Dimethyl Carbonate

Dimethyl Carbonate

16 DIMETHYL CARBONATE CHAPTER OUTLINE 16.1 Introduction 649 16.2 Direct Conversion of CO2 to Dimethyl Carbonate 16.2.1 Process Design 653 16.2.2 Proce...

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16 DIMETHYL CARBONATE CHAPTER OUTLINE 16.1 Introduction 649 16.2 Direct Conversion of CO2 to Dimethyl Carbonate 16.2.1 Process Design 653 16.2.2 Process Modeling and Simulation 655


VLE Data 655 Membrane Reactor 655 Separation Section 659 Flash Recovery and Predistillation Column (Dist-01) 659 Extractive Distillation (Dist-02) and Solvent Recovery (Dist-03) 661 Methanol Recovery (Dist-04) 662

16.2.3 Process Performance Evaluation 662 16.2.4 Economic Evaluation 665 16.2.5 Key Performance Indicators 668 16.3 DMC Synthesis by Propylene Carbonate Transesterification 16.3.1 Thermodynamics 670 16.3.2 Kinetics 671 16.3.3 PC Synthesis Process 672 16.3.4 DMC Synthesis Process 674 16.3.5 Key Performance Indicators 676 16.4 Conclusions 677 References 678


16.1 Introduction In the search for “green chemistry,” dimethyl carbonate (ðCH3 OÞ2CO, DMC) received interest because of its low toxicity and fast biodegradability (Delledonne et al., 2001; Keller et al., 2010; Ono, 1997; Pacheco and Marshall, 1997; Santos et al., 2014). DMC is considered as an important building block for different organic syntheses with major applications in the production of polycarbonates, as methylation agent, in use as solvent for lithium ion batteries and as fuel additive (Delledonne et al., 2001; Keller et al., 2010; Ono, 1997; Pacheco and Marshall, 1997; Santos et al., 2014) Applications in Design and Simulation of Sustainable Chemical Processes. https://doi.org/10.1016/B978-0-444-63876-2.00016-4 Copyright © 2019 Elsevier B.V. All rights reserved.




with a worldwide production capacity of 45 kt/a in 1997 (Pacheco and Marshall, 1997; Santos et al., 2014). Because of its environmental-friendly nature, it can replace hazardous chemicals such as phosgene and dimethyl sulfate (Santos et al., 2014). For the future, DMC is also considered as a potential fuel additive due to its high blending octane number, reduced CO and NOx emission, and high oxygen content (Pacheco and Marshall, 1997; Santos et al., 2014). It is estimated that this will result in a worldwide demand of 450 kt/a (Keller et al., 2010; Pacheco and Marshall, 1997). In 2018, the global DMC market was valued at 380.4 million USD (https://market.us/report/dimethyl-carbonate-market/). The traditional method for DMC production, in use until the end of the 20th century, was the pyridine-catalyzed phosgenation of methanol, according to the following global reaction: 2CH3 OH þ COCl2 /ðCH3 OÞ2CO þ 2HCl


The main drawbacks of the process are the use of toxic phosgene and the generation of large amounts of chlorhydric acid. Starting with the 1980s, the oxidative carbonylation of methanol became the preferred route (Wang et al., 2016). The raw materials can be obtained from coal, natural gas, or biomass. CO þ 0:5O2 þ 2CH3 OH/ðCH3 OÞ2CO þ H2 O


The Italian company EniChem developed an industrial process. The reaction takes place in liquid phase, at 10e30 bar and 100e130 C, using CuCl as catalyst. The reaction is exothermic, the heat of reaction being removed by evaporation of methanol (which also acts as a solvent). The main drawback of the process is the large amount of low-quality heat (around 100e130 C) which is lost. Additionally, the water produced in the reaction leads to catalyst deactivation and decomposes the DMC product. Dow Chemical patented a gas-phase process, where the reaction takes place in the presence of copper-based catalysts (CuCl2, or copper methoxychloride-based pyridine complex) on activated carbon, and KCl or MgCl2 as cocatalyst (Curnutt, 1991). The process is not applied on large scale because the selectivity and yield are quite low, and toxic CO is still used. In 1992, Ube Industries (https://www.ube.com) announced a new DMC process, where carbon monoxide and methyl nitrite react to form DMC and nitric oxide, on a PdCl2/C catalyst. No water is formed during this step; therefore, the stability of the catalyst is improved. Moreover, the problems related to separation of the methanoleDMC azeotrope are avoided. 2CH3 ONO þ CO/ðCH3 OÞ2CO þ 2NO



Nitric oxide is then oxidized back, with oxygen and methanol, to methyl nitrite and water in a separate nitrite regenerator. The reaction is fast (up to 2 s residence time) and takes place in liquid phase, without any catalyst. 2NO þ 0:5O2 þ 2CH3 OH/2CH3 ONO þ H2 O


The total reaction is the oxidation of carbon monoxide and methanol to DMC and water with dimethyloxyde being the major by-product. However, NO and RONO are toxic and corrosive. Urea alcoholysis is a promising method because it uses low-cost materials. The reaction takes place in two steps (at 100 and 180 C), via methyl carbamate. The separation and purification of DMC is easy. ðNH2 Þ2CO þ 2CH3 OH/ðCH3 OÞCO þ 2NH3


However, the reaction is not thermodynamically favorable. Thus, to be economically feasible, it requires an LHSV (liquid hourly space velocity) of 0.8, while 0.02 is the best result to date (Pacheco and Marshall, 1997). However, indirect urea alcoholysis (via ethylene- or propylene carbonate [PC]) appears to be a feasible process (Wang et al., 2016). Direct synthesis (reaction (16.6)) is the most attractive route (Delledonne et al., 2001; Keller et al., 2010), because of its simplicity and the use of cheap raw materials. 2CH3 OH þ CO2 /ðCH3 OÞ2CO þ H2 O


However, the process is not applied in industry due to severe equilibrium limitations arising from the stability of the CO2 molecule. This process will be examined in more detail in the next section. Synthesis of DMC by transesterification (Fig. 16.1) is based on the reaction of epoxides (or glycols) with CO2, leading to cyclic carbonates, which further react with methanol, to produce DMC in a transesterification reaction. This process will be studied in the second part of this chapter.

Figure 16.1 Synthesis of dimethyl carbonate by transesterification of propylene carbonate with methanol.




