Continuous diisobutylene manufacturing: Conceptual process design and plantwide control

Accepted Manuscript Title: Continuous Diisobutylene Manufacturing: Conceptual Process Design and Plantwide Control Author: Shudodhan Singh Thakur Ojasvi Vivek Kumar Nitin Kaistha PII: DOI: Reference:

S0098-1354(16)30341-6 http://dx.doi.org/doi:10.1016/j.compchemeng.2016.11.007 CACE 5595

To appear in:

Computers and Chemical Engineering

Received date: Revised date: Accepted date:

25-2-2016 2-11-2016 10-11-2016

Please cite this article as: Thakur, Shudodhan Singh., Ojasvi, ., Kumar, Vivek., & Kaistha, Nitin., Continuous Diisobutylene Manufacturing: Conceptual Process Design and Plantwide Control.Computers and Chemical Engineering http://dx.doi.org/10.1016/j.compchemeng.2016.11.007 This is a PDF file of an unedited manuscript that has been accepted for publication. As a service to our customers we are providing this early version of the manuscript. The manuscript will undergo copyediting, typesetting, and review of the resulting proof before it is published in its final form. Please note that during the production process errors may be discovered which could affect the content, and all legal disclaimers that apply to the journal pertain.

Second revision submitted to Computers and Chemical Engineering

Continuous Diisobutylene Manufacturing: Conceptual Process Design and Plantwide Control Shudodhan Singh Thakur, Ojasvi, Vivek Kumar and Nitin Kaistha Department of Chemical Engineering, Indian Institute of Technology Kanpur, Kanpur 208016 (INDIA) *

To whom correspondence should be addressed. Email: [email protected]; Phone: +91-5122597513; Fax: +91-512-2590104.

Highlights      

Simultaneous design encompassing flowsheet synthesis, optimization and plantwide control demonstrated for diisobutylene (DIB) manufacturing. Processes use tert-butyl alcohol (TBA) moderator, generated in-situ and recycled-toextinction, for DIB yield enhancement. Two novel flowsheets, FS1 (brand new) and FS2, using a decanter for water recovery and recycle developed. FS1 and FS2 shown to achieve both high reactor conversion and DIB yield, unlike extant literature flowsheet, FS0. FS1 found to be superior to FS2 in terms of capital cost, energy efficiency and fresh water consumption. Demonstration of innovative control loops for yield and conversion regulation for “green” process operation.

Abstract The complete process design cycle encompassing flowsheet synthesis, design and controllability evaluation is studied for continuous DIB manufacturing. The residue curve map tool is applied to synthesize two new flowsheets, FS1 and FS2, exploiting pressure swing distillation. A unique feature of these is the use of a decanter for recovery and recycle of water, which allows maintaining the reactor tert-butyl alcohol content at the desired level for suppressing the side reaction. Unlike the literature flowsheet (FS0), this innovation makes it possible to achieve both high conversion and high yield for an economically superior design. Between FS1 and FS2, FS1 is found to

be superior with significantly lower capital and energy costs as well as lower fresh water consumption. A “smart” control strategy for holding the single-pass reactor conversion and the overall process yield towards economic process operation is also developed. Rigorous dynamic simulations demonstrate the controllability of FS1 and FS2.

Keywords: Simultaneous process design, flowsheet synthesis, conceptual process design, sustainable process design.

Introduction Additives are routinely blended to gasoline cuts in refining operations to guarantee the oxygenate content and octane number. Until recently, methyl tertiary butyl ether (MTBE; octane # 110) was a preferred additive. Its use as an additive has however been banned in several US states due to groundwater contamination owing to its high water solubility1. The refining sector has thus sought economic technologies for more benign alternatives to MTBE. Since MTBE was produced by the reaction of isobutylene (IB) in the C4 feedstock with methanol, the preference is for technologies based on alternative chemistry on the available C4 and methanol feedstocks. In the above context, there is tremendous interest in process technologies for diisobutylene (DIB) manufacture. DIB is a dimer of IB, IB being available in the C4 feedstock. The DIB can be hydrogenated to isooctane (octane # 100) for use as a gasoline additive. The DIB may also be etherified with methanol to produce 2-methoxy-2,4,4-trimethyl pentane (C8ME), a benign additive (octane # 99) with significantly lower water solubility (0.014 wt%) compared to MTBE (4.3 wt%)2. DIB is produced via the dimerization of IB over an acid catalyst. To prevent further oligomerization to triisobutylene (TIB) and higher oligomers, the reaction must be performed in the presence of polar alkyl alcohols3. The preference is for tert butyl alcohol (TBA), produced in-situ via the reversible addition of water to IB. The presence of water and polar TBA in the reactor effluent however complicates its separation due to multiple azeotropes and liquid-liquid phase split.

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Even as there is tremendous industrial interest in DIB manufacturing, the open literature on the same is very sparse. These literature reports either patent catalysts and operating conditions for selective IB dimerization or fit kinetic models on experimental reaction data or patent feasible separation cum recycle schemes to recover nearly pure DIB from the reactor effluent. The various acid catalysts patented include zeolites, amberlyst and phosphoric acid (see e.g.4-6). The reaction kinetics reports mostly fit kinetic parameters on data obtained from experiments that do not use TBA or any other moderator for suppressing the further oligomerization of DIB (see e.g.7-8). These are then of little practical relevance due to significantly poor selectivity to DIB. The most relevant reaction kinetics report is the work by Honkela et al.3 that fits Langmuir-Hinshelwood kinetics to IB dimerization / trimerization in the presence of TBA moderator for suppressing IB oligomerization. The only notable work on reactor effluent DIB recovery is the patent by Wang et al.9 that exploits pressure swing distillation to cross distillation boundaries. Their reported process operating conditions are however for a low-conversion low-yield reactor design. Given that conversion and yield significantly affect the plant material balance (and hence overall economics), the feasibility of the flowsheet for a more practical / economical high-conversion high-yield process is suspect. The above brief literature survey brings out the need for a comprehensive process design study of a complete high-conversion high-yield DIB manufacturing process. Such a study should ideally consist of the following three steps, possibly with iterations between them: 1. Feasible flowsheet synthesis and screening (conceptual design) 2. Steady state economic design and 3. Controllability evaluation The benefits of the comprehensive study includes addressing flowsheet structure feasibility issues in light of phase equilibrium constraints (e.g. distillation boundaries), the underlying design tradeoffs in a feasible flowsheet alternative as well as the economic, controllability and sustainability pros and cons of the screened flowsheet alternatives for an informed decision on the best alternative.

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It is pertinent to highlight that controllability evaluation (Step 3 above) of a proposed process design is now considered an essential step in the complete process design cycle as in the current “go green” times10-12, the process industry is witnessing increasing levels of mass and energy integration (recycle) for zero waste discharge and enhanced energy efficiency. This can result in potential controllability issues13-15 with highly non-linear and multivariable interaction between the integrated units. There also exists an inherent conflict between optimum steady state design and dynamic controllability16. Further, the control system configuration can significantly affect the economic / “green” benefit of a given plant design17-18. Controllability evaluation is then essential so that design modifications for robust plant operability are undertaken at the conceptual design stage. The importance of systematically addressing the design-controllability interaction has long been recognized in the literature. It is however only in the past decade that methodologies that optimize a single criterion combining economic design and control objectives have been developed and demonstrated on relatively simple process systems. Since the optimization attempts to obtain the “best” design for a combined design + control objective function, the approach is aptly termed as simultaneous process design. The reader is referred to recent articles to appreciate the state-ofthe-art in the area of simultaneous process design19-23. The simultaneous process design paradigm is intellectually appealing in that it can provide a “strictly optimal” solution to the design+control problem. There is however the larger issue of (over)simplified process models (plug flow reactors, equilibrium trays, ideal mixing etc), very large (10-50%) parametric uncertainty (kinetics, vapour liquid equilibrium, enthalpies etc) in these simplified models coupled with substantial volatility in the raw-material/utility/equipment costs. In view of the large model uncertainties and price volatility, such guarantees on strict optimality of a recommended design make little practical sense. Furthermore, the approach is mathematically complex, requiring considerable time and effort in its proper formulation, and the computational burden is prohibitively expensive, even for small size plants. In non-academic industrial settings, the simple three-step sequential approach that can quickly provide a reasonable near optimum 3

controllable design thus remains a preferred choice. This work is an application of the sequential approach to the synthesis, design and control of a DIB manufacturing plant, a problem of much industrial interest. Its emphasis is on walking a process design engineer through the thinking goes behind making reasonable design decisions using existing process simulation tools. In the following, the residue curve map (RCM)24 tool is applied to synthesize two new feasible flowsheet alternatives, FS1 and FS2, for DIB manufacturing via IB dimerization with TBA moderator. We also briefly describe the basic flowsheet, FS0, reported in a process patent9. All flowsheets (FS0-FS2) employ pressure swing distillation to recover nearly pure DIB and recycle DIB free TBA. FS1 and FS2 employ an innovative decanter based water recovery and recycle scheme which allows achievement of both high-conversion and high-yield, which represents a significant improvement over the extant literature flowsheet, FS0. An economic design for FS1 and FS2 is then developed and compared in terms of its energy and material efficiency. We then synthesize a decentralized plantwide control system for the two new flowsheets with innovative loops for maintaining single pass reactor conversion and process yield towards economic operation. Their closed loop performance is evaluated for large throughput and C4 feed composition changes. A summary of the main findings concludes the work.

Conceptual Process Design Reaction Chemistry and Kinetics DIB is produced via the irreversible, exothermic dimerization of IB over acid catalyst as Reaction 1:

IB + IB → DIB

The DIB further oligomerizes with IB to form the trimer, triisobutylene (TIB), as Reaction 2:

DIB + IB → TIB

In the absence of an appropriate moderator, significant further irreversible oligomerization of TIB occurs representing expensive yield loss. To mitigate the same, the reaction is carried out in the presence of an alkyl alcohol, which competitively binds to the catalyst active sites. With the right 4

moderator concentration, further oligomerization is sufficiently suppressed with only DIB and small amounts of TIB being formed for economically viable process yields. Here, tert butyl alcohol (TBA) is used as the moderator as it can be generated in-situ via the reversible addition of water to IB as Reaction 3:

IB + H2O ↔ TBA

Honkela et al.3 have performed a comprehensive experimental study on the above reaction system, proposing a reaction mechanism and fitting reaction kinetic parameters to the experimental data. Here, we use their reaction kinetic model for a realistic process design study. The kinetic model details and parameters are provided in Table 1.

Flowsheet Input-Output Structure Based purely on the reaction chemistry, an overall process input-output (IO) structure as in Figure 1 is envisaged. The process flowsheet takes in the C4 feedstock (composition in Table 2) with the generated DIB and any minor amounts of TIB leaving as the main product stream, and the inert C4s in the feedstock leaving as an off-gas stream. Since all C4s are close boiling, separation and recycle of any unconverted IB from the other C4s is assumed to be unviable. Any unconverted IB then leaves in the C4 off-gas. The reversible TBA formation reaction allows the possibility of in-situ TBA generation. To understand the same, let us specify a non-zero positive value for the water component flow rate to the reactor at a given fresh IB feed rate. TBA would then form by the reversible TBA formation reaction above. The TBA is not allowed to leave the process and circulates back to the reactor. This causes the reactor feed TBA component flow rate (and hence concentration) to increase, in turn causing the TBA dissociation rate by the reverse reaction to increase. The steady state solution would correspond to that particular TBA recirculation rate for which the TBA formation by the forward reaction is exactly balanced by the TBA dissociation by the reverse reaction (zero net TBA generation). This is the classic “recycle to extinction” of a minor component. In case some water (a reactant) leaks out of the process, a make-up water stream must be used to compensate for the loss 5

and maintain the total water flow to the reactor at the desired/specified value. For this process, the vapour-liquid-liquid equilibrium (VLLE) is such that some water does leak out in the C4 off-gas stream. A make-up water stream is therefore necessary to close the overall plant water balance.

Residue Curve Map In this work, we consider only conventional “reaction followed by separation” flowsheets. Hybrid schemes involving, e.g. reactive distillation, are not considered. In order to realize the process IO structure of Figure 1, we need an appropriate separation section that processes the reactor effluent to: a) Recover DIB at the desired purity plus any small amounts of generated TIB; b) Recover and purge the C4s (inerts + unreacted IB); c) Recover and recycle the water d) Recover and recycle the TBA. In consonance with industrial practice, our preference is for a distillation based separation scheme. For distillation phase equilibrium calculations, the liquid phase activity coefficients are modelled using the UNIFAC method, as the reactor effluent is a mixture of both non-polar and polar compounds. The vapour phase is modelled as ideal. The synthesis of a feasible distillation sequence to accomplish the above separation tasks is dictated by thermodynamic phase equilibrium constraints such as azeotropes, distillation boundaries and liquid-liquid phase split. To understand these fundamental constraints, Figure 2a plots the atmospheric pressure residue curve map (RCM) star diagram of the IB-DIB-TBA-H2O system along with the liquid-liquid phase split envelopes. The RCM composition phase space for a quaternary mixture is a tetrahedron with four triangular faces. For easier visualization, the three triangular faces of the tetrahedron (excluding the base triangle) are opened and the RCMs for the four ternary combinations are shown in the 2-D plot in Figure 2. Such diagrams for quaternary mixtures are commonly employed in the literature (see e.g.25). The RCM is only indicative in nature as the actual 6

reactor effluent contains other C4s, as well as small amounts of TIB. Three minimum boiling binary azeotropes, namely TBA-DIB, DIB-H2O, TBA-H2O and a heterogenous ternary minimum boiling TBADIB-H2O azeotrope are evident. The azeotropes partition the composition space into distillation regions. These regions and the liquid-liquid phase split envelopes are shown in the plot. The azeotrope compositions and boiling points are noted in Table 3. Figure 2a suggests that water in the reactor effluent may be easily recovered and recycled via liquid-liquid phase separation. The liquid-liquid equilibrium (LLE) tie lines are such that the aqueous phase is always close to the water vertex so that it can be directly recycled back to the reactor without any further processing. The organic layer however is likely to contain noticeable amounts of water. The RCM also suggests that since IB is an unstable node, it (plus other C4s) may be recovered as a distillate product. Post water and C4 recovery, the recovery of pure TBA and pure DIB is complicated by the presence of distillation boundaries. Since only a small reactor TBA concentration is needed to moderate IB oligomerization3, the concentration of the stream to be processed post water and C4 recovery is likely to be in the DIB rich Region II (see Figure 2a RCM). For such a feed composition, pure DIB can be recovered as the bottoms (stable node). The distillate composition is however constrained by the Region I – Region II boundary to a TBA rich stream with significant DIB (plus column feed water). If this stream is sent back to the reactor for recycling TBA to extinction, the DIB in the reactor feed would promote TIB formation, adversely affecting DIB yield. A separation strategy innovation is needed to ensure the TBA recycle stream contains no DIB so that TIB formation remains negligibly small. As a first potential option, we check the pressure sensitivity of the azeotrope compositions and the distillation boundaries. Figure 2b shows the azeotropes, distillation boundaries and regions at 1 bar and 9 bars. All azeotrope compositions change significantly with pressure. In particular, the TBA-DIB azeotrope mol fraction changes from ~67 mol% TBA at 1 bar to ~75 mol% TBA at 9 bars. This suggests that a DIB rich mixture can be distilled at high pressure (HP) into pure DIB bottoms (stable node) and a distillate close to the HP Region I – Region II distillation boundary. This HP distillate composition lies in the low pressure (LP) 7

Region I and LP distillation would give a near pure TBA bottoms (stable node) for recycle to the reactor, and a TBA-DIB-water distillate. The distillate may be recycled to the separation section beginning ensuring any traces of C4s and water exit the process and do not build in the recycle loop. Based on the above insights, we now present feasible flowsheets for DIB manufacturing with the IO structure of Figure 1. Here we discuss three alternatives, labelled FS0, FS1 and FS2.

Basic Flowsheet (FS0) In their process patent, Wang et al.9 present a process schematic for continuous DIB manufacturing exploiting pressure swing distillation for product recovery cum TBA recycle. The schematic is shown in Figure 3. Fresh C4 feedstock and make-up water along with TBA recycle are heated and sent to a liquid phase packed bed reactor. The reactor effluent is sent to a high pressure (HP) complex column. On a C4s and TIB free basis, its composition is in HP Region II (see Figure 2) so that nearly pure DIB (stable node) plus any heavy TIB is recovered as the bottoms product. C4 rich off-gas leaves up the top of the HP column. A C4 free side stream from a tray below the column feed is withdrawn and sent to a simple low pressure (LP) column. The composition of this side stream (on a C4s and TIB free basis) is close to the HP Region I – Region II boundary (see Figure 2b), which is inside the LP Region I. The simple LP column then recovers nearly pure TBA down the bottoms (stable node), which is recycled back to the reactor for suppressing TIB formation. The distillate, which contains lights (C4s), DIB, TBA and negligible water, is sent back to the HP complex column. This allows the lights to exit up the top. The DIB and TBA too have downstream exit locations in the HP column bottoms and LP column bottoms, respectively. Note that at high pressure, we have an additional minimum boiling IB-H2O azeotrope which acts as an unstable node (Figure 2b). This allows small amounts of water entering the HP column to exit up the top. All components, other than TBA, entering the separation section thus have a way out of the process. As discussed earlier, the TBA is recycled to extinction. This basic flowsheet, FS0, is thus a feasible process alternative.

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A major limitation exists in FS0 with respect to closure of the overall water balance for the plant. The TBA recycle stream contains negligible water so that no water gets recycled around the reactor. Thus, all the water that does not get converted to TBA (Reaction 3) must exit the process to close the water balance. In FS0, this water exits in the C4 off-gas. The maximum allowed water mol fraction in this stream is quite small, limited by the HP IB-water azeotrope (Figure 2b). This implies a maximum limit to the amount of fresh water that can be fed to the process, beyond which closure of the overall water balance becomes infeasible. This maximum fresh water limit constrains the maximum achievable TBA concentration inside the reactor. A higher TBA concentration for further suppressing TIB formation to improve DIB yield is then not possible. Due to this fundamental limitation imposed by the nature of the VLLE for this system, we found that FS0, although feasible, cannot provide both high IB conversion as well as high DIB product yield. This is clearly illustrated in exploratory steady state simulation results in Table 4. For 98% DIB yield, the reactor IB conversion is low at ~70%. This implies significant loss of precious IB in the off-gas. On the other hand, for a high IB conversion of 95%, the DIB yield is only 92%, implying significant loss of IB to TIB. A careful examination of the original patent data reported confirms both low-conversion and low-yield. Ideally, we would like a process design that achieves both high IB conversion and high yield to DIB. Innovations to the basic flowsheet topology are then needed that allows as much fresh water to be fed, as necessary to suppress TIB formation and thus ensure high yield. Our exploratory FS0 simulations showed that as the fresh feed water rate is increased, the reactor effluent composition naturally moves inside the liquid-liquid phase envelope. The liquid-liquid tie lines are such that the aqueous phase composition is close to the water vertex. We then get a near pure aqueous phase, which may be recycled back to the reactor. The water and TBA recirculation rate would then self regulate via the TBA formation reaction (Reaction 3). Based on this basic insight, in the following, we develop two alternative feasible flowsheet topologies that allow both high IB conversion as well as high DIB yield. To the best of our knowledge, these flowsheet topologies are new and have not been reported in the extant literature. 9

Flowsheet Alternative 1 (FS1) FS1 is schematically shown in Figure 4. The C4 feedstock and total (make-up + recycle) water are mixed with the TBA recycle stream, preheated in a feed effluent heat exchanger (FEHE) to recover heat from the hot reactor effluent and then heated to the desired reactor inlet temperature in a preheater. The reaction chemistry occurs in a cooled, liquid-phase packed bed reactor (PBR). The PBR is a shell and tube heat exchanger with catalyst loaded tubes and hot pressurized water coolant on the shell side. The hot reactor effluent is cooled in the FEHE, mixed with the DIB-TBA (with some water) recycle stream, further cooled in a water cooled cooler, depressurized to 1 bar and then decanted. The decanter aqueous layer (nearly pure water) is recycled to the PBR while the organic layer is fed to a three-column separation train. The first column recovers C4s (inerts + unreacted IB) as the distillate (D1) with the heavies (DIB, TIB, TBA and water) leaving down the bottoms (B1). B1, with its composition in Region II of the HP RCM, is distilled in the HP column to recover DIB and any TIB as the bottoms (B2). Its distillate (D2) composition is close to the HP Region I – Region II boundary (Figure 2b). D2 is then TBA rich with significant DIB and some water. It is fed to the LP recycle column, for which the feed (D2) composition is in Region I due to the pressure sensitive shift in the distillation boundary. This column recovers near pure TBA as the bottoms (B3), which is recycled to the reactor, and a distillate (D3) composition close to the LP Region I – Region II boundary. D3 then contains DIB and TBA, some water plus any C4 traces that dropped down the first column. D3 should not be recycled to the reactor as it contains significant DIB which can react with IB to form the side product, decreasing the DIB yield. The feasible recycle location is the decanter so that the C4s have a way out through the first column distillate and do not build up in the recycle loop. Also the other major components (DIB, TBA and water) have a way out of the recycle loop. Flowsheet Alternative 2 (FS2) FS2, schematically shown in Figure 5, is similar to FS1, except that the first two FS1 columns are combined into a single HP complex column, as in FS0. The FS2 reaction section is the same as 10

FS1. In consonance with the distillation regions in Figure 2b, HP rectification of the organic layer from the decanter causes the distillate to approach the C4-water azeotrope very close to the C4 vertex (unstable node). The C4s (inerts + unreacted IB) leave the process as the distillate (D1) from the HP column, carrying some water along. Similar to FS0, DIB (+ any TIB) is recovered as the HP column bottoms and we get a side draw composition that is inside the LP distillation Region I. The side-draw (SD) is further distilled in an LP column to recover pure TBA (Region I stable node) as the bottoms (B2), which is recycled to the reactor for suppressing the side reaction. The LP column distillate (D2) has a composition close to the LP Region I - Region II distillation boundary (TBA-DIB and some water) and is recycled to the decanter for reasons similar to FS1. Steady State Process Simulation Aspen Plus is used to simulate and close the material and energy balances of the flowsheet alternatives. We do not further consider the basic flowsheet, FS0, as it cannot ensure both high conversion and yield. Simulation details and results are presented only for FS1 and FS2. We wish to process 481.55 kmol/h of the C4 feedstock, (IB component flow rate = 100 kmol/h). For both FS1 and FS2, a base-case converged simulation is first obtained and then optimized with respect to the dominant design variables. Both flowsheets contain three material recycles (water recycle and TBA recycle to reactor; and DIB-TIB recycle to decanter) and an energy recycle through the feed effluent heat exchanger. A minimum of two recycle tears are needed to converge each flowsheet. However convergence to steady state was very difficult using only two tears. We, therefore used four recycle tears directly associated with the three material and one energy recycle loops. This resulted in better convergence to steady state. The reactor is simulated as a liquid phase plug flow reactor with multiple catalyst loaded tubes. The exothermic reaction heat is removed to a constant temperature reactor shell (TC). TC is set equal to reactor inlet temperature (TRxrIn). A typical liquid-to-liquid overall heat transfer coefficient of 350 W m-2 K-1 is used. The reactor has 100 tubes of 12 cm diameter, which gives a reasonable pressure drop at the final recommended reactor design/operating conditions. To study the effect of 11

reactor volume (VRxr), we vary the length of the tubes at constant tube diameter and # of tubes. In our simulations, we impose a maximum reactor inlet temperature constraint of 110 °C to mitigate accelerated high temperature catalyst degradation. Literature reports suggest that at temperatures above 125 °C, significant catalyst degradation is likely26. In the separation section, the FS1 C4 recovery column pressure is set at 4.5 bar for a condenser temperature of ~40 °C, allowing the use of cooling water. In both FS1 and FS2, the HP product column pressure is set at 9 bar and the LP recycle column pressure is set at 1 bar. For this pressure difference, the TBA-DIB azeotrope compositions differ by more than ~8 mol%, which is considered large enough for pressure swing induced crossing over of the Region I – Region II distillation boundary in Figure 2b. At 9 bar, pure DIB boils at ~200 °C which allows a high pressure steam heated reboiler. At the chosen pressures, the top temperatures of the HP and LP columns in both FS1 and FS2 are high enough for a water cooled condenser. To appropriately size the columns, we note that the phase equilibrium is highly non-ideal so that Fenske estimates of the minimum trays are likely to be unreliable. A heuristic approach is therefore used with the number of trays in the columns initially set at a large value (50) and the simulation converged for given column splits. The number of trays in each individual column is then reduced and set at a value slightly over the number of trays below which the column duty exhibits a sharp rise. In the FS2 complex column, the side-draw cannot be located too close to the feed as then its C4 content is too high and significant TBA escapes up the top (yield loss). Also, as the side-draw is moved lower down, its TIB content increases, which is undesirable since all of the side-draw TIB ends up in the TBA recycle stream, and is likely to promote further oligomerization representing expensive yield loss. We thus locate the side draw appropriately below the feed so that its C4 and TIB content is negligible (ppm level) and the reboiler duty is not excessively large. To converge the columns to the desired split, the Aspen Plus design spec vary feature is used, as detailed in Table 5.

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Process Design We now focus attention on obtaining reasonable values for the design variables using engineering heuristics. The Douglas doctrine27 clearly states that since raw materials are orders of magnitude more expensive than energy, first priority must be given to mitigating raw material wastage to side products or loss in by-product streams. The savings due to efficient raw material utilization (material efficiency) is usually enough to pay for the extra energy / capital costs. All design variables that significantly affect material efficiency are thus dominant design variables with a substantial effect on the process economics. A heuristic two-step design procedure is then a natural outcome of the Douglas doctrine. In the first step, reasonable values for the dominant design variables are obtained for ensuring efficient material utilization. Next, the remaining design variables are adjusted for improved energy efficiency. We follow this two-step heuristic procedure here.

Dominant Design Variables for Material Efficiency In the proposed DIB flowsheets, FS1 and FS2, efficient utilization of the IB reactant requires ensuring a high single-pass reactor conversion, since any unreacted IB leaves in the C4 off-gas (raw material loss), as well as a high yield to DIB with negligible TIB by-product formation. For clarity, the reactor single-pass conversion (X) is defined as X = (FRxnInIB – FRxrOutIB)/FRxrInIB where FRxrInIB and FRxrOutIB are the IB component molar flow rates into and out of the reactor, respectively. The yield or selectivity (S) to DIB is defined as S = 2.(FRxrOutDIB – FRxrInDIB)/(FRxnInIB – FRxrOutIB) where the numerator is the IB reaction rate to DIB and the denominator is the total IB reaction rate. The yield/selectivity is the fraction of reacting IB molecules that convert to DIB, the desired product. The reactor conversion and yield are determined by its design and operating conditions, which then are the dominant design variables. We consider the reactor volume (VRxr), inlet temperature (TRxrIn) and total water flow to the reactor (FH2ORxr) as the dominant design variables that 13

must be chosen to ensure efficient IB utilization (material efficiency). The same reaction section design/operating conditions are applied to both FS1 and FS2 as the reaction section in both flowsheets are structurally identical. The competing dominant reactor design tradeoffs then include 1. As single pass conversions approach 100%, exponential VRxr blow-up competes with precious IB loss in C4 purge at lower conversions. 2. Selectivity-conversion tradeoff with selectivity dropping as single-pass conversion increases. 3. With increasing temperature, selectivity decrease competes with lower VRxr for given near complete single-pass conversion. 4. With increasing FH2ORxr, higher in-situ TBA and consequent selectivity increase due to oligomerization suppression competing with higher TBA recycle rates and higher VRxr for a fixed single pass conversion. For a quantitative appreciation of the above effects, Figure 6a plots the variation in single-pass reactor conversion and selectivity to DIB with VRxr at different values of TRxrIn. Expectedly, IB conversion (X) increases with VRxr with a sharp flattening on approaching 100%. Also, selectivity (S) decreases with VRxr. For a high conversion of say 95%, VRxr at a higher temperature is noticeably lower than at a lower temperature. However the selectivity at the higher temperature is lower. If we plot conversion vs selectivity at the different TRxrIn, the curves are almost identical (see Figure 6b). These overlapping S-X curves clearly show a sharp drop in selectivity for conversions above 98%. We therefore size the reactor for 95% conversion so that the recommended design is sufficiently away from this precipitous drop in selectivity. This represents a reasonable compromise ensuring both high conversion and high selectivity. Since selectivity at a given single pass conversion is almost the same regardless of TRxrIn, TRxrIn is set at the maximum allowed (110 °C) for the smallest VRxr. With single pass conversion and TRxrIn thus fixed, we evaluate the effect of FH2ORxr on selectivity. Figure 6c plots the variation in selectivity and TBA recycle rate with FH2ORxr holding TRxrIn at 110°C (maximum allowed) and adjusting VRxr to hold conversion at 95%. Expectedly, as FH2ORxr is increased, the TBA recycle rate increases, which causes the selectivity to increase. The improvement 14

in selectivity per unit FH2ORxr increase however diminishes significantly above 35 kmol/h FH2ORxr. FH2ORxr is therefore set at 43 kmol/h to ensure the design is towards the flat region of the S- FH2ORxr curve.

Other Important Design Variables for Energy Efficiency With the reactor design and operating conditions fixed as above for both FS1 and FS2, we focus attention on adjusting important column splits to minimize total reboiler duty (QTot). In FS1, there are six degrees of freedom corresponding to the three simple columns. Of these, first column distillate TBA mol fraction and bottoms cis-butene mol fraction (xTBAD1 andxC4=B1), HP column bottoms TBA impurity (xTBAB2) and LP column bottoms DIB impurity (xDIBB3) are specified to ppm levels (see Table 5). This leaves the HP column distillate DIB mol fraction (xDIBD2) and LP column distillate DIB mol fraction (xDIBD3) as the two specifications that may be adjusted for minimizing QTot. We found that QTot is minimized when xDIBD2 and xDIBD3 are moved as far apart as possible. Referring back to Figure 2b, this implies that xDIBD2 specification should be lowered to be as close as possible to the HP Region I – Region II boundary, approaching from Region II. Similarly, xDIBD3 specification should be increased to be as close as possible to the LP Region I – Region II boundary, approaching from Region I. The reduction in QTot as xDIBD2 is decreased and xDIBD3 is increased is evident in Figure 7. The chosen design xDIBD2 and xDIBD3 values are also shown in the Figure. The converged recycle column material balance lines on a TIB and C4 free basis confirms that at the chosen specifications, D2 and D3 compositions are very close to the HP and LP Region I – Region II boundaries, respectively. In FS2, we have five steady state dofs, three for the HP complex column and two for the LP column. Ensuring ppm level cis-butene impurity in side draw, ppm level TBA impurity in B1 (DIB product) and ppm DIB impurity in B2 (TBA recycle) takes away 3 dofs. This leaves the HP column sidedraw DIB mol fraction (xDIBS) and LP column distillate DIB mol fraction (xDIBD2) as the two remaining dofs that may be adjusted to minimize the total reboiler duty (QTot). Purely from the energy efficiency point of view, xDIBS and xDIBD2 should be chosen as large and as small as possible, respectively, limited by the flowsheet convergence limit. However, for too large xDIBS, the TBA 15

impurity in the HP column distillate increases exponentially, implying loss of yield to DIB. To keep TBA loss in the C4 off-gas small at 50 ppm (same as FS1), xDIBS is set at 0.29. This finalizes the design of FS1 and FS2 for an overall IO structure as in Figure 1. The salient design and operating conditions are shown in Figure 4 and Figure 5 for FS1 and FS2, respectively.

Economic Comparison of FS1 and FS2 To compare the economics of FS1 and FS2, the total annualized cost (TAC) defined as TAC = Total Equipment Cost/PBP + Yearly Operating Cost where PBP is the payback-period, is calculated. Here, we use a payback period of 3 yrs, a typical value used in process design studies. The equipment cost correlations, sizing details, price data and process economics are noted in Table 6. Note that pumping and piping costs are ignored in our analysis. This is in consonance with standard conceptual process design practice, as noted by Douglas in his seminal text27. Table 7 provides an at-a-glance comparison of FS1 and FS2. The TAC of FS1 at $2.31x106 yr-1 is noticeably lower (~24.7%) than the FS2 TAC of $3.07x106 yr-1. This is attributed to lower total capital cost (10.9%) and energy cost (30.5%) of FS1. Even as FS2 has fewer columns, its HP complex column is highly energy intensive (6.3 MW reboiler duty) so that its height, diameter and reboiler/condenser areas and consequent capital cost is very large. FS1 also consumes less fresh water as it loses comparatively less water in the C4 off-gas. This is because the C4 condenser pressure in FS2 is higher (9 bar) compared to FS1 (4.5 bar). Its IB-water azeotrope is then further away from the IB vertex on the IB-water edge in Figure 2b. The azeotrope (unstable node) pulls the distillate composition towards it causing higher FS2 water leakage. The Table 7 comparison of FS1 and FS2 suggests that FS1 should be preferred over FS2 as it is cheaper in terms of the total capital cost and more energy efficient with lower water consumption per kg DIB product. Further the C4 off-gas has comparatively better fuel value with lower water content. We also note that FS1 is amenable to column-to-column heat integration. The FS1 HP 16

product column top vapour temperature is ~144°C while the LP recycle column bottom temperature is ~90°C. The product column condenser duty (QC2) is 1.3 MW while the recycle column reboiler duty (QR3) is 0.68 MW. The product column vapor may thus be used to provide reboil to the recycle column to further reduce steam utility consumption. Since QC2 > QR3, an auxiliary condenser is needed to condense all the product column vapour. Such heat integration is not possible in FS2 as the HP Column 1 top removes light C4s so that the top vapour is only ~70°C. Overall, these arguments and results suggest that FS1 should be preferred over FS2.

Plantwide Control To test the operability of FS1 and FS2 in the face of large C4 feed composition and throughput changes, a traditional decentralized PID plantwide control system with the throughput manipulator at the process feed is designed. The plantwide control system for FS1 and FS2 is shown in Figure 8 and Figure 9, respectively. Salient features of the control system include: 1. Fresh C4 feed is flow controlled and the flow setpoint is the throughput manipulator (TPM). 2. The reactor preheater duty is manipulated to hold TRxrIn. 3. Multiple temperature measurements along the length of the reactor are available. We found that the hot spot temperature moves along the length of the reactor for a change in throughput or a feed composition change. However the change in hot spot temperature is relatively small (< 2°C). The hot spot temperature is then not a good variable for reactor heat balance control. Instead, the initial temperature rise along the length of the reactor is maintained by manipulating coolant temperature. Specifically, the difference, ΔTRxr, between the temperature sensor at reactor length 0.5 m (total reactor length is 5.5 m; base-case hot spot at ~2 m) and TRxrIn is held constant. 4. On the decanter, Buckley’s28 material balance control scheme is applied. The aqueous layer outflow holds the aqueous-organic interface level. The total outflow (organic + aqueous) is maintained by manipulating the organic outflow with the total level controller adjusting the total 17

outflow setpoint. This arrangement ensures the organic outflow moves in response to the organic layer hold-up and compensates for the one-way interaction with the aqueous layer level. 5. On all columns, the reflux drum and bottom sump levels are controlled by the associated distillate and bottoms flow rates, respectively. The column pressure is regulated by adjusting the condenser duty. The reflux rate is maintained in ratio with the column feed rate. The average temperature of three sensitive stripping trays is maintained by manipulating the reboiler duty. In the FS2 HP column, the side-draw rate is adjusted to maintain the average temperature of three adjacent sensitive trays between the feed and the side-draw. For completeness, Figure 10 plots the base-case temperature profiles for the columns in FS1 and FS2 and Table 8 notes the controlled temperature variables for the FS1 and FS2 columns. 6. The fresh water flow is manipulated to hold the total water flow (FH2ORxr) to the reactor. The fresh water thus makes up for water loss in the C4 purge stream. 7. TIB concentration in the DIB product stream (xTIBProd) is available with a sampling time and dead time of 5 mins each. This concentration is regulated by manipulating the FH2ORxr setpoint. This loop maintains the process yield by holding the TIB mol fraction in the product stream. Note that in practice, a downstream column would separate DIB and TIB and the TIB flow to DIB flow ratio indicates the process yield. The same may be maintained by manipulating FH2ORxr. This would bypass the need for an expensive composition measurement. 8. Large changes in reactor single pass conversion show up in a large relative change in the TBA recycle rate (FTBARcy). Accordingly, the flow ratio rTBA = FTBARcy/[FTBARcy+FH2ORxr+FC4], where FC4 is the fresh C4 feedstock flow rate, is regulated by manipulating the reactor inlet temperature (TRxrIn), in order to mitigate large changes in reactor conversion. This controller is simply marked as RC on the TBA recycle line in Figure 8 and Figure 9. Two first order lags of 2 hrs each are applied in series to smoothen out high frequency transients in rTBA and only “look at” long term changes in rTBA due to slow plantwide effects.

18

We note that the loops for product stream TIB composition and TBA recycle rate regulation are “smart” economic loops that help maintain high yield, high conversion operation over the expected disturbance space. Since holding both the conversion and yield at their base-case values can be infeasible for large disturbances, at least one of the two controllers should be P only, which allows appropriate offset for feasibility of the final steady state. Here, we use a PI controller for regulating xTIBProd while the rTBA controller is P only. These two economic loops have an interesting cause and effect relationship. Let us start with the TIB composition loop. An increase in product TIB composition indicates a reduction in process yield (economic loss). To suppress TIB formation, the fresh water feed rate is increased by the loop, which increases the TBA generation rate and hence the TBA recirculation rate as well as the reactor TBA composition. The increased reactor TBA composition suppresses TIB generation reaction which causes TIB composition to decrease and thus the yield to increase back up. Now the increased TBA recycle rate (and increased water rate) causes reactor dilution as well as slowing down of the main reaction, which causes the IB conversion to decrease. The precious IB lost in the off-gas then increases (economic loss). To mitigate the decrease in conversion, inferred via an increase in the rTBA, the reactor inlet temperature is increased. The controllers in FS1 and FS2 are tuned as follows. All valves / control inputs are at ~50% of their span at the base-case. Both the process variable (PV) and controller output (CO) are scaled between 0 to 100% so that the controller gain has units of %CO/%PV (i.e unitless). All flow controllers are PI and use a gain (KC) of 0.5 and a reset time (τI) of 0.3 mins for a fast servo response. All level controllers are P only and use a KC of 2 to smoothen out flow transients. The column pressure controllers are PI and use a small τI and a large KC for tight pressure control. All temperature loops use a 0.5 min process variable lag to account for sensor dynamics. A 2 min lag is also applied to the output of all column temperature controllers that manipulate the reboiler duty to account for the thermal lag associated with the reboiler. The temperature controllers are tuned using the relay feedback test with Tyreus-Luyben settings. In the FS2 complex column, first the 19

reboiler loop is tuned with the side-draw temperature loop on manual. Then the side-draw loop is tuned with the reboiler loop on automatic. The PI xTIBProd (yield) and P only rTBA (conversion) controllers are tuned by hit-and-trial for a smooth plantwide response to the load disturbances. Table 9 reports the salient controller tuning parameters applied in the closed loop dynamic simulations for FS1 and FS2. Time scale separation exists between the various control loops with the two economic loops being the slowest with an open loop time constant of ~5 hrs (data not shown) due to the slow plantwide recycle effects. Multivariable interaction on similar slow time scales also exists between these two loops. All other loops are comparatively much faster by at least one order of magnitude. This may be verified from the PI controller reset time values in Table 9. Thus all flows and column pressures may be assumed to be perfectly controlled at their setpoints (time constant ~30-120s). The P only column surge levels have a residence time of 5-10 mins while the decanter has a residence time of ~20 mins. The corresponding level loop dominant time constants are then of a similar order (5-20 mins). The much slower recycle loop transients then determine the transients in the process flows manipulated by the level loops. Similarly, all temperature loop (columns, reactor, cooler etc) dominant time constants are significantly faster than the slow economic loops. The overall plantwide response is thus dominated by the slow economic loops and the multivariable interaction between them. The presented closed loop performance results should be viewed in light of the above. The closed loop performance of FS1 and FS2 to a ±20% throughput step change and a ±5 mol% step change in the C4 stream IB mol fraction (other components remain in same ratio) is evaluated. These disturbances are more severe than in reality as throughput changes are almost always effected as ramps (and not steps) in the TPM setpoint and feed composition changes are also quite gradual. Effective rejection of these severe disturbances would instil greater confidence in the appropriateness of the proposed process design and control system. Figure 11 and Figure 12 show the transient response to the throughput change disturbance for FS1 and FS2, respectively. Smooth 20

transients are observed in the plant flows with the plantwide response completing in about 40 hrs. The final steady state xTIBProd (product quality) returns to its base-case value as a consequence of the action of the PI xTIBProd controller so that the process yield remains high. Also, the transient deviation in xTIBProd is within a band of ±0.2 mol% around its base-case value. To better understand the closed loop response, we obtained the dynamic response for a ±20% change in throughput with both the xTIBProd and rTBA loops on manual (constant FH2ORxr and TRxrIn) for the two flowsheets (data not shown for brevity). In FS1, the final steady state xTIBProd is higher (lower) for the large throughput increase (decrease). Further, there is almost no change in the final rTBA. Thus, when both loops are closed, the xTIBProd controller slowly increases FH2ORxr. This leads to an increase in rTBA since higher water increases TBA generation, and in consequence, the direct acting rTBA controller increases TRxrIn. In FS2, the final steady state value of xTIBProd is lower (higher) for a 20% throughput increase (decrease) with both the economic loops on manual. Also, unlike FS1, the final rTBA is noticeably higher (lower). The decrease in xTIBProd is not surprising as the higher rTBA implies enhanced suppression of the TIB formation reaction. When both the economic loops are closed, the rTBA controller increases TRxrIn. The increase in TRxrIn now causes xTIBProd to increase (yield decrease) so that the final value at which FH2ORxr settles is above the base-case value. Figure 13 and Figure 14 show the transient response to the ±5 mol% step change in the fresh C4 feedstock IB mol fraction for FS1 and FS2 respectively. Similar to the throughput change disturbance, a smooth transient response is also observed in the plots. The response completion time is about 60 hrs with tight regulation of the DIB product purity. In both FS1 and FS2, the action of the two economic loops results in the TRxrIn increasing (decreasing) and FH2ORxr decreasing (increasing) for a large IB mol fraction increase (decrease). This may be understood as follows. The large increase in IB mol fraction causes higher in-situ TBA generation so that xTIBProd decreases, which in turn decreases FH2ORxr. Due to higher TBA generation, rTBA increases, which in turn causes an increase in TRxrIn.

21

One of the more challenging aspects of economic process operation of FS1 and FS2 is maintaining the single-pass reactor conversion at a high value in the presence of large throughput or feed composition changes. The high single pass conversion mitigates loss of precious IB in the C4 purge stream. Since the IB composition is not measured anywhere in the process, the challenge lies in inferring large changes in single pass conversion from available process data. With the product stream TIB mol fraction (xTIBProd) controller on automatic for maintaining high DIB yield, changes in single pass conversion may be inferred from large relative changes in the TBA recycle rate, as inferred from the flow ratio, rTBA. The flow ratio controller adjusts the reactor inlet temperature (TRxrIn). To assess the effectiveness of this strategy, we tested the closed loop process performance with the xTIBProd and rTBA controllers on manual and with both the controllers on automatic in the face of throughput and C4 feedstock composition change disturbances. Table 10 reports the change in xTIBProd (equivalent to DIB yield) and % single pass reactor IB conversion (X) achieved at the final steady state for the disturbances. At the base-case design, xTIBprod = 2% and X = 95%. The data in the Table suggests that with the two controllers on automatic, the variability in the yield and conversion is significantly reduced compared to when the two controllers are on manual. In particular notice the dramatic decrease in conversion and yield for an increase in the C4 feed IB mol fraction. This dramatic decrease in conversion/yield is significantly mitigated with the two controllers on automatic. The controllers thus help realize substantial economic benefit for the plant in actual operation over a large disturbance space. Overall the plantwide control results suggest that both FS1 and FS2 are operable with a traditional decentralized plantwide control system achieving safe, stable and economic process operation. In the final analysis, FS1 is recommended over FS2 due to its significantly lower TAC.

22

Conclusions This work has demonstrated the application of the complete design methodology of flowsheet synthesis, economic design and plantwide control for continuous DIB manufacturing via catalytic liquid-phase dimerization of IB in the C4 feedstock. The dimerization is performed in the presence of TBA, generated in situ by adding water, to moderate the further oligomerization of DIB to achieve viable process yield to DIB. The reactor effluent is separated using pressure swing distillation to bypass the distillation boundary due to the TBA-DIB minimum boiling azeotrope. Two new flowsheets, FS1 and FS2, have been quantitatively evaluated here. The use of a decanter to recover and recycle water from the reactor effluent in FS1 and FS2 provides the ability to set the reactor TBA content at the desired level, and achieve both high conversion and high yield. This is not the case with the extant literature flowsheet FS0 which is shown to achieve either high conversion or high yield but not both. In both FS1 and FS2, the reaction section is common and the reactor is sized for a high single-pass conversion (95%) to mitigate unconverted IB loss in the C4 purge. The reactor inlet temperature is fixed at the maximum allowed constraint (110 °C) imposed by the catalyst to minimize the reactor volume and hence the catalyst cost. The total water flow to the reactor is chosen to ensure a high process yield to DIB. A comparison the flowsheets shows that FS1 is significantly cheaper than FS2 with a 10.9% lower capital cost and 30.5% lower energy consumption. The fresh water consumption in FS2 is also found to be lower. Traditional decentralized PID control with the throughput manipulator at the fresh C4 feed is effective for stabilizing the two processes for large throughput and feed composition changes. Two unique features of the plantwide control system for economic process operation are (a) manipulation of the total water flow to maintain yield; and (b) manipulation of the reactor inlet temperature to maintain the TBA recycle rate ensuring regulation of reactor conversion to near its design value for large throughput / C 4 feed composition changes.

23

Acknowledgements The financial support from the Department of Science and Technology, as well as Ministry of Human Resource and Development, Government of India, is gratefully acknowledged.

24

Literature Cited 1. Anderson ST, Elzinga A. A ban on one is a boon for the other: Strict gasoline content rules and implicit ethanol blending mandates. J of Env Econ Mgmt. 2014;67:258-273. 2. Struckmanna LKR, Karinen RS, Krause AOI, Jakobsson K, Aittamaa JR. Process configurations for the production of the 2-methoxy-2,4,4-trimethyl pentane-a novel gasoline oxygenate. Chem Eng Proc: Process Intesification. 2004;43:57-65. 3. Honkela ML, Krause AOI. Kinetic modeling of the dimerization of isobutene. Ind Eng Chem Res. 2004;43:3251-3260. 4. Tidwell CM. Isobutene separation with a molecular sieve. US Patent 3,531,539, Sept. 1970. 5. Drake CA. Catalyst supports, catalyst systems, and olefin dimerization. US Patent 5,064,794, Nov. 1991. 6. Bercik PG, Henke AM. Isobutylene dimerization process. US Patent 3,832,418, Aug. 1974. 7. DiGirolamo M, Lami M, Marchionna M, Pescarollo E, Tagliabue L, Ancillotti F. Liquid-phase etherification/dimerization of isobutene over sulfonic acid resins. Ind Eng Chem Res. 1997;36: 4452-4458. 8. Izquierdo JF, Vila M, Tejero J, Cunill F, Iborra M. Kinetic study of isobutene dimerization catalyzed by a macroporous sulfonic acid resin. Appl. Catal A. 1993;106:155-165. 9. Wang J, Sahay N, Loescher ME, Vichailak M. Separation of tertiary butyl alcohol From Isobutylene. US Patent 6,863,778 B2, Mar. 2005. 10. Hanes RJ, Bakshi BR. Sustainable process design by the process to planet framework. AIChE J. 2015;61:3320-3331. 11. Ojasvi, Kaistha N. Continuous monoisopropylamine manufacturing: Sustainable process design and plantwide control. Ind Eng Chem Res. 2015;54:3398-3411. 12. Fermeglia M, Longo G, Toma L, Computer aided design for sustainable industrial processes: Specific tools and applications. AICHE J. 2009;55:1065-1078 13. Bildea CS, Dimian AC. Stability and multiplicity approch to the design of heat-integrated PFR. AIChE J. 1998;44:2703-2712. 14. Kumari P, Jagtap R, Kaistha N. Control system design for energy efficient on-target product purity operation of a high purity Petlyuk column, Ind Eng Chem Res. 2014;53:16436-16452. 15. Suresh Babu, K., Pavan Kumar, M.V., Kaistha, N. Controllable optimized designs of an ideal reactive distillation column using genetic algorithm, Chem Eng Sc. 2009;64:4929-4942. 16. Luyben ML, Floudas CA. Analyzing the interaction of design and control-2: reactor-separatorrecycle system. Comput Chem Eng. 1994;18:971–994. 17. Jagtap R, Kaistha N. Economic plantwide control of the ethyl benzene process. AIChE J. 2013;59:1996-2014. 18. Jagtap R, Kaistha N, Skogestad S. Economic plantwide control over a wide throughput range: A systematic design procedure. AIChE J. 2013;59: 2407-2426. 19. Luyben WL. Principles and case studies of simultaneous design. Hoboken NJ: John Wiley and Sons; 2011. 20. Yuan Z, Chen B, Sin G, Gani R. State-of-the-art and progress in optimization based simultaneous design and control for chemical processes, AIChE J. 2012;58:1640-1659.

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21. Ricardez-Sandovol LA, Budman HM, Douglas PL. Integration of design and control for chemical processes: A review of the literature and some recent results, Annual Reviews in Control. 2009;33(2):158-171. 22. Sharifzadeh M. Integration of process design and control: A review, Chemical Engineering Research and Design. 2013;91(12):2515-2549. 23. Hussom JK. Challenges and opportunity in integration of design and control, Comput Chem Eng. 2015;81(4):138-146. 24. Doherty M.F., Malone M.F. Conceptual design of distillation system. New York: McGraw Hill, 2001. 25. Chen B-C, Yu B-Y, Lin Y-L, Huang H-P, Chien I-L. Reactive distillation process for direct hydration of cyclohexane to produce cyclohexanol, Ind Eng Chem Res. 2014;53:7079-7086. 26. Honkela ML, Krause AOI, Kolah A, Aittamaa J. Isobutene dimerization in a mini-plant scale reactor. Chem Eng Proc. 2006;45:329-339. 27. Douglas J. M. Conceptual Design of Chemical Processes. New York: McGraw Hill, 1988. 28. Buckley P.S. Techniques of process control. Wiley: Upper Saddle River, NJ, 1964.

26

Hot Utilities / Electricity

C4s

Make-up H2 O

C4spurge

Sustainable DIB process DIB product Energy recycles

Material recycles

Cold Utilities Figure 1. Overall process input-output structure

27

(+ minor TIB )

Figure 2(b). Distillation boundaries and their pressure sensitivity

Figure 2(a). Residue curve map Azeotropic boundary Residue curves LLE boundary

Azeotropic boundary at 1 atm LLE boundary at 1 atm Azeotropic boundary at 9 atm LLE boundary at 9 atm

28

Figure 3. Flowsheet schematic (FS0) with overall material balance

29

Figure 4. Flowsheet alternative 1 (FS1) with salient design and operating conditions

30

Figure 5. Flowsheet alternative 2 (FS2) with salient design and operating conditions

31

(a)

(b)

(c)

Figure 6. Reactor design trade-offs

32

(a)

(b) B3

D2 D3

Figure 7. Choosing FS1 product and recycle column specifications (a) Variation of QTot with column specs (b) Recycle column material balance line for low QTot

33

RC

Figure 8. Plantwide control structure for FS1

34

RC

Figure 9. Plantwide control structure for FS2

35

Figure 10. Column temperature profiles for FS1 and FS2

36

Figure 11. FS1 plantwide response to ±20% throughput change Solid: +20% FC4 ; Dashed: -20% FC4 Brown: Controlled Variable; Blue: Manipulated Variable; Red: Setpoint

37

Figure 12. FS2 plantwide response to ±20% throughput change Solid: +20% FC4 ; Dashed: -20% FC4 Brown: Controlled Variable; Blue: Manipulated Variable; Red: Setpoint

38

Figure 13. FS1 plantwide response to ±5 mol% change in fresh feed IB composition Solid: +5 mol% xIB; Dashed: -5 mol% xIB Brown: Controlled Variable; Blue: Manipulated Variable; Red: Setpoint

39

Figure 14. FS2 plantwide response to ±5 mol% change in fresh feed IB composition Solid: +5 mol% xIB; Dashed: -5 mol% xIB Brown: Controlled Variable; Blue: Manipulated Variable; Red: Setpoint

40

Table 1. Reaction chemistry and kinetics # 1

2

3

Reactions

Rate expression

k = k0.exp{-Ea/RT}

IB + IB  DIB

𝑟𝐷𝐼𝐵 =

2 𝑘𝐷𝐼𝐵 𝑎𝐼𝐵 (𝑎𝐼𝐵 + 𝐵𝑇𝐵𝐴 𝑎 𝑇𝐵𝐴 )2

kDIB = 0.82.exp{-3.0/RT}

DIB + IB  TIB

𝑟𝑇𝐼𝐵 =

𝑘𝑇𝐼𝐵 𝑎𝐼𝐵 𝑎𝐷𝐼𝐵 (𝑎𝐼𝐵 + 𝐵𝑇𝐵𝐴 𝑎 𝑇𝐵𝐴 )3

kTIB = 0.065.exp{-1.8/RT}

𝑘𝑇𝐵𝐴 (𝑎𝐼𝐵 𝑎𝐻2 𝑂 − 𝑎 𝑇𝐵𝐴 𝐾𝑎) (𝑎 𝑇𝐵𝐴 + 𝐾𝐻2 𝑂 𝑎𝐻2 𝑂 )

kTBA = 0.21.exp{-1.8/RT}

IB + H2O  TBA

𝑟𝑇𝐵𝐴 =

Thermodynamic Package : UNIFAC Ea in KJmol-1; k0in mole.(hr.gCat)-1 aiis activity coefficient of species i; KH2O = 1.5; BTBA = 7.0;Tref = 373.15K

Table 2. C4 feed stock composition Component details

Mol fraction

ISOBUTYLENE

0.208

ISOBUTANE

0.276

1-BUTENE

0.156

N-BUTANE

0.1

Cis-2-BUTENE

0.156

Trans-2-BUTENE

0.103

41

Table 3. DIB-TBA-H2O azeotropes at 1 bar # # of components Characteristic Mole Frac* Temperature 1

DIB-H2O

Heterogeneous

0.5267

65.73 oC

2

H2O-TBA

Homogenous

0.3731

79.24 oC

3

DIB-TBA

Homogenous

0.3490

77.98 oC

4

H2O-TBA-DIB

Heterogeneous 0.3327, 0.3335

72.17 oC

*: Use ∑xi = 1 for last component mol fraction

Table 4. Conversion-yield effects on FS0 FH2Omax Conversion Yield (kmol/hr) 26

~70%

~ 98%

23

~85%

~96%

18

~95%

~92%

Table 5. Design spec vary as applied to FS1 and FS2 columns Flowsheet 1 Column

Spec

Spec Value

Vary

xD1TBA cis-2-butene

50PPM 1PPM

Reflux ratio Reboiler duty

HP product column

xD2DIB xB2TBA

0.25 50PPM

Reflux ratio Reboiler duty

LP recycle column

xD3DIB xB3DIB

0.325 50PPM

Reflux ratio Reboiler duty

C4 purge column

B1

x

Flowsheet 2

xD1DIB HP product column

LP recycle column

xB1TBA

0.25 1000ppm 50ppm

Reflux ratio SD rate Reboiler duty

xD2DIB xB2DIB

0.325 50ppm

Reflux ratio Reboiler duty

SD

x

cis-2-butene

42

Table 6. Process equipment cost correlations, sizing details and process economics for FS1 and FS2 Process Economics Details 19

Equipment cost correlation and pricing details

Flowsheet 1 (FS1) Capital Cost Energy Cost 6 $ 10 $ 106

(equipment dimensions are shown in Fig 4/Fig 5)

C4 Purge Column

High Pressure DIB Product Column

Low Pressure Recycle Column Decanter Heat Exchanger Reactor

Tower sizing utility (Fair’s method) Tray spacing: 2 ft. 20% extra height for sump Tower cost: $17640(D)1.066(L)0.802 D: diameter , L: length Heat Transfer coefficients, condenser/reboiler : 0.852 / 0.568 kW K-1m-2 Differential Temperature, condenser/reboiler: 15 K / 35 K Capital cost : 7296(area)0.65 HP steam $9.83 GJ-1 / LP steam $7.78 GJ-1 Cooling water $0.16 GJ-1

Flowsheet 2 (FS2) Capital Cost Energy Cost 6 $ 10 $ 106

Reboiler

0.195

0.814

--

--

Condenser

0.253

0.012

--

--

Tower

0.447

--

--

--

Reboiler

0.110

0.33

0.310

1.63

Condenser

0.107

0.004

0.329

0.01

Tower

0.290

--

0.745

--

Reboiler

0.072

0.16

0.114

0.35

Condenser

0.126

0.004

0.222

0.01

Tower

0.242

--

0.356

0.20

--

0.26

--

$17640(D)1.066(L)0.802 U = 0.568 kW K-1m-2, ∆T = 35 K

HX1

0.066

0.15

0.066

0.15

U = 0.852 kW K m , ∆T = 15 K

HX2

0.120

0.004

0.120

0.004

0.14

0.003

0.14

0.003

0.04

--

0.04

--

0.05

--

0.05

--

2.45

1.50

2.74

2.16

-1

$17640(D)

-2

1.066

(L)

-1

0.802

-1

Catalyst

$8.33 kg yr

FEHE

U = 0.568 kW K m , ∆T = 35 K -1

-2

Total

Payback period 'n' = 3 years

TAC

TAC = Total capital cost / n + Total energy cost 43

2.31x106 Yr-1

3.07x106 Yr-1

Table 7. At-a-glace process performance and economics comparison table Parameter

FS1 0.02 60 ppm

FS2 0.02 40ppm

moisture content (kg) / kg of fuel gas

0.008

0.017

MW Steam/Kg DIB

4.52

5.73

MW Cooling/Kg DIB

2.98

5.29

TAC Total Capital Cost Total Energy Cost

$ 2.31x106 $ 0.81x106 $ 1.50x106

$ 3.07x106 $ 0.91x106 $ 2.16x106

TIB TBA

DIB product impurity Process performance

Economic

Table 8. Controlled temperatures for FS1 and FS2 columns Equipment Column

Purge column

FS1 CV

FS2 MV

(T24+T26+T28)/3 Reboiler duty

CV

MV

--

--

(T26+T27+T28)/3 Side draw rate HP product column (T22+T23+T24)/3 Reboiler duty (T40+T41+T42)/3 Reboiler duty LP recycle column (T +T +T )/3 Reboiler duty (T +T +T )/3 Reboiler duty 26 27 28 26 27 28

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Table 9. Salient controller parameters for FS1 & FS2*# FS 1 CV

KC

τi min

SP

PV Range&

MV Range&

TSCol1 TSHPCol2 x TIB B2→ FRxrH2O TSLPCol3 rTBA → TRxrIn TRxrin ∆TRxr

0.55 0.5 1 15 0.5 0.5 5

11.55 10 300 2.64 -10 0.5

69.85 oC 170.78 oC 0.98 84.12 oC 16.3 98 oC 10.516

35-100oC 85-255oC 0.90-0.99 0-155oC 0.005-0.04 55-128oC 0-15 oC

8.3MW 2.7MW 30-65 kmol/hr 1.27MW 80-120oC 2.5MW 70-115 oC

68-122oC 111-222oC 0.90-0.99 60-120oC 0.005-0.04 55-128oC 0-22 oC

.01-100 kmol/hr 12.2MW 30-65 kmol/hr 2.78MW 80-120oC 2.5MW 70-120 oC

FS 2 TRHPCol1 0.544 44.55 106.985 oC TSHPCol1 3.26 6.6 176.5 oC x TIB B2→ FRxrH2O 1 300 0.98 S T LPCol2 16.5 5.28 83.79oC rTBA→ TRxrIn 0.3 -16.5 TRxrin 0.5 10 98 oC ∆TRxr 5 0.5 9

*: All level loops use KC = 2 unless otherwise specified #: Pressure/flow controllers tuned for tight control &: Minimum value is 0, unless specified otherwise

Table 10. Effect of supervisory loops on conversion (X) and product impurity (xTIB) regulation FS1 FS2 Disturbance CS1 CS2 CS1 CS2 CS1 CS2 ΔX ΔX ΔxTIB ΔxTIB ΔX ΔX ΔxTIBCS1 ΔxTIBCS2 Throughput -20%

2.06

2.19

-0.27

0.0

3.52

2.30

0.89

0.0

Throughput +20%

-3.14

-3.09

0.26

0.0

-5.69

-3.57

-0.34

0.0

Feed IB molfr -5%

5.10

1.72

8.89

0.0

4.69

1.77

9.11

0.0

Feed IB molfr +5%

-7.95

-1.90

-0.92

0.0

-11.06

-2.38

-1.31

0.0

CS1: TBA ratio and TIB impurity controllers on manual CS2: TBA ratio and TIB impurity controllers on automatic

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