns are assumed to operate at 1 bar as it results in reasonable top vapor temperatures allowing a water cooled condenser. For given distillation column pressures and specified trays in the various column sections/LLX, the three flowsheet variants have a steady state operating degree of freedom of 8; one for the fresh feed, two for the MeOH column, one for the IPAc stripper, one for the extractive unit stripper, one each for the entrainer and LLX feed coolers, and one for the water (entrainer) rate into the LLX. Of these 8 dofs, the fresh feed processing rate (throughput), IPAc product purity and MeOH product purity specifications take away three. Further, the LLX feed and entrainer feed temperatures are constrained by the coolant temperature. For cooling water at 27 °C and a 10 °C approach temperature in the coolers, these temperatures are fixed at 37 °C. This leaves 3 steady state dofs. We use the bottom IPAc component mol flow of the extractive unit stripper (fB1IPAc), the methanol column bottom MeOH mol fraction (xB2MeOH) and the LLX entrainer (water) flow rate (S) as the specifications for these dofs. In addition to these three steady state specifications, we also need to fix a reasonable value for NLLX in all three variants and additionally NE in FS3c. Exploratory steady state simulations were performed to obtain reasonable values for the three remaining specifiations. In these simulations, NLLX was fixed at 10 and NE was fixed at 25 for FS3c. These are noticeably larger than the final recommended values. On the IPAc stripper, the design spec vary


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feature is used to hold the bottoms IPAc purity at 99.5 mol% by adjusting the bottoms rate. The purpose of the extractive unit stripper is to prevent IPAc leakage down the bottom. Accordingly, the design spec vary feature is used to hold the IPAc component flow rate down the bottoms (fB1IPA) at a small value (say 0.1 kmol/h) by adjusting the bottoms rate. On the MeOH column, design spec vary is used to hold the bottoms MeOH mol fraction (xB2MeOH) at a small value (say 0.1 mol%) by adjusting the reflux ratio. Also, design spec vary is used to hold the distillate MeOH purity at 99.5 mol% by varying the distillate rate. The flowsheet specifications are noted in Table 3 for clarity. We first consider the tradeoff with respect to the extractive unit stripper bottom IPAc component flow rate, fB1IPAc. All of the IPAc that leaks down the extractive unit ends up in the downstream MeOH product stream (D2). The tight MeOH product purity constraint of 99.5 mol% implies a total impurity (IPAc + water) of 0.5 mol%. For a fresh feed MeOH component rate of 50 kmol/h, the MeOH product rate is expected to be ~50 kmol/h. The maximum total impurity rate in the MeOH product stream then is ~0.25 kmol/h. If fB1IPAc approaches ~0.25 kmol/h from below, then the water leakage up the top of the MeOH column (C2) must approach trace amounts requiring high reflux and hence reboiler duty. On the other hand, if fB1IPAc approaches trace amounts, the extractive unit stripper reboiler duty will shoot up. Figure 7 plots the QTOT vs fB1IPAc variation for FS3a. QTOT increases up as fB1IPAc approaches its two extreme limits and is flat away from the extreme limits. Similar flat behavior is also observed for FS3b and FS3c. Thus an fB1IPAc value in this flat region is nearly optimal. We fix fB1IPAc = 0.15 kmol/h which is in the flat region for the three flowsheet variants. Next, we consider setting an appropriate value for the entrainer impurity mol fraction, xB2MeOH. Since B2 is the entrainer stream, large impurity levels in the stream would compromise the efficacy of the entrainer and the entrainer circulation rate will then shoot up to achieve the desired separation. Figure 7 plots the variation in the solvent rate (S) required to hold the extractor unit top draw MeOH impurity at 1 mol% as xB2MeOH is varied. The graph is shown for a standalone extractor with NLLX = 8. From the Figure, an



entrainer MeOH impurity level of 0.1 mol% seems appropriate as it gives a solvent rate close to the ideal case of a pure entrainer.

Dominant Design Variable Specifications The only remaining design variables to be fixed are NLLX, NE (for FS3c) and S. NLLX and NE improve the extraction capacity. These have a plantwide effect on the material and energy balances since increased extraction capacity will require lower entrainer circulation rate for a nearly MeOH free extractor unit top organic draw. Plantwide simulations are performed to obtain their values for near minimum QTOT. For a given extraction capacity, there is a minimum entrainer rate below which the separation is infeasible. As the entrainer rate is increased above this minimum, the extractive unit reboiler duty is expected to decrease. On the other hand, since water drops down the unit with the MeOH, the MeOH column feed rate is higher so that its duty increases. One therefore expects a minimum in the QTOT vs S curve. This minimum is clearly evident ion Figure 9 that plots the curves for the three flowsheet variants as NLLX is increased from a simple decanter (NLLX = 1) to multiple LLX trays (NLLX > 1). For FS3c, NE is kept large at 25 trays. From the first three subplots, NLLX values of 8, 4 and 4 appear reasonable for FS3a, FS3b and FS3c, respectively, since on increasing NLLX, the further decrease in minimum QTOT becomes too small (diminishing returns). To fix NE for FS3c, the fourth subplot (Figure 9d) plots the QTOT vs S curve with NLLX = 4 and NE varied from 1 to 5. A value of 5 appears to be reasonable based on the diminishing returns principle. For clarity, the recommended number of trays in each of the column sections is reported in Table 4 for the three flowsheet variants. Using the above analysis, the (near) best possible design for the three flowsheet variants corresponding to minimum QTOT are obtained. The near minimum column reboiler duties, QTOT, and solvent rates, S, for these near best designs are also noted in Table 4. Clearly, FS3a is the most energy efficient for the nominal equimolar MeOH-IPAc binary fresh feed. FS3a is therefore the finally


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recommended flowsheet. Its salient design and operating conditions are shown in Figure 10. The streamwise flow data is provided in Supporting Information. The interested reader may use it to verify the convergence of the plant material balances. For completeness, the salient material balance lines for the best designs of the three flowsheet variants are shown in Figure 11. Notice that in each case, the organic draw from the extractive unit is nearly MeOH free. Also the bottoms from the extractive unit is nearly IPAc free. Also note that the liquid-liquid phase envelope is at the LLX temperature of 37 °C. Utilizing the material / energy flows at the recommended steady state, the other equipment size parameters are determined. All column diameters are obtained using the Aspen Plus built-in tray sizing utility based on Fair's method 22. At the nominal steady state, the column maximum vapor superficial velocity is set at 50% of the flooding velocity. This requires bubble cap trays as sieve trays are inappropriate due to weeping limitations at low vapor rates. With such overdesign, each column can withstand, at most, a ~100% increase in boil-up (vapor rate) before the onset of entrainment flooding. Condenser heat transfer areas are obtained using an overall heat transfer coefficient of 500 W m-2 °C-1 23 and a ΔTLMTD of 15 °C. Similarly, reboiler heat transfer areas are obtained using an overall heat transfer coefficient of 1200 W m-2 °C-1 23 and a ΔT of 30 °C. To obtain liquid-liquid heat exchanger area, an overall heat transfer coefficient of 300 W m-2 °C-1 23 is used. To provide additional heat removal capacity, all heat transfer areas are sized for a heat duty that is 50% higher than the heat duty at the recommended design steady state (heat transfer area overdesign). The LLX is diameter is calculated using a total liquid throughput of 35 m3 liquid h-1 ft-2 with a nominal 4 ft length per LLX stage 24. Table 5 summarizes the size and utility consumption of all the major equipment in the flowsheet. The distillation columns use standard 2 ft spacing between trays with 20% extra height for the bottom sump. For completeness, capital cost and operating cost estimates are also provided in the Table. These estimates have been obtained from the cost correlations and price data detailed in Supporting Information. The total annualized cost of the



process using a typical pay-back period of 3 years is $0.7×106 yr-1. From the engineering point of view, the plant consumes 0.747 kg low pressure steam per kg feed processed.

Plantwide Control A conventional decentralized plantwide control system is now synthesized. The control system must be designed to process a given feed, a common scenario in integrated plant complexes. Accordingly the fresh feed rate (F) is the throughput manipulator (TPM). As the fresh feed rate changes, the water (entrainer) recirculation rate around the plant must change in tandem. Accordingly, the solvent rate, S, is held in ratio with F. The remainder of the control system is straightforward. The mixed vapor condenser duty is manipulated to hold the condenser pressure at 1 bar. This fixes the C1 and C3 pressure to 1 bar. On the LLX, Aspen Plus has an internal stagewise aqueous and organic material balance solver that fixes the LLX organic and aqueous draw rates. In practice, the interface level is maintained by adjusting one of the draw rates (organic or aqueous), depending on which is the continuous phase. For example, if water is the continuous phase, then the organic-aqueous interface will exist near the top of the LLX. This interface level shall be maintained by adjusting the top organic draw rate. Note that water circulates around the plant in a closed loop with negligibly small losses. The plant water inventory is reflected in the MeOH column sump level. In view of the negligible water loss, the sump level is not controlled. The assumption is that water loss accumulated over sustained process operation will result in a gradual decrease in the level. This loss is conveniently made up by recharging the sump with appropriate amount of water, say once every few weeks, as the level reduces to a sufficiently low value. On the other columns, the sump level is maintained by manipulating the bottoms rate. Additionally, the reflux drum level on the MeOH column is maintained by manipulating the distillate rate. Similarly the C1 - C3 condensate drum level is controlled by manipulating the drum outlet rate. A sensitive tray temperature, corresponding to a large change in tray composition, is maintained by manipulating the reboiler duty on the columns.


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Specifically Tray 5, Tray 14 and Tray 5 temperatures (top-down numbering) are regulated for C1, C2 and C3, respectively. Note that the controlled tray temperature in C2 is in the rectifying section to achieve tight distillate MeOH product purity control. The complete control structure is shown in Figure 12. The steady state simulation is exported to dynamics post sizing of the reboilers and C2 condenser for a 5-10 min hold up at 50% level. A flow driven simulation is performed since Aspen dynamics does not currently simulate a pressure driven LLX . A 2 minute lag is applied to all controller outputs to an energy stream to account for heat transfer dynamics. Thus, all column pressure and temperature controller outputs are lagged by 2 mins. Further, a 30 s lag is applied to all temperature measurements to account for sensor dynamics. All process flows (material or energy) and levels are spanned from 0 to twice their nominal design value. All flow controllers are PI and use a gain of 0.5 and 0.5 mins for a fast and snappy servo response. Similarly, all column pressure controllers are PI. The pressure controllers are tuned to the Zeigler Nichols (ZN) settings using the relay feedback autotuner. All level controllers are P only and use a gain of 2 (%/%). The water recycle cooler temperature controller is tuned to its PI ZN autotuner settings. On the columns, the tray temperature control loops are tuned to the Tyreus-Luyben 25 PI settings from the autotuner with the column pressure and level loops on automatic. The salient controller tuning parameters thus obtained and used to perform rigorous closed loop dynamic simulations are provided in Supporting Information. The closed loop performance of the plantwide control system is evaluated for a ±20% step change in throughput and a ±10 mol% step change in the fresh feed MeOH composition. The throughput step change transient response of salient process variables (PVs) is shown in Figure 13. The response shows smooth transients in the process flows with the plantwide response completing in about 10 h. Also tight regulation of the MeOH and IPAc product purities is achieved with negligible (< 0.1 mol%) deviation at the final steady state.



The composition change plantwide transient response is shown in Figure 14. Note that this is a more severe disturbance in that the load in aqueous and organic processing loops must appropriately change in response to the feed composition change. As seen from Figure 13, a smooth plantwide response is obtained with response completing in about 10 h. Also, acceptably tight control of the MeOH and IPAc product purities is achieved during the transients with a small (

Synthesis, Design and Control of an Azeotropic Distillation System for

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