16.2 Direct Conversion of CO2 to Dimethyl Carbonate The use of CO2 as building block for DMC production (16.6) is economically and environmentally interesting. It reduces the CO2 emission into the atmosphere and decreases the demand for fossil fuels, and because CO2 is considered as waste product, it is still an inexpensive feedstock (von der Assen et al., 2013) useable for producing valuable chemicals, such as methanol (Kiss et al., 2016). At ambient conditions, the equilibrium constant of reaction (16.6) is in the order of 105. Starting with a mixture methanol/ CO2 in the ratio 10/1, at 50 atm and 50 C the CO2 equilibrium conversion is about 5% (Kuenen et al., 2016). The reaction is exothermic; therefore, the equilibrium conversion drops quickly at higher temperatures. The equilibrium conversion could be increased by performing the reaction at much higher pressures. The selective removal of reaction products (e.g., by the use of dehydrating agents) can also shift the equilibrium to the product side, resulting in higher conversions. Sakakura and Kohno (2009) categorized two types of dehydrating agents: recyclable agents, such as acetals, butylene oxide, or molecular sieves, and difficult-to-recycle agents, such as orthoesters or dicyclohexyl carbodiimide. Because of high costs and the difficulty to regenerate this last group, acetals are preferred. However, at higher acetal concentrations, DMC yield is suppressed due to the formation of side products (Tomishige and Kunimori, 2002). A sustainable, process intensification alternative is to perform the reaction in a membrane reactor. By combining the reaction with a membrane separation, one can instantly and continuously remove one or more products from the reaction mixture. Li and Zhong (2003) investigated three types of membranes in a membrane reactor and observed an improved methanol conversion compared with the use of a conventional catalytic reactor. The use of more selective membranes (Choi et al., 2002; Li and Zhong, 2003), an increase in pressure (Li and Zhong, 2003), and a better integration of membrane and reactor (Mengers et al., 2014) can further enhance the conversion. Another challenge is the formation of DMC/methanol and DMC/water azeotropes (Luo et al., 2000; Camy et al., 2003). This requires nonstandard separation techniques to obtain pure DMC, such as membranes (Won et al., 2002), pressure-swing distillation (PSD) (Wang et al., 2010), or extractive distillation (ED) (Wang et al., 2010). Won et al. (2002) showed that pervaporation


using chitosan membranes circumvents DMC/methanol azeotrope formation. Wang et al. (2010) compared the performance of PSD with ED. Although both technologies can break the DMC/methanol azeotrope, ED uses 71% less energy compared with PSD. The challenges discussed above set the boundaries for commercial operation. Therefore, this section focuses on the design of a DMC process plant for the direct conversion of CO2 and methanol to DMC. The focus is on the technoeconomical evaluation of the overall process, showing the implications of the technical limitations on the process economics. This is achieved by combining reasonable simple modeling of the membrane reactor with Aspen Plus simulations. Furthermore, the influence of the most important parameters is analyzed, to be considered in future investigations for enhancing the profitability of the process.

16.2.1 Process Design We consider the direct conversion of CO2 and methanol to DMC with a DMC production capacity of 20 ktonne/year and >99wt% purity. The production capacity is an average size of a DMC production facility (Buysch, 2000) and the purity is a commonly used industrial grade. The process for conversion of CO2 to DMC is presented in Fig. 16.2 (Kuenen et al., 2016). The feed (stream #1, consisting of methanol and CO2) is mixed with the vapor and liquid recycles to obtain a methanol: CO2 feed ratio of 4:1. This stream enters the membrane reactor at 50 atm and 135 C and the reactants are converted with 100% selectivity to DMC. Water, methanol, and traces of CO2 and DMC selectively permeate through the membrane. The product stream (retentated#2) with DMC leaves the reactor and enters a flash drum where the gaseous components (CO2) are removed from the liquid components (DMC, water, and methanol). To assure that only CO2 is in the gaseous phase, the separation is performed at 25 C. Meanwhile, the pressure in the flash drum decreases to 1.5 atm. The obtained CO2 is compressed to 50 atm and recycled to the membrane reactor. The liquid product stream leaving the flash drum (#4) is distilled in the predistillation column (Dist-01). Water and most of the methanol are removed (#8) and the DMC product stream (#7) is concentrated. The product stream leaves the condenser of Dist-01 at the top as liquid with DMC/methanol in an



Figure 16.2 Schematic overview of the process design for the direct conversion of CO2 to dimethyl carbonate. Compressors (C-0x) are assumed to be multistage compressors with internal cooling.


azeotropic ratio, and CO2 leaves the top of the column as vapor (#6). Water and most of the methanol are removed via the bottom of the column. To minimize the loss, the CO2 stream and water/ methanol stream are both recycled back to the membrane reactor. The DMC/methanol mixture (#7) is separated in an ED column (Dist-02) by dissolving DMC in an extractant. DMC is obtained together with the extractant at the bottom (#12) of the column. Last traces of CO2 (vapor phase) and most methanol (liquid phase) are obtained at the top and recycled back to the membrane reactor. Finally, in the third distillation column (Dist-03), DMC is recovered from the extractant and obtained at the top (#13) at a purity of >99wt.%. The recovered extractant is obtained as the bottom product (#14) and recycled back to Dist-02. Any minor extractant losses are compensated for by the supply of fresh extractant to the ED column (#9). In the membrane reactor, water, methanol, and traces of CO2 and DMC permeate through the membrane. To minimize the methanol loss, this stream is recycled back to the reactor, after removal of water in the fourth distillation column (Dist-04). To prevent accumulation of extractant, a purge is applied at the bottom product stream of Dist-01.

16.2.2 Process Modeling and Simulation VLE Data DMC forms azeotropes with both methanol (Luo et al., 2000) and water (Camy et al., 2003). The SoaveeRedlicheKwong (SRK) property model was used to describe the VLE data available in literature for DMC/methanol (Luo et al., 2000) and DMC/water mixtures (Camy et al., 2003). All the binary interaction parameters related to the SRK property model are available in the pure components’ databank of the Aspen Plus. Fig. 16.3 compares the predictions of the model with the experimental data (Kuenen et al., 2016). Membrane Reactor The simulation of the process was performed in Aspen Plus v8.0. Because Aspen Plus does not include a membrane reactor, this is simulated as a cascade of reactor (REQUIL)dseparator (SEP2) combinations as presented in Fig. 16.4 (Kuenen et al., 2016). In the reactor, the reaction proceeds to equilibrium. In the separator, the components are separated. The product outflow of the separator enters the reactor of the next block, followed by the next separator, etc.




Figure 16.3 Experimental and simulated vaporeliquid equilibrium (SRK model), 1 atm. Left: dimethyl carbonate (DMC)/methanol; Right: DMC/water.

Figure 16.4 Schematic representation of the membrane reactor simulation and the key input and output parameters.

Each reactor block is represented by the REQUIL model implemented in Aspen Plus. This requires an expression for calculating the equilibrium constant defined by (16.7). Keq ¼

½H2 O½DMC ½CO2 ½MeOH



with Keq ¼ 6.12e0.016 T (Kuenen et al., 2016). The separation of water from the reaction mixture requires membranes with high water/gas selectivity. Polymer membranes based on sulfonated poly (ether ether ketone) (SPEEK) showed excellent dehydrating performances (Sijbesma et al., 2008). For the direct conversion of CO2 and methanol toward DMC, SPEEK is capable to selectively remove water vapor while retaining CO2


and methanol (Khan et al., 2011; Sijbesma et al., 2008; Silva et al., 2005). This prevents the equilibrium to establish and favors the DMC formation. The separator models assume that the SPEEK membrane removes 98% of the produced water. To account for the removal of other components, membrane selectivity toward methanol, CO2, and DMC are included as well. A H2O/methanol selectivity of 25 is assumed and a conservative value of 1000 is selected for H2O/CO2 selectivity. The H2O/DMC selectivity was considered to be 250. More details regarding the various literature data which support these choices can be found in Kuenen et al. (2016). Recently, membrane technologies for the separation of methanol and DMC were also reported by Holtbruegge et al. (2013, 2014). These new membranes have a slightly better separation performance above 30wt% MeOH (although the interesting region below 20wt% MeOH was not investigated), while the total flux is in the same order of magnitude as the ones reported by Won et al. (2002). The percentage of water permeating through the membrane per cascade ðxH2 O  permeateÞn can be determined using the following equation:   FH2 O$ fH2 Oproduced n  (16.8) ðxH2 O  permeateÞn ¼  fH2 Oreactor out n   with FH2 O is the removal fraction of water, fH2 Oproduced n is the amount in the reactor of cascade n [kmol/hr],   of water produced and fH2 Oreactor out n is the total amount of water leaving the reactor and entering the separation unit of cascade n [kmol/hr]. This stream also contains water that was not removed in the previous cascades. To determine the fraction of DMC permeating through the membrane, (xDMCpermeate)n, the removal fraction of H2O ðFH2 O is divided by the H2O/DMC selectivity SH2 O=DMC. FH 2 O ðxDMCpermeate Þn ¼

  $ fH2 Oproduced n

SH2 O=DMC ðfDMCreactor out Þn


Because the reactants are already present before entering the membrane reactor, their overall permeation (not per cascade, but over the complete reactor) is determined: " #! fCO2 FH2 O xreactantspermeate ¼ 1  nj zCO2 (16.10) SH2 O=j fj




here nj is the stoichiometric ratio of the reaction [e], zCO2 is the conversion of CO2 [], fXO2 is the amount of CO2 fed to the reactor [kmol/hr], and fj is the amount of j fed to the reactor [kmol/hr].  When j is CO2, then fCO2 fj equals 1, while if j is methanol, then fCO2 fj equals 1/4. Finally, to determine the amount of reactants that permeates through the membrane per cascade ((xreactantpermeate)n), we calculate the fraction of water that is removed per cascade relative to the total amount of water that is removed in the membrane reactor module and multiply this with xreactantspermeate. Fig. 16.5 (left) shows the calculated  CO2 conversion as function of the water removal fraction ðFH2 O for different methanol/CO2 feed ratios (Kuenen et al., 2016). A strong increase in CO2 conversion especially at higher water removal fractions is observed. This effect is more pronounced at higher methanol/CO2 feed ratio. Both effects stem from the equilibrium constant (Eq. 16.7), which shows a quadratic dependency on the methanol concentration. Fig. 16.5 (left) suggests that it is beneficial to operate the process at higher methanol/CO2 feed ratios. However, an excess of methanol decreases the DMC concentration in the reactor effluent due to dilution. This will lead to higher energy requirement for product purification. To find the optimum feed ratio, we consider the DMC concentration at the outlet of the membrane. Fig. 16.5 (right) shows the DMC concentration in the retentate as function of the methanol/CO2 feed ratio (Kuenen et al., 2016). The maximum DMC concentration is obtained at a feed ratio of approximately 3.5e4. Below this feed ratio, the conversion is low, while above this ratio the dilution by methanol is dominant. Consequently, we assume a feed ratio of 4 in the simulations, as this gives the highest DMC

Figure 16.5 Left: CO2 conversion as a function of the water removal fraction for different methanol/CO2 feed ratios; Right: dimethyl carbonate concentration in retentate as function of the methanol/CO2 feed ratio.


Table 16.1 Summary of the Membrane Reactor Specifications. Parameter

MR Conditions

Treactor Preactor Methanol/CO2 feed ratio  Removal fraction H2O ðFH2 O SH2 O=DMC SH2 O=MeOH SH2 O=CO2

135 C 50 atm 4 98% 250 25 1000



e e e e Chitosan SPEEK SPEEK

Aresta et al. (2008) Aresta et al. (2008) e e Won et al. (2002) Silva et al. (2005) Sijbesma et al. (2008), Khan et al. (2011)

concentration. Table 16.1 summarizes the selected operating conditions of the membrane reactor module as discussed earlier (Kuenen et al., 2016). Separation Section The separation train following the reactor was selected based on the specific properties of the MeOH/DMC system. Several technologies were evaluated such as membranes, PSD, extraction, and ED. Because of limited literature information and the high-purity specification for DMC (99wt%), membranes are discarded. For PSD, the effect of pressure on the azeotropic composition of MeOH-DMC was evaluated and the results showed only a small change in composition (7% change for a DP of 4 bar) and a boiling point difference between the azeotrope and pure methanol of only 1 C, which makes PSD very energy intensive. Extraction is described in literature but for a specific situation not applicable for the DMC process proposed here. ED was finally selected because literature reports several entrainers specific for the MeOH/DMC separation. The most promising entrainers were investigated and finally methyl isobutyl ketone (MIBK) was selected. Flash Recovery and Predistillation Column (Dist-01) After the membrane reactor, a simple gaseliquid flash separation is performed to separate the CO2 (gas) from the other components. The liquid stream outlet from the flash has to be predistilled due to the low DMC concentration attainable at the outlet of the




membrane reactor. By predistillation, the excess of methanol and remaining water are separated from DMC. The concentrated DMC stream is further concentrated by ED. The top distillate of Dist-01 has a near-azeotropic composition of DMC/methanol. The fraction of methanol in the azeotrope increases with pressure; thus, to concentrate DMC, Dist-01 should operate at low pressure. The column pressure is set at 1.5 atm to maintain a certain driving force over the process (>1 atm) while minimizing the amount of methanol leaving together with DMC. Before the flash vessel, the reactor effluent is cooled from 135 to 25 C, as a further decrease in temperature leads to an economic penalty due to cooling agent requirements. At higher temperatures, more DMC is lost in the vapor phase and recycled together with CO2 to the membrane reactor, thus being detrimental to the equilibrium reaction. Optimally, the flash vessel operates at low pressures to achieve high CO2 recovery, but lower pressures also increase the amount of DMC lost with CO2 into the vapor phase. However, at higher pressures, more CO2 enters Dist-01 and leaves the column as vapor distillate, taking also a part of DMC with it. Fig. 16.6 (left) shows the amount of DMC recycled via the flash drum and Dist-01 as a function of the operating pressure of the flash vessel as determined by simulations, considering a fixed pressure of 1.5 atm in Dist-01 (Kuenen et al., 2016). Clearly, a significant amount of DMC is recycled via the flash vessel and Dist-01. To minimize the loss, the optimal flash vessel pressure is 1.3e1.5 atm. As Dist-01 operates at 1.5 atm, the flash vessel is also set to this pressure.

Figure 16.6 Left: Percentage of dimethyl carbonate (DMC) recycled in the flash drum and Dist-01 as a function of the flash vessel pressure (calculated by simulations at a fixed column pressure of 1.5 atm in Dist-01); Right: DMC recovery after the solvent recovery step from the extractive distillation as function of the MIBK/feed ratio, as determined by Aspen Plus simulations.

Chapter 16 DIMETHYL CARBONATE Extractive Distillation (Dist-02) and Solvent Recovery (Dist-03) The concentrated DMC product stream from Dist-01 enters Dist-02, which is the ED column. Because of the formation of an azeotrope between DMC and methanol, it is not possible to obtain pure DMC by conventional distillation. Several technologies are available to purify DMC, such as membranes (Won et al., 2002), PSD (Wang et al., 2010), and ED (Wang et al., 2010). Because of the limited information available regarding membranes for DMC separation, this separation technology was discarded. Wang et al. (2010) compared the use of PSD and ED to break the DMC/methanol azeotrope and showed that ED uses 71% less energy as compared with PSD, thus making it an obvious choice to use for DMC/methanol. Several entrainers are available for ED, e.g., dimethyl oxalate (Pacheco and Marshall, 1997), phenol (Wang et al., 2010), 2-ethoxyethanol, and 4-methyl2-pentanone (methyl isobutyl ketone, MIBK) (Matsuda et al., 2011). The last two were investigated in more detail by Matsuda et al. (2011) who observed that MIBK is more selective to DMC. However, the drawback of MIBK is that it is more difficult to separate MIBK again from DMC, as compared with 2-ethoxyethanol. Nevertheless, considering that a less expensive entrainer is preferred over a lower reboiler duty in the solvent recovery step, MIBK was selected as entrainer for the ED. The ED column was optimized by determining the DMC recovery in Dist-02 as function of the MIBK/feed ratio, as shown in Fig. 16.6 (right). We selected an MIBK/feed ratio of 0.8, leading to a recovery of w99.5% DMC. At lower solvent to feed ratios, a relatively large amount of DMC is lost, while at higher ratios an excessive amount of MIBK is needed to recover the last DMC tracesdleading to an economic penalty in the solvent recovery column. The energy efficiency can be improved by heat integration of the condenser of Dist-02 with the reboiler of Dist-01. Because a minimum temperature difference of 15e20K is required to drive heat transfer (Seider et al., 1999; Turton et al., 2003) and considering that the reboiler of Dist-01 operates at 76 C, the temperature of the condenser of Dist-02 is set to 100 C, resulting in a corresponding operating pressure of 3.85 atm. This temperature is equal to the boiling point of the mixture containing mainly methanol. At these conditions, most methanol is recovered in the vapor phase of the distillate along with the last traces of CO2, which is undesired. By increasing the column pressure, the boiling point of methanol increases, and more methanol is retained in the liquid distillate,




while less methanol is lost in the vapor distillate. Hence, the column pressure in Dist-02 is set at 7 atm, after which a further increase in pressure has only minimum effect on the recovery. In the solvent recovery column (Dist-03), DMC is separated from MIBK, which is recycled back to Dist-02. A conventional distillation column is used, being designed with the specification that >99wt% DMC is recovered at a purity of >99wt% DMC. Higher DMC recoveries result in higher energy requirements in Dist-02, as it becomes practically impossible then to prevent methanol from going along with DMC and MIBK to Dist-03. Again, the condenser of Dist-03 is heat integrated with the reboiler of Dist-01 to increase the energy efficiency. The required condenser temperature in Dist-03 is minimum 100 C, giving a temperature difference of 15e20 C (Seider et al., 1999; Turton et al., 2003). This is reached when the pressure of Dist-03 column is set at 1.5 atm. Methanol Recovery (Dist-04) The permeate flow of the membrane reactor consists mainly of water and methanol with traces of DMC and CO2. To recover the products, this stream is separated in a distillation column, which is, based on the relative volatility of methanol/water and the feed composition, the standard technology to separate methanol from water. Alternative separation technologies (e.g., molecular sieves) become economically interesting when the water concentration in the feed is below 15wt% (Wintek, 2014). The distillation column Dist-04 operates at 1 atm to achieve >99wt% methanol recovery with a purity >98.5wt% at the top of the column and a water recovery of >99.5wt%. The methanol recovery at the top and the water recovery at the bottom of Dist-04 are based on the optimum energy efficiency in Dist-04. Higher recoveries require much higher energy inputs, which is economically not interesting. Except for some small traces of water, the other 1.5wt% is CO2 and DMC that are also removed at the top of Dist-04. Table 16.2 provides a complete overview of the different input variables for the separation units (Kuenen et al., 2016). These input data are set in Aspen Plus and the process is optimized to produce 20 ktonne/a of DMC with a purity of 99wt%.

16.2.3 Process Performance Evaluation Table 16.3 provides the overall mass and energy balance of the new DMC process (Kuenen et al., 2016). Notably, the membrane reactor module reaches a CO2 conversion of 6.3% toward DMC.



Table 16.2 Input Variables for the Flash Drum and the Distillation Columns. Unit



Flash drum Dist-01 Dist-02

P P Tcondenser P MIBK/Feed ratio Tcondenser P P

1.5 atm 1.5 atm >100 C 7.0 atm 0.8 mol/mol

Dist-03 Dist-04

>100 C 1.5 atm 1.0 atm

Table 16.3 Mass Balance of the DMC Process, Calculated by Aspen Plus Simulations. DMC (mol%) Water (mol%) CO2 (mol%) Methanol (mol%) MIBK (mol%) F (kmol/hr) T ( C) P (atm)









#9 #10 #11


#13 #14


e e 32.4 67.6

1.5 0.0 19.8 78.6

0.2 29.5 0.4 69.9

1.7 0.0 5.5 92.7

0.4 0.0 97.2 2.4

1.9 0.0 91.4 6.7

11.5 0.0 4.6 83.8

0.0 0.0 0.0 99.8

e e e e

0.1 0.0 48.1 51.8

0.1 0.0 2.5 97.4

12.6 0.0 0.0 0.3

97.7 0.0 0.0 2.1

0.1 0.0 0.0 0.0

0.2 0.2 0.6 99.0

e 84.9 135 50

0.1 2071.5 135 50

0.0 93.4 58 1

0.1 1749.3 3.4 1.5

0.0 322.2 3.4 1.5

0.0 93.1 20 1.5

0.0 239.3 20 1.5

0.1 1416.9 75.6 1.5

100 0.05 90 7

0.0 12.7 100 7

0.0 198.6 100 7

87.1 219.8 192.6 7

0.2 28.1 102.9 1.5

99.9 191.6 131 1.5

0.0 65.5 52 1

Overall, 84.9 kmol/h of reactants enter the process in a ratio of 2.1:1 methanol/CO2. This ratio is higher than the stoichiometric requirements and is caused by the methanol loss via the product stream and the purge, while all CO2 is recycled. The retentate stream (#2), which contains only 1.5 mol% of DMC, is purified via different steps to eventually obtain 97.7 mol% DMC (equal to >99wt%) as main product (#13). The key reason for the low DMC concentration is the low conversion and the excess of methanol used in the



process. Because of low DMC concentrations, it is very energy intensive to purify the outlet stream to the desired DMC purity. As shown in Table 16.3 (Kuenen et al., 2016), a crucial part of the separation process occurs in Dist-01, where DMC is concentrated to DMC/methanol azeotropic composition. In this distillation column, only 14 mol% of the feed is ending up in the product stream (#7) and 81 mol% is leaving Dist-01 via the bottom of the column (#8) containing mainly methanol. In Dist-02, only 12 mol% of the feed (#7) leaves the distillation column as product stream (#12). As shown in Table 16.4 (Kuenen et al., 2016), the energy usage in these columns is extremely large. Dist-01 and Dist-02 use 70% of the total heat required in the overall process, in both cases mainly to separate methanol from DMC. Especially the large energy demand of Dist-01 is caused by the low DMC concentration in the feed stream to this column. Furthermore, due to the DMC/methanol azeotrope, a maximum DMC purity of 11.5 mol% is attainable in the liquid distillate (#7). The rest is mainly methanol that is removed over the top of the column along with DMC, according to the azeotropic composition. Evaporation of methanol causes the high energy demand in Dist-01. Cooling also demands a substantial amount of energy, and 74% of this energy is used to cool down the retentate stream (#2) from 135 to 25 C before the flash drum. This stream contains mainly unreacted CO2 and methanol and cooling prevents that DMC is lost in the vapor phase and recycled to the membrane reactor. In addition, the heating consumes a significant amount of energy: 79% of the heating duty being used to heat up the

Table 16.4 Overview of Energy Usage Through the Process, Calculated by Aspen Plus Simulations. Equipment Distillation columns

Pumps Compressors Heat exchangers

Energy [MW] Dist-01 Dist-02 Dist-03 Dist-04

Cooling Heating

14.12 5.78 2.23 2.61 0.16 1.56 10.63 3.82


recycled liquid stream from the bottom of Dist-01 to the desired membrane reactor temperature of 135 C. Overall, the process requires a high amount of energy compared with the amount of DMC produced, mainly due to the low conversion and the excess of methanol present in the process. The conversion in the membrane reactor depends not only on the temperature and pressure but alsoon the methanol/CO2 feed ratio, removal fraction of water ðFH2 O , and the selectivity of the membrane. The influence of the H2O/methanol membrane selectivity SH2 O=MeOH is limited, as increasing this selectivity from 5 to 1000 results in an increase of the conversion of only 0.76%, of which 0.61% is already gained by increasing the selectivity up to 25. Preventing the permeation of methanol causes the methanol concentration in the retentate to increase. In principle, this shifts the equilibrium toward the conversion of DMC, but the effect is minimal. Thus, it is of little importance to have highly selective membranes to improve the chemical conversion.

16.2.4 Economic Evaluation A cost analysis was performed to determine the profitability of the process, using the design specifications of the main process equipment given in Table 16.5 (Kuenen et al., 2016), in terms of specific equipment size (diameter, height, volume, power, transfer area, etc). Table 16.6 provides the investment costs of all major equipment units in the process (Kuenen et al., 2016). These investment costs were determined using the overall factor method (Seider et al., 1999). Price estimation of most equipment was done via the correlations in Peters et al. (2003) and the equipment sizing results obtained from Aspen Plus. These prices are corrected to the current prices using a Chemical Engineering Plant Cost Index (CEPCI) of 586 (2014) and multiplied with a Lang factor of 6 to take into account installation, piping, electrical, etc (Seider et al., 2010). According to the information in Table 16.6, a total investment of 45.5 M$ is needed to build this DMC plant. Notably, the total investment costs are relatively equally divided over the membrane reactor, the distillation columns, the compressors, and the heat exchangers. Table 16.7 gives an overview of the operating costs and the revenues for the production of 20 ktonne/year of DMC (Kuenen et al., 2016).




Table 16.5 Design Specification Process Equipment Throughout the Process. Equipment


Key Specifications

Height (m)

Diameter (m)

Distillation Columns

Dist-01 Dist-02 Dist-03 Dist-04 e C-01 C-02 P-01 P-02 P-03 P-04 P-05 Reboiler Dist-01 Reboiler Dist-02 Reboiler Dist-03 Reboiler Dist-04 Condenser Dist-01 Condenser Dist-02 Condenser Dist-03 Condenser Dist-04 He-01 He-02 He-03 He-04 He-05 He-06 He-07 He-08

50 79 67 49

27.5 42.0 36.0 27.0 10.6 e e e e e e e e e e e e e e e e e e e e e e e

3.16 2.92 2.00 1.58 3.32 e e e e e e e e e e e e e e e e e e e e e e e

Flash drum Compressors Pumps

Heat exchangers

stages stages stages stages

1530 kW 30 kW 4 kW 133 kW 28 kW 8 kW 13 kW 471 m2 279 m2 104 m2 142 m2 1861 m2 175 m2 113 m2 326 m2 600 m2 10 m2 99 m2 92 m2 9 m2 41 m2 15 m2 5 m2

The main conclusion is that the DMC process is not profitable under the current conditions. The main values of the operating costs are based on the factors given by Seider, Seader, and Lewin (Seider et al., 1999). For the raw materials, CO2 has zero costs hence not taken into account, while the methanol price was estimated at 445 $/tonne (icis.com, 2014) and the DMC sale price at 1000 $/tonne (alibaba.com, 2014b). Note that for typical



Table 16.6 Overview of the Investment Costs. Equipment Type

Membrane reactor Flash vessels Distillation columns Compressors Pumps Heat exchangers Total investment cost

Costs (Peters et al., 2003)

Costs Corrected for 2014

Total Costs including Installation




e 0.2 1.3 1.1 0.1 e 2.7

1.7 0.3 2.0 1.5 0.1 2.0a 7.6

10.1 1.8 11.7 9.2 0.8 11.9 45.5


Prices from Matche.com, (2014).

Table 16.7 Overview of the Operating Costs. Costs (M$/year) Raw materials Utilities Electricity (0.06 $/kWh) Chilled water (4 $/GJ) Cooling water (0.05 $/m3) Hot waterd95 C (0.05 $/m3) Steamd2 bar (5 $/tonne) Steamd5 bar (7 $/tonne) Steamd20 bar (12 $/tonne) Operations (labor related) Maintenance Operation overhead Property taxes Depreciation Depreciation (8%) Membrane life time (3 years (Criscuoli et al., 2001)) Cost of manufacture General expenses Total production cost Sales (DMC) Revenue

Costs (M$/year) 6.63 4.18

0.83 0.77 0.40 0.004 0.60 0.54 1.05 1.32 3.66 0.74 0.91 4.07 3.50 0.56 21.51 2.31 23.82 20.00 3.82



chemical processes, the costs of raw materials account for 50% or more of the total costs. However, in this direct synthesis DMC process, the cost of raw materials is only 28%. This is caused by the high costs for utilities, maintenance, and depreciationdof which the last two are both related to the high investment costs. Maintenance and depreciation are both related to the investment costs and these costs are high due to the low conversion and large excess of methanol present in the process. Lower excess of methanol will result into smaller distillation columns because less methanol needs to be heated or less MIBK is needed to remove DMC from methanol. Nevertheless, at a lower reactant ratio, less DMC is produced resulting in lower conversions and low DMC concentrations at the reactor outlet. The bottom line is that the DMC concentration in the reactor effluent is practically too low to make this process profitable. To investigate the profitability of the process, the effect of the DMC price on the return on investment (ROI) is investigated. For the base case (100%), the process reaches its break-even point at a DMC price of 1215 $/tonne. To reach an ROI >0.2, the DMC price has to go up to 1730 $/tonne, which means that the DMC price has to increase by 73%, before this process becomes economically attractive. In the hypothetical situation that the investment costs could be reduced by 50%, the current price of DMC of 1000 $/tonne is still not economically interesting, being barely sufficient to make the process profitable.

16.2.5 Key Performance Indicators The key performance indicators selected are the energy consumption expressed as kWh/kgDMC and the net CO2 emission expressed as kgCO2/hr, because they are the main points that need attention before this process becomes commercially interesting. Table 16.8 shows the results for these indicators (Kuenen et al., 2016). According to these numbers, the specific energy requirements per kg DMC are high and the process generates more CO2 than it consumes. To have zero CO2 emission, the energy demand must decrease by 61%. Clearly, this process cannot be considered as green, despite the green chemical reaction at its heart. To make this process economically viable, the conversion toward DMC needs to be significantly enhanced. An outgoing reactor effluent with a DMC concentration close to the DMC/methanol azeotropic composition prevents the use of the energy demanding predistillation column. To reach high DMC concentrations (>10 mol%), conversion has to increase by 6e7 times because the current DMC concentration in the reactor


Table 16.8 Energy Requirements and Net CO2 Emission of the DMC Production From the Direct Conversion of CO2a,b.


Net CO2 Emission Based on Fossil Fuels  ðkgCO2 =hr

Net CO2 Emission Based on Renewablesc  ðkgCO2 =hr

16.15 13.61

3505.4 2619.5

2783.9 1898.0

Energy Usage

No heat integration Heat integration

1 MW ¼ 139.84 kgCO2 =hr (Kiss and Ignat, 2012).


Ethermal ¼ 3∙Eelectrical



No CO2 emission is accounted for electrical energy produced by renewables.

effluent is only 1.5 mol%. At the current status, the reactor effluent contains only 0.03 mol% of water. Therefore, further dehydration using a membrane reactor to increase the conversion 6e7 times seems quite unrealistic.

16.3 DMC Synthesis by Propylene Carbonate Transesterification Ethylene and propylene carbonates are interesting intermediates for DMC synthesis. Among the organic carbonates, ethylene and propylene carbonates are exceptional compounds, as they can be quite easily obtained from the reaction of ethylene- or propylene glycol (EG, PG) with CO2 or urea (while the other carbonates require reacting the alcohol with the more reactive phosgene). Then, the cyclic carbonate can be transesterified with methanol, to produce DMC while regenerating the glycol. In this section, we will consider the design of a process for DMC synthesis by transesterification of propylene carbonate (PC). The whole reaction network is shown in Fig. 16.7. Firstly, urea is obtained from CO2 and ammonia. Urea is the carbonation agent in the reaction with PG, leading to PC and ammonia (the later can be recycled to the urea synthesis step). Finally, transesterification of PC with methanol leads to the DMC product and regenerates the PG, which is recycled. This section will consider only the last two processes (highlighted in Fig. 16.7),




Figure 16.7 Dimethyl carbonate synthesis by transesterification of propylene carbonate.

as urea production is a well-established technology. Shi et al. (2017) consider the design of a similar process, but the product is the MeOH/DMC azeotrope.

16.3.1 Thermodynamics The components involved in the process are ammonia, methanol, urea, PG, and PC. These are available in the pure components database of Aspen Plus v10.0, from where their physical properties are taken. Table 16.9 presents the normal boiling points of the pure

Table 16.9 Boiling Points of Pure Components (1 atm). Boiling Point (1 atm) Destination NH3 DMC/MeOH azeotrope (0.14/0.86 M) MeOH DMC Urea PG PG/PC azeotrope (0.98/ 0.02 M) PC

33.34 C 63.78 C

Recycle to urea synthesis process Break or recycle within DMC synthesis process

64.7 C 90.2 C 133 C 187 C 187.7 C

Recycle, within DMC synthesis process Product Recycle within PC synthesis process, or pass through the DMC synthesis process and recycle Recycle from DMC synthesis to urea synthesis Recycle, from DMC synthesis to urea synthesis

240 C

Recycle, within the DMC synthesis process


components (taken from the Aspen Plus database) and their azeotropes (calculated by the NRTL model). Table 16.9 also presents the destination of the chemical species, after the separation steps are performed. Considering the boiling points, it appears that in the PC synthesis process the separation of NH3 is easy. Complete conversion of urea is not required, but high conversion is desirable because this will simplify the separation section of the PC synthesis plant, small amounts of urea in the PC stream being probably tolerated by the DMC plant. Within the DMC synthesis process, separation of PG (to be recycled to the PC step) appears to be easy, although some amounts of PC are expected in this recycle due to formation of PG/PC azeotrope. The DMC/MeOH azeotrope is troublesome. One way to deal with it is by using advanced separation techniques, as explained in the previous section. Another alternative, which will be explored here, is to perform the reaction under such conditions that the reactor mixture allows separation into high-purity DMC (the product) and DMC/MeOH azeotrope (recycled). A good prediction of the vaporeliquid equilibria is required for reliable design of the separation section. The NRTL model can represent the nonideality of the liquid phase, but it requires good binary interaction parameters. For this reason, experimental data were collected from the NIST database, and the NRTL binary interaction parameters were regressed using Aspen Plus facilities. For the MeOH-DMC mixture, the experimental data of Qun-Sheng et al. (2014) and Matsuda et al. (2019) was used. For the DMC-PC and DMC-PG mixtures, the data from Luo et al. (2001) were used, while Mathuni et al. (2011) provided the data for regressing the PCePG binary interaction parameters. Fig. 16.8 compares the experimental data (markers) with the predicted values of the boiling and dew and points (Txy diagram). The agreement is excellent.

16.3.2 Kinetics PC can be obtained from the reaction between PG and urea (urea alcoholysis). The reaction takes place in the presence of MgO as heterogeneous catalyst. For a catalyst concentration of 2%wt., Shi et al. (2017) provide the following reaction rate expression:    3  562:6 r1 ¼ k$cPG $curea ; with kðTÞ kmol m s ¼ 0:02646$exp  T (16.11)




Figure 16.8 Dimethyl carbonateeMeOH vaporeliquid equilibrium. Markers: experimental data; linesdcalculated values, NRTL model with regressed parameters.

The reaction between PC and methanol is catalyzed by strongly basic quaternary ammonium ion exchange resins with hydroxide counter ions. Pyrlik et al. (2012) performed batch and continuous experiments and provided experimental data for several catalysts, among which the Amberlyst 26OH appears the most active. Here, we assume a power-law reversible kinetics:  1 xDMC xPG 2 r ¼ kxPC xMeOH 1 (16.12) 2 Kx xPC xMeOH By regressing the experimental data provided by Pyrlik et al. (2012), the following kinetic parameters were determined. Fig. 16.9 compares the experimental and calculated PC conversion. kð313KÞ ¼ 0:4166  103 kmol=kg=s Kx ð313KÞ ¼ 0:2: Note that the transesterification reaction can also be performed using other basic catalysts, such as sodium methoxide (Holtbruegge et al., 2013), Ce-based catalysts (Kumar et al., 2016), synthetic hydrotalcites (Murugan and Bajaj, 2010), or Verkade super bases (Williams et al., 2009).

16.3.3 PC Synthesis Process The PC synthesis process is presented in Fig. 16.10, where key sizing and mass and heat balance results are also included. Fresh urea and PG (recycled from the DMC synthesis process) are fed, in


Figure 16.9 Kinetics of PCeMeOH transesterification at 40 C. Comparison between the experimental (markers) and calculated (line) PC conversion.

Figure 16.10 Process for propylene carbonate synthesis by urea alcoholysis.




about the stoichiometric ratio, to a stirred tank reactor, operated at 180 C and 10 bar. The reactor has a volume of 20 m3 and contains the MgO catalyst (2%wt, about 400 kg). The PC conversion exceeds 95%. The reactor effluent is cooled to 50 C, and the pressure is reduced to 0.5 bar, when a large amount of ammonia byproduct is vaporized. The liquid is sent to the second reactor operated under similar conditions, where the PG conversion increases to 99.3%. The reactor effluent is sent to the distillation column COL-1, where NH3 is separated as vapor distillate. The column has 10 theoretical stages, 1.4 m diameter and is operated at 2 bar (such that the heat removed in the condenser could eventually be used for generating low-pressure steam) and a small reflux ratio. The bottom product, free of ammonia but containing small amounts of PG and urea, is sent to the DMC synthesis plant.

16.3.4 DMC Synthesis Process In the DMC synthesis process, the transesterification reaction between PC and methanol (MeOH) leads to DMC (the final product) and PG (recycled to the DMC synthesis process). The reaction is reversible, but the equilibrium does not impose severe limitations (Kx ¼ 0.2 at 40 C). The main hurdle to be overcome is the MeOH/DMC azeotrope. This can be separated into high-purity components by various techniques, for example by ED. However, it should be noticed that only DMC must be obtained at high purity. Methanol can be recycled to the reactor together with some amounts of DMC (for example, close to the azeotropic composition). The cost of the recycle will increase (higher recycle flow rate obtained as distillate), but this will be compensated by avoiding the separation of the DMC-MeOH azeotrope. Fig. 16.11 shows the PC equilibrium conversion versus the MeOH/PC feed ratio. The conventional wisdom is to perform the (trans)esterification reactions at a large excess of the alcohol (which is usually the cheaper reactant). Indeed, the conversion of PC is low when PC is the limiting reactant (the MeOH/PC ratio below 2, the stoichiometric value), but increases to about 0.65 for MeOH/PC ratio above 10. In Fig. 16.11, several points are shown on the equilibrium line. For each point, the DMC mole fraction in the DMCeMeOH mixture (that would result after removing the heavy PC and PG species) is presented. Remarkably, for MeOH/PC feed ratios lower than about 3.5, the mole fraction of DMC in the mixture exceeds the azeotropic composition (14% molar DMC), while the equilibrium conversion is still acceptable. This would allow separation of the reactor effluent into MeOH/DMC


Figure 16.11 Dependence of the PC equilibrium conversion (40 C) versus the MeOH/PC feed ratio.

azeotrope (to be recycled within the process), high-purity DMC (the product), and heavy components (PC and PG). The later mixture can be split into components which are recycled within DMC process (PC) and to the PC process (PG). Fig. 16.12 shows the process for synthesis of DMC by the transesterification of PC with methanol. PC (from the PC process and recycled) is mixed with methanol (fresh and recycled). At reactor inlet, the MeOH/PC ratio is about 1. This value is lower compared with the stoichiometric value 2 and to the critical value 3.5 which allows separation of high-purity DMC (see Fig. 16.11). The reaction proceeds in a tubular reactor, operated at 40 C and employing 1270 kg of catalyst. The PC conversion is about 15% (reactant in excess), whereas the methanol conversion is about 30% (limiting reactant). The reactor effluent is sent to COL-2, where the easy (DMC, methanol)/(PC, PG) split is performed. The column has 26 theoretical stages, 1.8 m diameter and is operated at rather low reflux ratio. The distillate, containing DMC and methanol is sent to COL-3. Here, the DMC product is obtained at high purity (99.5%wt.) as bottoms, while the distillate is the DMC/MeOH azeotrope which is recycled to the reactor. The column has 20 theoretical stages, 1.8 diameter and is operated at 0.85 reflux ratio. The bottom stream of COL-2 is sent to COL-4, where the PG/PC split is performed. The distillate contains PG (98% molar) and is sent to the PC synthesis plant. The bottom stream, high-purity PC, is recycled to the transesterification reactor. The column has 40 stages, 1.4 m diameter, and is operated at a reflux ratio of 2.47.




Figure 16.12 Process for dimethyl carbonate synthesis by propylene carbonateemethanol transesterification.

16.3.5 Key Performance Indicators The economic performance of the process was assessed using Aspen Plus Economic Analyzer. The total equipment installed cost is estimated at 4.95  106 USD. The main items are the urea alcoholysis reactors (each of about 0.45  106 USD), the distillation columns of the DMC synthesis plant (COL-2, COL-3, and COL-4, each at about 1  106 USD). The total capital cost is estimated at 12.32  106 USD. The total utilities cost is estimated at 4.9  106 USD/year, with the main contribution coming from high- and low-pressure steam (63% and 27%, respectively). The specific energy requirements are 3.7 kWh/kg DMC, which is 22% of the direct DMC synthesis. Note that Aspen Plus Energy Analyzer identifies targets for energy savingsdabout 20% for heating utilities and 23% for cooling utilities. Developing a heatintegrated flowsheet is left as an exercise for the reader. To calculate the profitability of the process, we assume the following prices: uread247 USD/tonne (www.indexmundi.com); methanold445 USD/tonne (icis.com, 2014); and DMCd1000 $/tonne (alibaba.com, 2014b). The revenues that could be


obtained by selling the NH3 obtained from the process are neglected. Under these assumptions, the cost with raw material is 480 $/tonne DMC, while the utility costs are 138 $/tonne DMC. The cost with capital depreciation is 1.232  106 USD/ year (assuming a payback period of 10 years), which amounts to about 40 $/tonne DMC. Thus, the total specific costs are 658 USD/tonne DMC, which leads to an estimated profit of 342 USD/tonne DMC (about 11  106 USD/year).

16.4 Conclusions This chapter presents two alternatives for DMC synthesis. Firstly, a novel process design for the direct conversion of CO2 to DMC is developed. A membrane reactor is used to remove selectively the water by-product, and thus to overcome the equilibrium limitations. The results described here allow a valuable insight in the limitations of this newly proposed process and provide some potential solutions to overcome the envisaged problems. Thus, only low CO2 conversion is achieved even when using an excess of methanol, resulting in a membrane reactor effluent with a low DMC concentration. Purifying this diluted reactor outlet stream to the desired product concentrations demands large separation units and a significant amount of energy (specific energy use of 13.61 kWh/kg DMC). This leads to high investment and utility costs, which make this process design not profitable. To make this process economically attractive, the focus for new membrane reactors could be on the removal of DMC (instead of water) from the reaction area. Ending up with a concentrated DMC product stream could reduce significantly the costs of the separation process. The second process proceeds in two steps: (a) ureaePG alcoholysis to PC, followed by (b) transesterification of PC with methanol. The main issue is separating DMC from the DMC/methanol mixture, which is difficult due to azeotrope formation. The smart solution suggested here consists in performing the transesterification reaction at MeOH/PC ratio well below the stoichiometric value, such that the reaction proceeds at an excess of PG, while MeOH is the limiting reactant. Thus, the composition of the reactor effluent falls in the region where the MeOH/DMC ratio allows separation into high-purity DMC product and MeOH/DMC azeotrope to be recycled. The specific energy use is only 3.7 kWh/kg DMC. The process becomes profitable, with estimated revenues of 342 USD/tonne DMC. The inefficiency of the direct DMC synthesis is due to the low chemical reactivity of the CO2 molecule. To increase the CO2




conversion in the reaction with methanol, more severe reaction conditions are needed, for example, pressures much higher than 50 bar. On the other hand, the first step of the indirect routedurea synthesisdis a well-established technology in which rather drastic conditions (140e175 bar, 190 C) are employed to transform CO2 into a more reactive species, at a competitive cost. Thus, the next steps (urea alcoholysis and PC transesterification) can be performed under milder conditions, with the result of lower investment and operating costs.

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