Membrane Reactor for one-step DME synthesis process: Industrial plant simulation and optimization

Journal of CO₂ Utilization 22 (2017) 33–43

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Journal of CO2 Utilization journal homepage: www.elsevier.com/locate/jcou

Membrane Reactor for one-step DME synthesis process: Industrial plant simulation and optimization

T

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Marcello De Falco , Mauro Capocelli, Alberto Giannattasio Unit of Process Engineering, Department of Engineering, Università Campus Bio-Medico di Roma, via Álvaro del Portillo 21, 00128 Rome, Italy

A R T I C L E I N F O

A B S T R A C T

Keywords: Membrane reactor Selective membrane One-step DME synthesis Industrial plant optimization CO2sequestration

DME production represents a possible route of CO2 valorisation in the decarbonization pathway. This paper provides an insight into the basic design and process analysis of a DME production plant by implementing Membrane Reactors to convert CO2 at acceptable values. The paper provides the process analysis of the overall plant architecture related to an innovative DME production process configuration, called Double Recycling Loop DME (DRL-DME), enables the utilization of a pure CO2 stream as sweeping gas in the permeation zone (PZ) and a double recycling loop to reintroduces the un-reacted syngas and the outlet stream from the PZ. The mathematical model of the entire plant is presented and implemented to study the effect of the process parameters on the performances. This paper brings a step forward by determining the possible configuration and the performance of a whole industrial plant, able to realize the DME synthesis from CO2-rich streams in membrane reactor with recirculation. The whole process includes the utilization of a pure CO2 external feed, the steam production and all the necessary separation units. By fixing the set point of the process control variables, the recirculation streams are calculated though a mathematical algorithm founded on previous literature results. Through the mathematical model it is possible to estimate the CO2 conversion and DME yield as well as the main features of the plant, by varying some important process parameters (e.g. feedstock composition, reactor pressure, H2/COx).

1. Introduction The well-known issue of Greenhouse Gases (GHGs) emissions is leading to an unacceptable increase of the earth's average temperature with potentially catastrophic consequences on the ecosystems [1]. This emerging and shared concern has been promoting the development of international agreements to impose stringent medium-term targets on the equivalent CO2 emissions [2,3]. The national regulation bodies are defining strategies to discourage the CO2 production and to increase the percentage of energy generation by renewable sources, alternative to the fossil fuels, in order to mitigate the global warming [3]. On the other hand, the decarbonization of the present production processes based on fossil fuels goes through the CO2 capture and sequestration (CCS) adopted with retrofit solutions [4–6]. The most consolidated solutions implement the amine scrubbing units at stationary power plants [7,8] that produce a low-cost and relatively pure CO2, a valuable feedstock of nearly zero cost for conversion to fuels and chemicals [9]. Olah et al., in fact have described the possibility of migrating towards a Methanol Economy based on the carbon dioxide recycle to methanol, dimethyl ether and subsequently to synthetic hydrocarbons,

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as the future paradigm for renewable fuels in a low-carbon economy [10]. According to this view, the future productive capacity should be accompanied by the use of CO2 as a building block in organic syntheses to obtain valuable chemicals and materials (CCU − Carbon Capture and Utilization), an overall market of CO2 valorization that could overcome the present industrial use of CO2 of one order of magnitude [9–11]. In the context of CCU, the carbon dioxide is seen as a reactant of a chemical process to be converted into a desired product [11]. Therefore the CO2 can be used for syngas production through the dry reforming process [12,13], for methanol synthesis [10], for microalgae cultivation [14], for methane production [15], as well as C1-building for polymer synthesis [16]. In this paper, the “CO2 valorization” technology analyzed is the Dimetyl Ether (DME) production through a CO2-rich feedstock. DME is an organic compound mainly used as a reagent for the synthesis of widely applied products as the dimethyl sulfate (a methylating agent), methyl acetate and light olefins [17,18]. The growing interest on DME production is due to the characteristics similar to those of liquefied petroleum gas (LPG) and to the excellent combustion property, that make the DME a potential candidate as a fuel in the compression-ignition engines [19]. It has been experimentally demonstrated that

Corresponding author. E-mail address: [email protected] (M. De Falco).

http://dx.doi.org/10.1016/j.jcou.2017.09.008 Received 27 July 2017; Received in revised form 6 September 2017; Accepted 11 September 2017 Available online 19 September 2017 2212-9820/ © 2017 Elsevier Ltd. All rights reserved.

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Nomenclature cp D d F Fi Fsweep G GHSV Hmix Kj Ki Keq L P P Po Pa Pj pj ∼ PH2

∼ PH2 O ri

R Re Ri Re Sw T Tv yi YDME U

Specific heat of the gaseous mixture at constant pressure, J mol 1 K−1 Reactor inside diameter, m Particle diameter, m Total molar flow rate, kmol/s molar flow rate of the component i, kmol/s sweeping gas flow rate in the permeation zone, kmol/s binary interaction parameters Gas hourly space velocity, h−1 Enthalpy of the gas mixture, kJ Rate constant of the reaction j Adsorption equilibrium constant, Pa−1 Equilibrium constant Reactor length, m Reactor pressure in fixed bed side, Pa Permeation zone pressure, Pa Pressure in cooling fluid side, bar Atmospheric pressure, bar Partial pressure of speciesj in fixed bed side, Pa Partial pressure of speciesj in permeation side, Pa

XCO2 z

Greek letters ε η ρ ρc

R P in Out

CO + H2O ↔ CO2 + H2 2CH3OH ↔ CH3OCH3 + H2O

ΔHo = −41.0 kJ/mol ΔHo = −23.0 kJ/mol

Relative to the reactor Relative to the global process Inlet Outlet

[25]. The integration of selective membrane in the catalytic reactors represents one of the most promising and interesting solution to this aim [29–34]. Such reactor configuration, called Membrane Reactor (MR), allows the selective removal of water steam produced by reactions (1) and (3), thus supporting continuously the methanol and DME production. Hydrophilic selective membranes are classified in: i) microporous zeolite membranes; ii) amorphous microporous membranes; iii) polymeric membranes. According to Rodhe et al. [26,31], who reviewed the water selective membranes behaviour for the Fischer-Tropsch synthesis (at operating conditions similar to those of the DME production), only microporous zeolite membranes can be applied at temperature > 200 °C with satisfactory permeation performance (H2O permeance within the range 10−7 − 10−6 mol/s m−2 Pa−1 and a H2O/H2 selectivity > 10). Many works have clearly demonstrated the benefits of the selective membrane integration due to improvement of CO2 conversion, DME yield and selectivity [32–34]. The Group of Diban et al. [33,34] also agrees that zeolite membranes are the only to withstand the operating conditions required for DME synthesis. On the other hand, their work underlines that lower DME yields are obtained for CO2-rich feedstock when zeolite membranes, allowing the permeation of reactants such as methanol, are implemented [31]. Significant differences in H2O-selectivity over methanol are reported in the literature in relation to the MR configuration. Attending to the selectivity of the zeolite membranes evaluated by Diban et al. [33,34], the MeOH selectivity was of the same order of magnitude to that of H2O. According to them, reproducing a sieving mechanism of mass transport through zeolite membranes that promotes the permeation of the smallest molecules such as H2 and H2O, is the way to realize ideal H2O selectivity MR and to enable the CO2 valorisation by DME synthesis in MR [34]. Unfortunately at the moment, this type of zeolite membrane are not ready for the implementation in MR because of the low hydrothermal stability shown in previous tests [31,35]. To overcome the issue related to the MeOH permeation, recent innovations include also the development of a catalytic membrane reactor

adding DME to the diesel fuel allows an enhancement of the combustion properties, with a reduction of smoke, NOx and CO2 emissions [20]. DME is also able to stores energy more conveniently and safely compared to the hydrogen in fuel cell applications [21,22] in the transportation sector and can also be converted into ethylene and/or propylene, among the main building blocks of many synthetic hydrocarbons. Perez-Uriarte et al. recently investigated the effect of the catalyst properties, particularly matrix acidity and its properties, in the DME conversion to light olefins [23,24]. Traditionally, DME has been produced in a two-step process from purified methanol, in turn obtained from syngas [25]. More recently, the combination of methanol synthesis and dehydration has been developed to synthesize DME directly from syngas (CO, CO2, H2) in a single catalytic tubular reactor. The direct process allows for a higher COx conversion and for a simpler reactor design, therefore resulting in lower DME production costs [25]. The simplified reactions scheme considered in this paper and in previous literature works [26–30], also, includes the CO2 hydrogenation, the water-gas shift and the methanol dehydration, as described in the following: ΔHo = −49.4 kJ/mol

Packed bed void fraction Permeation side to reaction side pressures ratio Density of gas phase, kg/m3 Catalyst bed density, kg/m3

Subscripts and superscript

Hydrogen permeance through the selective membrane, mol/s m−2 Pa−1 Steam water permeance through the selective membrane, mol/s m−2 Pa−1 Rate of the reaction i, kmol/(kg ∙ s)

CO2 + 3H2 ↔ CH3OH + H2O

Ideal gas constant External tube radius, m Internal tube radius, m Reynolds number Sweeping gas to inlet total flow rate ratio Temperature, K Steam temperature in cooling fluid side, K Mole fraction of the component i DME yield Overall heat transfer coefficient between pipe and shell, W/(m2 K) Conversion of CO2 Axial coordinate, m

(1) (2) (3)

The process is exothermic, is promoted at low pressure and can be supported by a bi-functional catalyst as Cu-ZnO-Al2O3/HZSM-5 for the simultaneous promotion of methanol synthesis, methanol dehydration and consequently of CO and CO2 conversion even at high temperatures [27–29]. The increase of the CO2 content in the inlet feedstock has a detrimental effect to the DME synthesis process performance resulting into lower syngas conversion and into lower DME yield [29,30]. Therefore, if the DME production process is fed with a CO2-rich feedstock, innovative configurations are required to achieve satisfying performance and recycling/valorising the by-product carbon dioxide 34

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streams to control the composition, the process simulation depends on the development of a mathematical algorithm, including the cited models, that represents the core of the present work. The capability of the plant to convert a CO2-rich syngas in pure DME is evaluated by means of the commercial softwares Matlab and Aspen Plus. The effect of the process parameters on the performances are evaluated and the optimized plant configuration is discussed. The simulation allowed to understand the effect of the feedstock composition at the reactor inlet, which influences the flow-rate of the recirculation and the performance of the separation units in the DRL Plant. The theoretical performance described in the literature [30–34,39], focused on the theoretical DME synthesis from CO2-rich streams in membrane reactor with recirculation, are partially reproduced and in some cases improved by the performances of the more DRL-DME complex process that, in conclusion, could represent the prodrome of the basic design of an industrial plant.

with solid-acid catalyst, namely, F-4SF resin that demonstrated a methanol conversion of 36% and 100% selectivity toward DME [29,36]. A dual-layer zeolite membrane (mild acidity layer for the methanol dehydration and hydrophilic Na-LTA layer to remove the water) has showed higher methanol conversion (91%) and 100% DME selectivity [37]. Also Fedosov et al. obtained similar results by testing the catalytic dehydration of methanol with a selective and thermally resistant membrane made of a zeolite layer synthesized on a metal–ceramic support [38]. Our research group has been focusing on a complete study of DME production process in the framework of the CO2 valorisation, by starting from the thermodynamic analysis [29] to the simulation of the reaction performance with the integration of microporous zeolite membranes [39]. Recently, Diban et al. investigated the influence of the GHSV on the process addressing the mass transport rate of all the components present in the reactive system through zeolite membranes accounting for the sweep gas recirculation [33,34]. These results show that the application of zeolite membrane promoted the transformation of CO2 into DME and promise interesting prospects for the application of the MR including the gas recirculation to recycle the CO2 as DME on a large scale [34]. On the other hand, no work has been published on the analysis of the complete process of DME production from CO2-rich feedstock, also considering the CO2 closed loop as well as the final product purification, the separation of the condensable, the distillation of the different streams, the reactants recycling and the possibility of heat recovery. On these basis, the present paper provides the proof of concept of a novel DME production process and the first mathematical model for its description, for the investigation of its technical feasibility and for the sensitivity analysis of the overall plant architecture. Fig. 1 depicts the overall schematic of the innovative plant configuration, implementing the Double Recycling Loop DME (DRL-DME) production process. This is based on:

2. Process description Fig. 2 shows the Process Flow Diagram of the configuration studied. It is composed by the Membrane Rector (MR) unit, the Steam Generation Unit, the Separation & Recirculation Unit and the Distillation unit. The recirculation units realize of the separation, mixing and recirculation (in-series operation units) of different streams to prepare the feedstock and the sweeping gas sent to the MR in the reaction zone (RZ) and the permeation zone (PZ), respectively. The plant inlet streams are the reactant syngas (Stream 1), produced by a fossil feedstock or derived from the biomass gasification, and a pure CO2 stream (Stream 21) integrating the sweeping gas (stream 21) in the PZ with a portion of the PSA off gas (Stream 20). The two recycles characterizing the Double Recycling Loop are the streams 5 and 17, fed to the reactor from the separator S-1 and from the PSA unit, respectively. Therefore, the inlet composition of the MR is adjusted by recirculating the streams from the first separation stage (Stream 5) and the H2-rich stream, coming from the PSA. The Stream 5 is the recirculated vapour obtained after the partial condensation stage occurring in the HE-1. It is considered in equilibrium with the Stream 6 at the set-point temperature of the condenser (5–7 °C). This temperature is adjusted in order to obtain a negligible composition of DME (the most volatile compounds among the condensable) in the vapour fraction (Stream 5). In this configuration, the CO2-rich stream (Stream 21), mixed with the recirculated Stream 20, sweeps the permeated H2O/H2 flux out. The water in Stream 13 is separated by condensation, while the CO2 and H2 are fed to the PSA stage to separate the two components. The CO2 stream (Stream 18) is then recycled to the PZ of the reactor. A portion of the CO2 stream can be recirculated as reactant into the RZ of the membrane reactor inlet by the Stream 19. By the DRL-DME plant

• a membrane reactor (MR), representing the technological core of the • •

overall process, to obtain the desired DME yield and selectivity, also at high CO2 content in the inlet feedstock as previously demonstrated by our research group [39]; the use of a pure CO2 stream as sweeping gas in the membrane reactor permeation zone (PZ); a double recycling loopthat reintroduces both the un-reacted syngas and the outlet stream from the PZ in the process.

In the following, the DRL-DME plant configuration is described and the related mathematical model is proposed. The model description is characterized for each of the section of the DRL-DME process. Since the process includes a series of separation units and recirculation of gas

Fig. 1. Block scheme of the Double Recycling Loop (DRL) one-step DME production process with a membrane reactor.

35

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Fig. 2. Process Flow Diagram of the DRL-DME production plant (solid line = reactants-products, dashed line = membrane reactor sweeping gas line; dotted line = reactor cooling line).

coherence between the heat available and the boilers’ heat duty is verified.

architecture, the CO2 composition in the reactor feedstock is much higher than the syngas, since part of the sweeping gas is recycled to the reaction zone. The H2 (Stream 15) is recycled to the RZ of the membrane reactor inlet after being compressed by the CC-1 to reach the RZ pressure. H2/COx in the MR is controlled by adjusting the external CO2 (Stream 21) in order to achieve a convenient CO2 conversion and DME yield. Despite the ability of the selective membrane to ensure high conversions even in the presence of large concentrations of CO2, such a concentration has to be < 40% since, above this value, the reactor performance is not satisfactory in terms of conversion and DME yield [32,39,41]. The outlet from the MR reaction zone (Stream 3) is composed by the un-reacted syngas (CO, H2 and CO2) and by the condensable products (DME, methanol and un-permeated water steam) and is fed to a first separation unit S-1. In this unit, the mixture is cooled down to produce a liquid fraction (Stream 6: DME, methanol, water and dissolved carbon dioxide) separated from the non-condensable components (Stream 5) recirculated to the membrane reactor inlet (first recycling loop). The liquid mixture is then fed into the Distillation Unit consisting of 3 distillation columns (C-1, C-2, C3). The column C-1,operating at the same pressure of the reactor (minus the pressure drops) in order to enhance the separation, is aimed to separate DME and CO2 (with traces of CO and H2) from CH3OH and H2O. The second column C-2 is fed by the bottom residues (Stream 8) of column C-1 and separates CH3OH from H2O. The third column C-3 is fed with the distillate of the column C-1 and separates an almost pure CO2 (available at high pressure and a storage ambient temperature) from the bottom product (DME). Hence, the plant output streams are: pure DME (Stream 10) obtained as residue from the distillation column C-3, liquefied CO2 (Stream 9) and pure methanol (stream 11) separated from water (stream 12) in the distillation column C-2. In the Steam Generation Unit, the membrane reactor is cooled down by the evaporation of the saturated water to produce the steam (@ Tv = 470–500 K) that can be supplied to the reboilers of the distillation columns (HE-5 and HE-7). During the plant simulation, the

3. Mathematical modelling The reactor simulation implements the steady-state mass, energy and momentum balance equations for the membrane fixed-bed reactor assuming a 1-D plug-flow [32,39]. The reactor is packed with bi-functional catalytic particles of Cu-ZnO-Al2O3/HZSM-5·The mass balances of the six components for the fixed-bed side are expressed by Eqs. (4)–(9), where the trans-membrane molar fluxes for hydrogen and water are indicated as JH2 and JH2O. The sweep-gas is flowing through the PZ in counter-current to the reactants in the RZ. For sake of simplicity, an infinite H2O/CH3OH permeation selectivity is assumed to allow the convergence of plant modelling.

dF˙H2 O = ρc (1 − ε )·π (Re2 − Ri2)·(r1 − r2 + r3) − JH2 O ·2πRi dz

(4)

dF˙H2 = −ρc (1 − ε )·π (Re2 − Ri2)·(3r1 − r2) − JH2 ·2πRi dz

(5)

dF˙CO = −ρc (1 − ε )·π (Re2 − Ri2)·(r2) dz

(6)

dF˙CO2 = −ρc (1 − ε )·π (Re2 − Ri2)·(r1 − r2) dz

(7)

dF˙CH3 OH = ρc (1 − ε )·π (Re2 − Ri2)·(r1 − 2r3) dz

(8)

dF˙C2 H6 O = ρc (1 − ε )·π (Re2 − Ri2)·(r3) dz

(9)

where F˙i is the mass flow rate of the i-th component, z is the reactor axial coordinate, ρc is the ratio of mass catalyst on the reaction volume, r1, r2, r3 are the rates of the reactions (1), (2) and (3), respectively, Re and Ri are the external and internal tubes radii and ε is the void fraction in the catalyst bed. The heat balance inside the reactor, neglecting the 36

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heat flux from the reaction zone to the permeation zone, is described by Eq. (10) where ΔHir is the heat of reaction and the enthalpy of the gas mixture that has been written in relation to the temperature inside the reactor T by calculating an average specific heat:

component, T is the mixture temperature and the binary interaction parameters are calculated as follows and reported in Table 1.

Gij = exp(−αij τij )

nc

τij = cij + dij (T − T0) +

d ⎡∑ F˙j c pj·(T − Trif ) ⎤ nr ⎞ ⎦ =⎛ ⎣ j=1 r 2 2 ⎜∑ (−ΔHi (T )) ri⎟·ρc ·(1 − ε )·π·(Re − Ri ) dz i = 1 ⎠ ⎝ − U (T − Tv )·2πR e (10)

r2 = K2

3

(11)

PW − PCO2 PH2/ K eq2 PCO ) (1 + K CO2 PCO2 + K CO PCO + √ (K H2 PH2) )

(12)

P2 P r3 = K3 ⎛⎜ CH 3OH − CH 3OCH 3 ⎞⎟ P K eq3 ⎠ ⎝ W

(13) th

where Pj represents the partial pressures of the j component (j = H2O, H2, CO, CO2, CH3OH, DME) in the gaseous mixture, K1, K2 and K3 are the kinetic constants and Kj the Langmuir adsorption kinetic constants of the respective compounds on the catalyst [27,40,42], calculated in function of the operating temperature according to the Arrhenius’ law [27,39]. The trans-membrane molar flux for hydrogen and water is proportional to the partial pressure difference across membrane and to the permeability according to: ∼ JH2 = PH2 (p H2 − p H2 ) (14)

∼ JH2 O = PH2 O (p H2 O − p H2 O )

∑j xj τji Gji ∑k xk Gji

+

∑

(15)

j

x j Gij

∑k

xk Gkj

*(τij −

∑m xm τmj Gmj ) ∑k xk Gkj

+ eij ln T + fij T

(18)

F=D+R

(19)

FxF = DxD + RxR

(20)

FhF + QR = DhD + RhR + QD

(21)

where F is the feed molar flow, xF is the molar fraction of DME in F, D is the distillate mole flow, xD is the molar fraction of DME in D,R is the residue molar flow,xR is the molar fraction of DME in R, hF is the feed enthalpy, hD is the distillate enthalpy, hR is the residue enthalpy, QR is the reboiler heat duty and QD is the condenser heat duty. Performance parameters of the distillation column considered in the sensitivity analysis are the composition of the distillate and residuexD, xR as well as the DME recovery defined as αDME = RxR/FxF. Since the process requires the product recirculation to adjust the feedstock composition as well as the sweeping flow rate, the overall process simulation can be obtained through the implementation of the algorithm reported in Fig. 3. The recycled streams into the RZ are 5 and 17 whereas the input stream, controlling the sweep gas flowrate, is the external CO2 (stream 21). The algorithm starts by assuming the syngas composition (Stream 1), the desired H2/COx ratio at the RZ inlet (Stream 2) and other fixed parameters discussed in the following (e.g. η, SW, etc.). The solution process continues by calculating the main parameters regarding these recycles: the composition of the inlet streams in terms of CO2/COx and the flowrate of the external CO2stream. These two parameters are the chosen ones for the convergence and the algorithm termination. By solving the models of the MR unit and the following separation processes (Recirculation Unit), the gas composition of the MR inlet (Stream 2) as well as the flowrate of the Stream 21 can be calculated. These values are compared to the ones adopted in the previous calculation stages. The iterations stop when the difference between the values simulated at the iteration stage i+1 and the ones obtained at the stage i is confined below a fixed limit as depicted in Fig. 3(ε≤1%).Once the correct flowrate and composition of the recirculating streams are

where p and p are the partial pressure in the retentate and permeate ∼ ∼ sides, respectively, and P H2 and P H2O are the permeability across the membrane of hydrogen and water steam [43]. The cooling of the catalytic tubes takes place by the means of natural circulation of boiling water between the steam drum and the reactor. The vapor produced by the MR cooling can be used in the reboilers of column C-1 and C-3. For a deeper description of MR mathematical model including the balance of heat and momentum, one might refer to [39]. In both condensation and distillation units the physical equilibrium model to obtain reliable multi-component vapor-liquid equilibrium is the NRTL, as recommended by Song [44]:

ln γi =

T

The condensation curves (of the heat exchangers, HE in Fig. 2) were obtained though the heat and material balances coupled with the cited equilibrium conditions and by fixing two degrees of freedom (temperature and pressure). The process analysis of the distillation units was limited to the individuation of the reflux ratio and the reboiler/condenser duty by fixing the number of stages and by knowing the feed temperature and composition. In order to obtain reliable results through the Aspen Plus simulations, the number of theoretical stages was pre-determined by implementing a short-cut method (i.e. theMcCabe & Thiele method) considering a system of two components, the heaviest of the lighter components and the lightest of the heavier components of the feed stream.The overall heat and material balance for the pseudo-two-component system can be written as:

PCO2 PH2 (1 − PCH 3OH PW / K eq1 PCO2 PH3 2 ) (1 + K CO2 PCO2 + K CO PCO + √ (K H2 PH2) )

bij

With Tii = 0 and Gii = 1

The heat flux toward the shell is evaluated considering an overall heat transfer coefficient U and the temperature difference between the interior of the reactor at T and the external steam at Tv according to the relation reported in a recent study [39]. The reaction rates for methanol synthesis, water gas-shift and methanol dehydration on the bi-functional catalyst are described by Eqs. (11)–(13) [27]:

r1 = K1

(17)

(16)

where γ is the activity coefficient, x is the mole fraction of the Table 1 Binary Interaction Parameters for the NRTL model from the Aspen Physical Property System. Component i Component j

C2H6O-1 H2O

C2H6O-1 CH3OH

CH3OH H2O

CH3OH CO2

H2O CO2

C2H6O-1 CO2

AIJ AJI BIJ BJI CIJ

−0.27592 3.61795 622.773 −562.157 0.359115

2.4526 −1.2715 −541.881 480.918 0.3

−2.626 4.8241 828.387 −1329.54 0.3

0 0 −48.6751 126.042 0.3

0 0 291.41 177.401 0.3

0 0 586.89 −423.458 0.3

37

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Fig. 3. Flowchart of the algorithm for the simulation of the overall DRL-DME Process Flow.

papers in the literature [45].

determined, the simulation continues by calculating the fed to the Distillation Unit and therefore by individuating the flow rate and composition of the final products. To have a precise description of the plant, the MR model has been developed in Matlab™ environment according to the procedure described in [39], whereas all the separation stages has been simulated in Aspen Plus ™ environment. The willing of a rigorous description of the plant led to the cooperation of the two software in the Microsoft Excel ™ environment as reported in very few

4. Results and discussions The reactor behaviour was evaluated varying the process parameters (as the reactor operating pressure and the H2/COx ratio of the Stream 2) by quantifying the following values referred both to the reactor and to the overall plant performance: 38

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R XCO = 2

P = XCO 2

R YDME =

P YDME =

content in the feedstock, at fixed H2/COx and recirculation through the parameter SW, may lead to lower yields due to lack of carbon monoxide as reagent. In fact, this particular behaviouris connected to the fact that, for sake of simplicity and clarity the parameter SW has been kept constant over the present simulations.

(2) (3) FCO − FCO 2 2 (2) FCO 2

(22)

(1) (22) (9) + FCO − FCO FCO 2 2 2 (1) (22) + FCO FCO 2 2

(23) 4.3. Separation stages analysis & overall performances

(3) 2FDME (2) (2) FCO + FCO 2

(24)

The overall process simulation allows to justify and to frame in a real plant context, the promising results of De Falco et al. related to the membrane reactor [39] and Diban et al. including the recirculation of the unreacted gaseous mixture [34]. The plant simulation has been performed, following the algorithm of Fig. 3,by fixing the design specification of the condensation stages as well as the column C1 and C2. The distillation column C3 is of minor interest since CH3OH and H2O only represent 2–3% of the stream 7 to the condensation stages and its feasibility is linked to a techno economical analysis since the overall product streams and the economic value of the products could be not comparable to the actual costs of investment. The remaining process variables are reported in the captions. The variation of the feed temperature has negligible effect on xD and xR, therefore the feed (Stream 6) has not been preheated (by implementing an additional internal heat recovery) and is fedwith few degrees of subcooling. Table 3 summarizes the design specification finally adopted for the distillation units including the calculation of reboiler duty and the number of theoretical stages assumed. By varying the reflux ratio and the recovery ratio, the sensitivity analysis investigates the changes in purity of DME as residue, the presence of CH3OH in the residue and the purity of CO2 as a distillate to determine the optimal operating conditions. The effect on CO2 purity is negligible because of the high relative volatility between the top and bottom components and is not reported. The effect on the recovery ratio and reflux ratio on the DME purity and reboiler duty is visible in Fig. 7 and 8, respectively. The global results including the flow rate and composition of the main streams are reported in Table 4.Worth of mention that the vapour, produced by cooling the MR tubes as described in our previous paper [39], at a saturated temperature Tv = 500K, can be integrated in the process to supply the reboiler duty for both columns C-1 and C-2 since its flowrate is more than 64 t/h for a total produced heat around 28 Mkcal/h, enough to fulfil the heat requirement of the plant (Table 4).

(10) 2FDME (1) FCO 2

(1) (21) + FCO + FCO 2

(25)

where the apex indicates whereas the parameter refers to the reactor (R) or to the overall process (P). The operating conditions, the fixed process parameters as well as the properties of the conventional fixed bed reactor are summarized in Table 2. The effect of temperature on the direct DME synthesis in MR has been previously addressed [39]. Based on the literature results, the temperature in the reactor mainly depends on the cooling media temperature (at fixed geometrical parameters, gas composition and flow-rate) that is fixed at Tv = 500 K. The permeation zone operating conditions are defined by the following parameters: 4.1. Effect of the operating pressure R R and YDME is illustrated in The effect of the RZ pressure on the XCO 2 Fig. 4.Globally, the pressure has a positive effect on the MR performance as stated by the thermodynamics of the process [29,41]. In fact, the CO2 hydrogenation reaction is promoted by the high pressure, thus increasing the methanol content and, consequently, promoting the reaction of methanol dehydration. Higher pressures support the methanol production but does not influence the methanol dehydration and, consequently, the DME synthesis (slightly reducing the DME selectivity) [29]. Membrane reactor performances improvement is also due to the enhancement of the pressure driving force across the membrane (Eqs. (14)–(15)).

4.2. Effect of the feedstock composition The most important parameter, as suggested in the literature, is the H2/COx ratio of the inlet stream to the reactor. It depends on the feedstock, the recycle as well as on the external CO2 (stream 21). In the present innovative process scheme, it is controlled by the flow-rate of Stream 19. Therefore, the plant simulation has been carried out at different feedstock composition by fixing the value of theH2/COx ratio for the Stream 2 (reactor inlet) equal to 3, the desired design specification to reach as reported in Table 3.Figures 5 and 6 report the effect of the feedstock composition as contour plot in the H2eCO2 plane. Firstly, it can be said that, the lower the CO2 content of the feedstock, the higher the flowrate of the external CO2 fed to adjust the sweeping gas flow rate. Moreover, the higher H2 content in the Stream 1 increases the need of the external CO2 stream. This aspect is visible in Fig. 5 (a). Moreover, by increasing both the H2 and CO2 content in the stream 1, the CO2 conversion increases as showed in Fig. 5 (b). The blank portion of the plot corresponds to not possible syngas composition since the sum of the molar fractions should be equal to 1 at tops. On the other hand, the recirculation brings a non-linear effect in the YDME as showed in Fig. 6. In fact, at fixed H2/COx in the reactor, the YDME varies slightly with the H2 and CO2, ranging between 0.5 and 0.7 with a stronger dependence on the H2 initial content. This is due to the recirculation of different streams that allow to reach a H2/COx ratio of 3 at the inlet. This allows to balance the yield independently from the feed composition in the investigate range and operating conditions. On the other hand, the higher CO2 inlet composition at low H2 content could limit the YDME making the DRL process less convenient. Minimum yields are reached at higher CO2 content in Fig. 6. Too high H2 and CO2

4.4. Effect of the H2/COx Given the peculiarity of the process, the performances of the MR Table 2 Operation conditions and properties of conventional DME fixed bed reactor and overall DRL_DME Process for the parametric study.

39

Variable

Value

D L ρc ε dp H2/COx in MR GHSV Tv Pin Tin η Sw Tout (HE-1) TC (C1) TC (C3) Tout (HE-2)

0.038 m 1m 1900 kg/m3 0.33 0.002 m 3 6000 h−1 500 K 60 bar 532 K 0.058 5 270 K 373 K 270 K 270 K

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R Fig. 6. Effect of feedstock composition (Stream 1) on the YDME . P=60 bar, H2/COx=3, Feedstock flow rate (Stream 1) = 1kmol/s, GHSV=6000 h−1, the values of the other process parameters reported in Table 2.

Fig. 4. Effect of pressure in the RZ on the XCO2 and YDME. H2/COx=3, Feedstock flow rate (Stream 1) = 1kmol/s, GHSV=6000 h−1, the values of the other process parameters reported in Table 2.

difference between the overall performances and the one of the MR alone. Figs. 8 and 9 reports the XCO2 and YDME for both the reactor boundary limits and the overall plant as defined by Eqs. 22–25. Stated that the increase of the H2/COx also leads to an increase of the CO2 conversion and DME yield [39]. To perform these simulations, a rich H2 feedstock has been chosen to have a large variability on H2/COx available in the reactor. CO2 conversion and DME yield vary from 17% to 60% approximately for both the reactor and the overall process as H2/COx increases from 1 to 3. The difference between the curves allows to detect the effect of the recirculation and actual efficiencies of the separations for the selected P operating conditions. The CO2 conversion of the plant XCO 2 is lower than R the simple XCO2 because of the process internal CO2 losses in the separation stages, mainly absorbed in the water at higher pressure. These bear an increase in the denominator of equation 23 with respect to that P of the equation 22. On the other hand, the overallYDME is favoured by the internal recycling loop of reaction products and is higher than the R showing the positive effect of the recirculation on the process mere YDME performance. This aspect, though not too visible due to the effective separation efficiency and the feedstock conditions and product recirculation blocked by all other process parameters, confirms the trends reported by Diban et al. [34]. Looking also at the relative small size of the necessary equipment it can be said that the recirculation in the DME

Table 3 Design Specification and Simulation results of the Distillation Columns. COLUMN

C-1

C-3

Light Component (LC) Heavy Component (HC) P [bar] QD [Mkcal/h] TD [K] QR [Mkcal/h] TR [K] Number of Theoretical Stages (# feed stage) Feed Flow [kmol/h] Distillate Flow [kmol/h] (LC Mole Frac) Residue Flow [kmol/h] (HC Mole Frac)

DME CH3OH 60 −1.417 373 9.021 495 7 (2) 1215 1170 (0.474) 45 (0.468)

CO2 DME 60 −3.924 270 2.403 407 10 (7) 1170 522 (0.983) 556 (0.994)

alone (as described in the previous sections and in the literature) cannot be directly extended to the overall plant performances because of the different recirculation streams following several separation stages and the presence of an external CO2 stream (in addition to the CO2 content of the feedstock). Therefore, it is interesting to investigate the

Fig. 5. Effect of feedstock composition (Stream 1) on the external CO2 stream flowrate [kmol/h] (a) and on the XCO2 in the reactor (b). P = 60 bar; H2/COx=3, Feedstock flow rate (Stream 1) = 1kmol/s, GHSV=6000 h−1, the values of the other process parameters reported in Table 2.

40

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Fig. 7. Sensitivity Analysis of Cascade Distillation: effect of the reflux ratio and distillate recovery on the DME purity in the residue; the values of the other process parameters reported in the Tables.

Table 4 Material balance of the DRL-MR: Main streams composition and flow-rate with respect to the process flow diagram of Fig. 2. Simulation realized at H2/COx = 3 and process parameters reported in Tables 2 and 3. STREAM

1

5

17

2

3

6

9

10

11

8

P [bar] T [K] yDME yCH3OH yH2O yCO yCO2 yH2 Mole flow [kmol/h]

60 523 0 0 0 0.05 0.15 0.80 3600

60 270 0.013 2.5 10−5 3.2 10−5 0.037 0.22 0.73 1305

60 676 0 0 0.02 0 0.30 0.68 2632

60 551.6 0.002 4.5·10−4 0.011 0.029 0.217 0.741 7537

60 501 0.371 0.015 0.014 0.033 0.419 0.148 2520

60 270 0.458 0.019 0.018 0.017 0.463 0.025 1215

60 270 0.003 0 0 0.013 0.983 0.001 523

60 270 10−4 0 0 0.151 0.526 0.323 92

60 407 0.996 0.003 10−5 0 9.9· 10−4 0 552

60 485 0.094 0.451 0.455 0 3·10−6 0 48

Fig. 8. Sensitivity Analysis of Cascade Distillation: effect of the reflux ratio and distillate on the reboiler duty; the values of the other process parameters reported in the Tables.

41

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dioxide enters the reactor through the recirculation from the treatment of the sweeping gas. The management of the overall process simulation required the use of a complex algorithm to solve the models of several interconnected sub-units with the help of three integrated commercial softwares. For this reason and for the purpose of seeking greater clarity, some simplified hypotheses were taken and discussed. Such hypotheses will subsequently be removed to aim at a economical feasibility of the process that will account for the estimation of the cost of investment. In conclusion, the mathematical model of the overall process enabled the process analysis and the sensitivity analysis of the main parameters’ effect. The algorithm, implemented to carry out the sensitivity analysis of the process, is presented and applied to investigate the effect of the main process parameters and to perform the basic design of the separation and recirculation units. The results show how it is possible to obtain 60% CO2 conversion and a 60% YDME by implementing a double recycling loop and an external feed of pure CO2. Our simulations show that the overall process in the DRL_DME configuration could reproduce the reactor performances working at high CO2 content as suggested by the recent literature [33,34,39]. These results highlight the relevance of further research and development on zeolite membranes with improved H2O permselectivity as one of the major challenges in processes for CO2 capture and transformation into DME. Other major outcomes of the sensitivity analysis are that:

Fig. 9. Effect of the H2/COx in the MR (Stream 2) on the CO2 Conversion in both the plant and the mere reactor. Feedstock composition yH2 = 0.8; yCO2 = 0.15, feedstock flow rate (Stream 1) = 1kmol/s, GHSV=6000 h−1, the values of the other process parameters reported in Tables 2 and 3.

• