Techno-Economic Analysis of a Chemical Process to Manufacture

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Techno-Economic Analysis of a Chemical Process to Manufacture Methyl-#-Caprolactone from Cresols David Lundberg, Daniel Lundberg, Marc A. Hillmyer, and Paul J. Dauenhauer ACS Sustainable Chem. Eng., Just Accepted Manuscript • DOI: 10.1021/ acssuschemeng.8b03774 • Publication Date (Web): 28 Sep 2018 Downloaded from http://pubs.acs.org on September 30, 2018

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ACS Sustainable Chemistry & Engineering

Techno-Economic Analysis of a Chemical Process to Manufacture Methyl-ε-Caprolactone from Cresols

David J. Lundberg1,3*, Daniel J. Lundberg1,3*, Marc A. Hillmyer2,3, Paul J. Dauenhauer1,3† 1

University of Minnesota, Chemical Engineering and Materials Science, Amundson Hall, 421 Washington Ave. SE, Minneapolis, MN 55455. 2 University of Minnesota, Department of Chemistry, Smith Hall, 207 Pleasant Street SE, Minneapolis, MN 55455. 3 Center for Sustainable Polymers, a National Science Foundation Center for Chemical Innovation. Smith Hall, 207 Pleasant Street SE, Minneapolis, MN 55455. *Authors contributed equally. † Corresponding Author: [email protected] Keywords: Caprolactone, Cresol, Elastomer, Hydrogenation, Oxidation

Abstract. Methyl-ε-caprolactone is a monomer used in the manufacture of elastomeric biodegradable polymers with enhanced properties. We present here the conceptual process design for the production of methyl-ε-caprolactone from p-cresol, a bio-renewable feedstock obtained from lignin. The two-reaction process consists of cresol hydrogenation to methyl-cyclohexanone followed by Baeyer-Villiger oxidation to methyl-ε-caprolactone. Details for designing an optimized process include unit operation design of two reactors, one decanter, one flash tank, and five distillation columns. Distillation and integrated heat transfer were determined via process simulation, with the objective of optimizing the process for net present value. For cresol obtained at $1.00 kg-1, the minimum selling price of methyl-ε-caprolactone (defined as the product selling price that results in a zero net present value of the entire 30-year project) was $3.521, $2.798, and $2.557 kg-1 ($1.600, $1.272, and $1.162 lbm-1) at three process scales of 10, 30, and 60 kTon yr-1 of pcresol feed. A sensitivity analysis of the major process variables identified catalyst selectivity to methyl-εcaprolactone in the Baeyer-Villiger oxidation reaction as the key parameter for improving process economic potential.

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Introduction. The transition to a sustainable plastics economy requires the discovery of new, highperforming polymers that can be economically manufactured from renewable resources[1]. Across all polymer types and market sectors, the potential of plastics from plants includes the opportunity to sustainably obtain the feedstocks (i.e., monomers) that are assembled into polymeric structures with strategic end-of-life options including biodegradation or recycling[2]. The past decade has seen the emergence of drop-in replacement monomers from biomass including ethylene[3], p-xylene[4.5], acrylonitrile[6], isoprene[7,8], and butadiene[9,10], while entirely new polymers based on FDCA[11,12] (furan dicarboxylic acid), caprolactones[13,14], and a broad range of diols[15,16] comprise new materials with advanced elasticity, biodegradability, or barrier properties. Implementation of these biomass-derived materials will depend on the economics and environmental impact in all three life-cycle phases: synthesis and manufacturing, application, and end-of-life processing[17]. Emerging advanced polymers utilize branched cyclic lactones such as methyl-ε-caprolactone to produce biodegradable elastomers. Conventional poly(ε-caprolactone) (PCL) from straight-carbon-chain ε-caprolactone monomer produces a semi-crystalline material that can be effectively blended with other polymers such as starch or poly(ethylene oxide)[18]; the ester moieties in the polymer backbone also impart biodegradability[19]. The branched variant, poly(methyl-ε-caprolactone) (PMCL) in addition to other alkylgroup-containing

caprolactones

such

as

poly(propyl-ε-caprolactone),

also

benefits

from

biodegradability/hydrolyzability[20], but the addition of branching leads to a completely amorphous material with low glass transition temperature (Tg ≈ –60 °C).[21,22,23,24]

PMCL has also been utilized in the

preparation of unique block co-polymers which self-assemble into polymersomes[25]; combined polymer blocks also include poly(ethylene glycol)methyl ether[26], poly(n-isopropylacrylamide)[27], poly(lactic acid)[28,29], and poly(ethylene oxide)[30,31]. PMCL-derived block polymers have comparable performance to Kraton brand specialty polymers which can sell for as much as $7 per kg [32],[33]. We are particularly enthusiastic about the ability to generate resilient, strong and degradable elastomers using methyl-εcaprolactone[28,34]. The manufacture and application of these polymers relies on a viable chemical pathway between a sustainable biomass-derived feedstock and methyl-ε-caprolactone produced at high purity. As depicted in Figure 1, a chemical route proposed by Roman[35], Hillmyer[29] , Sels[36], and Hammond[37] utilizes alkylphenols such as those obtained from lignin including m- or p-cresol or guaiacols. As part of a two-step process, alkyl-phenols are initially hydrogenated to form alkyl-cyclohexanones; for example, cresol is hydrogenated to form methyl-cyclohexanone. In the second reaction step, Baeyer-Villiger oxidation (BVO) of alkyl-cyclohexanones forms alkyl-caprolactones; methyl cyclohexanone is oxidized to methyl-εcaprolactone. These combined reaction steps provide flexibility to a range of lignin-derived alkyl-phenols while combining two selective chemistries capable of producing alkyl-caprolactones with a net high yield.

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The hydrogenation of alkyl-phenols aims to selectively break aromaticity and form a cyclic ketone. Upon initial hydrogenation of the six-carbon ring, the hydroxyl group undergoes keto-enol tautomerization and forms a carbonyl[38]; catalyst and conditions capable of selectively halting hydrogenation prior to carbonyl reduction with supported Pd, Pt and Rh have been identified[39,40,41]. For example, palladium on hydroxyapatite (HAP) has achieved 97% selectivity to cyclohexanone from phenol at less than five bars of hydrogen pressure in water at 100 °C.[39] Higher selectivity to cyclohexanone from phenol and various alkyl-phenols (including o-, m-, and p-cresol and p-cumenol) has been reported by Xu et al. using a HAP supported palladium catalyst under atmospheric hydrogen pressure in water at 75 °C[40]. Other approaches have utilized a combined heterogeneous Pd catalyst with homogeneous Lewis acids such as AlCl 3 to achieve >99.9% selectivity to cyclohexanone at complete conversion of phenol[41]. The intermediate methyl-cyclohexanone can be oxidized to methyl-ε-caprolactone via a BVO reaction. While BVO is commonly conducted with organic peroxides such as trifluoroperacetic acid or meta-chloroperoxybenzoic acid[42], lower cost hydrogen peroxide has been evaluated for oxidation with heterogeneous Lewis acid catalysts including Sn-containing zeolites[43,44], which can activate ketones in aqueous media[45]. Roman evaluated a range of Sn-containing microporous and mesoporous materials and demonstrated conversions as high as 99% with >98% selectivity to lactones from cyclic 2-adamantone using H2O2[43]. Hammond and coworkers showed that tin-beta (Sn-BEA) catalyst and hydrogen peroxide achieved 90% selectivity to methyl-ε-caprolactone at 20% conversion of methyl-cyclohexanone in batch experiments and was suitable to operate continuously[37]; higher conversion lowered the overall yield of alkyl-caprolactone product.

The major byproduct of methyl-cyclohexanone BVO is methyl-

hydroxyhexanoic acid, which results from the hydrolysis of methyl-ε-caprolactone (See Figure 1). We note that methyl-hydroxyhexanoic acid can be converted back to methyl-ε-caprolactone, but this process involves multiple fractionation and reaction steps and is not considered here. [46] In this work, the combination of sequential alkyl-phenol hydrogenation and oxidation via heterogeneous catalysis of cresol will undergo techno-economic analysis as part of a chemical process with reaction, separation, and ancillary process equipment. Hydrogenation within a trickle-bed catalytic reactor will convert cresol to methyl-cyclohexanone, which is purified and sent to a liquid phase fixed bed oxidation reactor. The resulting methyl-ε-caprolactone is then purified by distillation. Key economic parameters including the minimum selling price of methyl-ε-caprolactone will be identified, and a rigorous economic evaluation of each process block will be conducted.

Methods. Based on the chemistries of alkyl-phenol hydrogenation and BVO, a process has been developed to assess the economic and technological feasibility of the conversion of p-cresol into 4-methyl-εcaprolactone (MCL). The developed process combines the two subsequent reactions, recycle, and product

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purification steps for the hydrogenation of p-cresol to 4-methyl-cyclohexanone (MCH) and then the BVO of MCH to MCL. The process is divided into four blocks shown in Figure 2: (1) hydrogenation of p-cresol to MCH and vapor recovery, (2) MCH purification, (3) BVO of MCH to MCL with recycle of unreacted MCH, and (4) the purification of MCL. Process Chemistry and Thermodynamic Simulation. Aspen Plus (V8.6 Aspen Technology) was employed to simulate the reaction, separation, and mass balances of the designed process. The PengRobinson equation of state was used to simulate all separation operations. Distillation columns were modeled using the Aspen RadFrac module. Accounting of mass and cash flows were conducted with Microsoft Excel. Cresol Hydrogenation. In the designed process, liquid phase hydrogenation of p-cresol to MCH was performed under atmospheric pressure of hydrogen using a Pd/HAP catalyst. Based on previously reported results for this catalytic system it was approximated that 98% conversion of p-cresol and 97% selectivity to MCH can be achieved at 75 °C. Catalyst activity was directly taken to be 0.014 mol gcat-1 hr1

from the work of Xu et al.[40] While this result was reported for the hydrogenation of p-cresol in water, n-

dodecane was instead used in the first reactor as solvent to avoid the azeotrope formed from water and MCH. It is noted that Xu et al. demonstrated improved kinetics for the hydrogenation of phenol in ndodecane as compared to water. Based on these results, a reactor was modeled for the liquid phase hydrogenation of p-cresol (R-1). The configuration of a three-phase trickle-bed reactor was implemented for R-1 as shown in Figure 2; in practice, a chemical reactor that maximizes gas/liquid contacting for sufficient hydrogen dissolution in the reaction solvent would be used. It was assumed that deactivation of the Pd/HAP catalyst was mainly caused by formation of coke and not sintering due to the relatively low reaction temperature.[47] Catalyst lifetime was approximated to be at least 15 hours[40]; reactant feed will be regularly switched between a parallel reactor, while the coke deposits in the empty reactor are removed by oxidation at 450 °C for six hours48. Initially, a mixture of p-cresol and n-dodecane with a reactant to solvent molar ratio of 0.187 (half the solubility limit of p-cresol in n-dodecane at 75 °C, as estimated in Aspen Plus) was heated to reaction temperature (75 °C) and fed to the top of the reactor. Hydrogen was fed to the bottom of the reactor. Excess hydrogen was collected from the top of the reactor and recycled using a blower (B-1). Liquid effluent from the reactor was cooled to 45 °C and fed to an atmospheric pressure flash tank (F-1) where 8% of dissolved hydrogen was removed from the condensate. The flashed vapor was recycled through a recycle stream via the blower. Methyl-Cyclohexanone Purification. The liquid flash stream containing MCH, side product 4methyl-cyclohexanol (MCOH), unreacted p-cresol, and n-dodecane solvent was pumped to a distillation column (C-1) where n-dodecane and p-cresol were recovered from the distillation column bottoms (stream

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8) and recycled back to the hydrogenation reactor feed. The vapor distillate (stream 10) contained hydrogen and MCH that was recycled through the vapor recycle system. MCH and MCOH were recovered in the liquid distillate (stream 9), combined with recycled MCH and MCOH from C-4, and fed to a second distillation column (C-2) where MCOH was removed in the bottoms (stream 11) and pure MCH (99 mol%) was recovered in the liquid distillate (stream 12) and sent to the BVO reactor (R-2). The vapor distillate from C-2 (stream 13) was recycled through the vapor recycle system. A 5% fraction of recycled vapor from F-1, C-1, and C-2 was purged to prevent accumulation of impurities in the process[49]. Dehydrogenation of MCOH back to MCH was not considered in this process due to the high selectivity achieved in R-1. MCL Production. MCH (stream 12) was mixed with a 50 wt% aqueous hydrogen peroxide solution (stream 14), heated to the reaction temperature of 50 °C, and fed at 1.0 bar over a Sn-BEA catalyst in a packed bed flow reactor (R-2). Reactor performance was based on the results of Yakabi et al.[50]; catalyst space time yield was based on batch experimental results of 24.5 g-lactone kgcat-1 hr-1 per cubic centimeter of reactor volume. In the reactor, the hydrogen peroxide reacted to completion with 20% of the fed MCH to generate 4-methyl-ε-caprolactone (MCL). 10% of the generated MCL was further hydrolyzed to produce 4-methyl-6-hydroxyhexanoic acid (4-M-6-HHA). Yakabi et al. report the formation of trace amounts of 4-methyl-adipic acid at 100% conversion[50]; this potential byproduct formation when operating at 20% conversion was assumed negligible in the design. Trace amounts of 4-methyl-adipic acid present in the reactor effluent would not affect the downstream separation process, since it would be removed alongside 4-M-6-HHA in C-3 and C-5. MCL Purification. By the method of recycle outlined in Figure 2, 100% net conversion of MCH achieves 90% selectivity to MCL in the BVO reactor. The effluent from R-2 was sent to a vacuum distillation column (C-3) operated at 0.2 bar from which 4-M-6-HHA and MCL were recovered in the bottoms (stream 16) and sent to a second vacuum distillation column (C-5) operating at 0.056 bar. Polymerization grade MCL (99.9 mol%) was recovered in the distillate from this column (stream 16), while the major byproduct of 4-M-6-HHA was removed in the bottoms (stream 18). Vacuum distillation was used to ensure minimal MCL losses due to polymerization by reducing operating temperatures within C-3 and C-5. The chosen column pressures were comparable with previously reported purification of εcaprolactone via vacuum distillation to minimize polymerization.[51,52,53,54] MCH Recycle/Recovery. The liquid distillate from C-3 (stream 17) contained 80% of the MCH fed to R-2 and a significant amount of water from co-fed H2O2 aqueous solution as well as byproduct water from the BVO reaction. Water must be removed from the MCH before it is recycled to R-2 in order to prevent additional hydrolysis of MCL. This stream was cooled to 45 °C and sent to a decanter (D-1) where 90.7% of the water was removed in an aqueous stream. The aqueous phase from D-1 was sent to wastewater treatment (stream 20) and the organic phase (stream 21) was sent to a distillation column (C-4), the distillate

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of which contained a mixture of MCH and water that was recycled back to the decanter (stream 23). MCH with an impurity of MCOH was obtained from the bottom of C-4 (stream 22) and needed to be pumped back earlier in the process to C-2, where remaining MCH was purified to prevent accumulation of MCOH in the process.

Results and Discussion. The designed chemical process of Figure 2 was developed to minimize the total project cost over the lifetime of 30 years. Key technical decisions were assessed based on their impact on the minimum selling price (MSP) of the product methyl-ε-caprolactone, defined as the price that results in zero net present value of the entire project using the method of capital and operating cost discounting. Full economic details are provided in the supporting information. Capital and Operating Costs. Based on the simulation results, the capital costs and variable operating costs of most process equipment (reactor vessels, flash tanks, distillation columns, decanter, and blower) in addition to fixed operating costs for the plant (labor, insurance, maintenance, etc.) were estimated using Aspen Process Economic Analyzer (V8.4 Aspen Technology) in 2013 USD. Vacuum pumps are not standard simulation blocks in Aspen and were sized separately (see Supporting Information section 6). Total catalyst loadings in both reactors were calculated using data from the literature (see Supporting Information sections 4-5). All equipment costs were indexed to the year 2017 using the Chemical Engineering Plant Cost Index[55]. Utilities and feedstocks were assumed to be available for purchase at the prices listed in Table 1. The base case processing rate of p-cresol was selected to be 10,000 metric tons per year, which could produce as much as 37,500 tons of PLA-PMCL-PLA thermoplastic elastomers, where ‘PLA’ represents ‘poly(lactic acid)[28]. This product quantity is equal in mass to approximately 2% of the global styrene block copolymer (SBC) thermoplastic elastomer market[56]. Lignin conversion via reductive or acid-catalyzed depolymerization can produce alkyl phenols with up to 30% yield[57]. Over 30 kTon yr-1 of lignin would need to be converted to supply the feedstock for the base case process. This represents nearly a fifth of the total capacity of a single biorefinery, which can process up to 146 kTon yr-1 of lignin[58]. The total capital and operating costs for the process are listed in Table 2. For the base case plant size capable of processing 10,000 metric tons per year of p-cresol, the total installed equipment costs were $25.0 million, whereas total capital investment (TCI), which includes costs such as project contingency, site construction and engineering, etc., was $53.7 million. Capital costs were concentrated in the purification blocks where large molar flow rate of mixed streams were separated to high purity streams. For plant sizes of 30 and 60 kTon yr-1, TCI increased to $91.8 and $140.9 million, respectively, with over 72% of installed equipment costs coming from purification blocks. The process operating costs for varying plant size (10, 30 and 60 kTon p-cresol yr-1) are listed in Table 3. For a plant size of 10,000 tons of p-cresol per year, the total yearly operating cost of the process

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was estimated to be $19.1 million, of which feedstock costs comprised 68.6% of the total. Feedstock cost was dominated by the purchase of p-cresol, which accounted for over 90% of total feedstock costs. Fixed operating costs were $3.7 million per year (which included labor, maintenance, insurance, etc.). For plant sizes of 30 and 60 kTon yr-1, fixed operating costs become a smaller proportion of total operating costs, and steam costs similarly dominated total utility costs. Minimum Selling Price. The minimum selling price (MSP) of MCL (defined as the price that results in a zero net present value of the entire project) was calculated using the discounted cash flow methodology[59] based on economic parameters and assumptions listed in Table S2; those of most interest are: 30 year plant lifetime, 15% minimum annual rate of return (MARR), and 35% tax rate. Based on this analysis, the MSP of MCL for the base case (10,000 tons p-cresol per year) was calculated to be $3.521 kg1

($1.597 lbm-1). The sensitivity of MSP to economic parameters such as the MARR or “hurdle rate”, tax

rate, and process lifetime is illustrated in Figures S1 to S3. Figure 3 depicts the contribution to minimum selling price (MSP) of methyl-ε-caprolactone by process block identified by row; the contribution towards MSP of capital, fixed operating, utility, and feedstock costs are identified by color. The largest single contribution to MSP derived from feedstock costs ($1.529 kg-1), consistent with a large-volume commodity chemical process. The next largest contributors to the MSP were capital costs ($1.351 kg-1), fixed operating costs ($0.395 kg-1), and utility costs ($0.245 kg-1). The ketone and caprolactone purification blocks (including capital and operating costs) accounted for the majority of the process contribution to the MSP ($1.637 kg-1). Overall, the ketone purification process block accounted for the largest contribution of the process to the MSP of MCL. This is due to the two high-volume separations being performed; recovery of unreacted p-cresol and n-dodecane solvent from R-1 effluent in C-1, and the removal of byproduct MCOH from MCH in C-2. Purification of MCH is unavoidable if hydrogenation selectivity is below 100%, and the recycle of stream 22 to C-2 further increases the cost of the column. Chemical Process Base Case Design. This base case design provided an initial technically-feasible process which could manufacture methyl-ε-caprolactone to the purity specifications of 99.9%. The combination of hydrogenation and BVO of p-cresol feedstock required four significant technical design decisions impactful to the overall base case process design. Solvent Selection. The selection of a solvent for the two reactors (hydrogenation in R-1 and Baeyer-Villiger oxidation in R-2) directly affects catalyst performance while also impacting the cost of separations. In reactor R-2, purified methyl-cyclohexanone is combined with an aqueous stream of hydrogen peroxide; prior to C-2, the selected solvent will impact the catalytic performance of reactor R-1, flash tank F-1, and distillation column C-1.

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The base case design selected n-dodecane as a solvent for hydrogenation based on its inertness during reaction and ease of separation from the reactants and products. Xu et al. probed the applicability of different solvents on the hydrogenation of phenol over Pd/HAP and found linear alkanes to exhibit the best performance[40]. Alternative linear alkanes may be applicable in this process, but the hydrogenation solvent is desired to have a sufficiently low vapor pressure to remain in the liquid stream of flash tank F-1 while also being sufficiently heavy to be recovered from the bottoms of C-1 in stream 8 to recycle back to reactor R-1. Ultimately, n-dodecane was found to be the optimal linear alkane solvent due to its boiling point similar to that of p-cresol (see Supporting Information for additional details). Additionally, the cost associated with C-1 could be lowered by running the reaction neat. It is noted that Liu et al. have demonstrated the solvent-free hydrogenation of phenol to cyclohexanone under an atmospheric pressure mixture of a H2 and CO2 over a supported palladium catalyst (>95% selectivity and 70% conversion), but the use of water as solvent improved reaction kinetics60. If the hydrogenation reaction could maintain its selectivity and conversion while being operated without solvent, the size and operating cost of column C-1 would decrease significantly (saving $11.5 million in TCI, and $68.90 hr -1 in utility costs) and lower the base case MSP by $0.335 kg-1. Heteroazeotropic Distillation of Reactor R-2 Effluent. To recover and recycle the unreacted MCH from reactor R-2, heteroazeotropic distillation was required. The distillate of C-3, containing water, MCH, and trace amounts of MCOH was predicted by Aspen to form a heteroazeotrope, which is an azeotrope where the vapor phase coexists with two immiscible liquid phases. Traditionally, binary heteroazeotropes are separated as a two-distillation-column process wherein one column is fed the mixture to be separated, and the distillate of both columns is condensed, decanted, and each component-rich phase is refluxed back to a different column. Each of the two components is obtained as the bottoms of the respective column in which its phase is refluxed. It is important to note that heteroazeotropes are always minimum boiling mixtures[61], thus explaining the counterintuitive result of obtaining pure streams of each component as bottom streams out of the columns. This process has been described in the industrially important drying of n-butanol62, ethanol63, and acetic acid64, though the latter two make use of an entraining component and thus deal with ternary systems. A modified version of heteroazeotrope distillation was modeled in this process to recover MCH. Only a single column was required to obtain separation streams of water and the MCH/MCOH mixture. Aspen predicts the phase behavior of the water, MCH, and MCOH stream to form an extremely pure aqueous phase (>99.999 mol% water) and an organic phase containing 93.4 mol% MCH, 1.1 mol% MCOH, and 5.4 mol% water. This organic phase was sent to column C-4, where the distillate was collected and returned to be phase split in the decanter; the bottoms stream, comprising 98.7 mol% MCH and 1.2 mol% MCOH, was recycled back to the feed of column C-2.

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Recycle of Methyl-Cyclohexanone. Stream 22 could potentially be recycled back to reactor R-2; however, this would require an additional distillation column to remove MCOH byproduct. This additional distillation column would increase TCI by $3.9 million and lower utility costs negligibly ($1.12 h-1). A simpler approach selected in this design was to recycle stream 22 to distillation column C-2, where MCOH could be separated with existing process equipment. MCH was then purified and sent to reactor R-2 to complete the recycle stream. Recovery of 4-Methyl-6-Hydroxyhexanoic Acid.

As mentioned in the introduction, the

reconversion of the byproduct 4-M-6-HHA back to MCL was not included in the designed process depicted in Figure 2. Processes proposed for the industrial reconversion of 6-hydroxyhexanoic acid to ε-caprolactone require multiple reaction and separation steps, primarily the esterification and subsequent cyclization of the hydroxyacid[46,53], as well as long residence times which limit the recovered yield of ε-caprolactone to below 80%. Lab-scale synthetic routes towards this reconversion (primarily through Lewis acid catalysis [65,66]) also require long reaction times and would necessitate multiple separation steps. Given the base case performance of the BVO reaction, full recovery of the MCL from the 4-M-6-HHA byproduct could at most increase the total molar yield of MCL to 95.1%. Sensitivity Analysis. A sensitivity analysis was conducted on the base case process design of 10 kTons yr-1 to determine which variables have the largest effect on process economics. Individual process and economic variables were varied by ±10% of their base case value (with all other variables held constant), and a new minimum selling price of methyl-ε-caprolactone was calculated. Changes to the base case MSP from this sensitivity analysis are depicted in Figure 4. The overall molar yield of MCL from pcresol was found to have the largest impact on MSP. When the overall molar yield was varied over a range from 77.1% to 94.3%, the MSP of MCL varied from $3.912 kg-1 to $3.201 kg-1. Variation of ±10% of the following process parameters had comparable impact on the MSP: capital costs (48.3-59.1 $MM), minimum annual rate of return (13.5-16.5%), cresol purchase price ($0.90-1.10 kg-1), plant size (9.0-11.0 kTon yr-1), and operating costs (6.03-7.37 $MM yr-1). Less sensitive parameters were the tax rate (31.538.5%) and hydrogen peroxide price ($0.72-0.88 kg-1), while the palladium and hydrogen price variations had negligible economic impact. The clear conclusion from process sensitivity is the importance of catalytic selectivity to caprolactone via BVO in reactor R-2; improved yield significantly impacts the product sale price. Oxidation Reactor R-2 Performance. In the proposed process at 10 kTon yr-1 p-cresol processed, as presented in Figure 2, the maximum molar yield of MCL was 87.3% (the product of 97% selectivity to MCH in R-1 and 90% total molar conversion of MCH in reactor R-2). The base case design had an overall MCL molar yield of 85.6% based on catalytic performance of both reactors R-1 and R-2. To further demonstrate the impact of catalyst performance in reactor R-2, the minimum selling price of MCL was

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calculated as a function of the single-pass conversion and selectivity to MCL in reactor R-2 as depicted in Figure 5. Increasing the single-pass conversion of reactor R-2 results in a lower recycle flow rate to achieve an equivalent total molar conversion which leads to lower molar flow rates through all of the following process blocks—including columns C-2, C-3, and C-5; the resulting decreased capital and operating costs of these columns can reduce the MSP of MCL by almost $1 kg-1 between 10% and 50% single-pass conversion of methyl-cyclohexanone. Similarly, improved selectivity of the oxidation catalyst from 80% to >99% has the potential to improve the MSP of MCL by more than $1 kg-1. Economic Impact of Plant Size. The size of the chemical plant was evaluated for technical design and economic potential. For a range of 5 to 60 kTon yr-1, the same combination of unit operations outlined in Figure 2 was maintained with varying size as required for greater throughput. The initial base case design of 10 kTon yr-1 of p-cresol feedstock is identified in Figure 6 with the red hollow square data point, with smaller and larger chemical plants identified as black squares. The base case MSP of $3.521 kg-1 for MCL at 10 kTon yr-1 increases significantly to ~$4.50 kg-1 ($2.045 lbm-1) at the smaller plant size of 5 kTon yr-1. Alternatively, if plant size were increased to 60 kTon yr-1, the MSP of MCL would decrease to $2.557 kg-1 ($1.162 lbm-1). This production rate would account for ~10% of the global SBC thermoplastic elastomer market or approximately 5% of the global thermoplastic elastomer market. Increased plant sizes require a larger total capital investment and higher investor risk, as detailed in Figure S4. Alternative Feedstocks. The implementation of the two-stage hydrogenation/oxidation process with p-cresol is one potential process iteration, but there exists a broader range of potential alkyl-phenol mixtures that can serve as feedstock to manufacture alkyl-caprolactone monomers. Isomers of p-cresol (o,m-cresol, or mixtures of these components) will undergo almost identical processing, with only minor differences in separation and catalysis. Sels and co-workers have also proposed the extension of this chemistry to other lignin-derived monomers which contain larger alkyl groups (e.g., propyl) and/or additional oxygenated functionality (e.g., ethers, multiple hydroxyl groups)[36]. The structure and size of the alkyl group on branched caprolactone monomer allows for alternative poly(alkyl-caprolactone) materials with properties tunable by the size and structure of the alkyl group. Additionally, the higher oxygen content in lignin-derived monomers introduces additional requirements for hydrogenation in the initial reactor, R-1[34]. Despite these challenges, extending the proposed process of Figure 2 to additional alkyl-phenols from lignin could provide a biorenewable pathway to alkyl-caprolactone monomers with economic potential.

Conclusion. The chemical manufacturing of methyl-ε-caprolactone was evaluated for technical process design and economic analysis by the combination of hydrogenation of p-cresol to methyl-cyclohexanone followed by Baeyer-Villiger oxidation. A two-stage process was proposed based on the two chemistries

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with accompanying separations to produce methyl-ε-caprolactone at 99.9% purity.

Processes were

proposed at three scales of 10, 30, and 60 kTon yr-1 of p-cresol feed, and the associated minimum selling price of product methyl-ε-caprolactone for a break-even process was $3.521, $2.798, and $2.557 kg-1, respectively, based on a feedstock price of $1.00 kg-1 of p-cresol. The overall yield of the optimized basecase process was 85.7%, and sensitivity analysis of the major process variables identified improvements in overall selectivity via catalytic performance of the Baeyer-Villiger oxidation catalyst as the key opportunity for improving the overall economics of the process.

Supporting Information. The supporting information is available free of charge on the ACS Publications website at DOI: xxxx Process flow diagram stream flow rates, economic analysis parameters, sizing and costing of capital equipment, reactor design, process sensitivity to economic parameters, and process optimization details. Acknowledgements. This work was supported by the National Science Foundation through the University of Minnesota Center for Sustainable Polymers under award number CHE-1413862. Additional financial support was provided by the Minnesota Corn Growers Association.

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Figure 1. Process chemistries for the conversion of p-cresol to 4-methyl-ε-caprolactone

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Figure 2. Process flow diagram for the conversion of p-cresol to 4-methyl-ε-caprolactone via hydrogenation (B-1, R-1 & F-1), ketone purification (C-1 & C-2), ketone oxidation (R-2, D-1 & C-4), and caprolactone purification (C-3 & C-5).

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Figure 3. Economic contribution to minimum selling price (MSP) [$ kg-1 of product methyl-ε-caprolactone] of process stages based on capital and operating costs. Assumptions include a p-Cresol purchase price of $1 kg-1, a plant size being fed 10,000 ton yr-1 of p-cresol, a MARR of 15%, a tax rate of 35%, 30 year project lifetime, and total molar yield of p-cresol to MCL of 85.6%.

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Figure 4. Sensitivity of base case (10,000 ton yr-1) minimum selling price (MSP) of methyl-εcaprolactone (MCL) to 10% changes in operational/economic parameters.

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Figure 5. Oxidation Reactor, R-2, Economic Performance. Minimum selling price (MSP) of methylε-caprolactone (MCL) as a function of the selectivity to MCL and single-pass conversion in oxidation reactor, R-2, which converts methyl-cyclohexanone by Baeyer-Villiger oxidation to MCL for the base case process (10,000 ton yr-1).

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Figure 6. Minimum selling price (MSP) of methyl-ε-caprolactone (MCL) relative to chemical plant size based on the processing rate of feedstock p-cresol (in metric kilotons per year).

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Table 1. Process Feedstock and Utility Prices p-Cresol Purchase Price [$ kg-1]67 1.00 -1 49 Hydrogen Purchase Price [$ kg ] 1.57 Hydrogen Peroxide Solution Purchase Price [$ kg-1]68 0.8 -1 69 Palladium Price [$ kg ] 32,000 -1 a Cooling Water [$ kg ] 3.08 x 10-5 -1 a Steam, 100 PSI [$ kg ] 0.0179 -1 a Steam, 165 PSI [$ kg ] 0.0215 Steam, 400 PSI [$ kg-1]a 0.0258 -1 a Electricity [$ kWh ] 0.0775 Wastewater Treatment [$ kg-1 organic removed]59 0.33 a Estimated using Aspen Process Economic Analyzer (V8.4 Aspen Technology).

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Table 2. Capital Cost in Millions of U.S. Dollars for a Chemical Process of Varying Plant Size (10, 30, 60 kTon yr-1) Process Block Cresol hydrogenation MCH Purification MCH Oxidation + Ketone Recycle MCL Purification Total installed equipment cost Total Capital Investment

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10 kTon yr-1 3.4 13.0 2.3 7.7 25.0 53.7

Capital Cost ($MM) 30 kTon yr-1 9.0 23.1 4.2 13.5 49.8 91.8


60 kTon yr-1 17.0 37.1 5.1 20.9 80.1 140.9

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Table 3. Process Operating Costs with Varying Plant Size (10, 30, 60 kTon yr-1) Raw Material Operating Cost ($MM) 10 kTon yr-1 30 kTon yr-1 p-Cresol 10.0 30.0 H2 0.6 1.8 H2O2 Solution 2.5 7.5 Steam 2.1 6.4 Cooling water 0.1 0.2 Electricity 0.1 0.1 Fixed Costs 3.7 6.0 Total Operating Cost 19.1 52.0 Note: Wastewater treatment is a negligible cost.

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60 kTon yr-1 60.0 3.6 15.0 13.9 0.4 0.2 7.7 100.8

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For Table of Contents Use Only

Synopsis: Process design for the renewable production of methyl-caprolactone reveals economic viability of new biodegradable materials with advanced properties.

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Figure 1 170x242mm (300 x 300 DPI)


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Figure 2 414x232mm (300 x 300 DPI)



Figure 3 304x143mm (300 x 300 DPI)


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Figure 4 341x156mm (300 x 300 DPI)



Figure 5 395x303mm (300 x 300 DPI)


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Figure 6 278x232mm (300 x 300 DPI)



Graphical Abstract 203x83mm (300 x 300 DPI)


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Techno-Economic Analysis of a Chemical Process to Manufacture

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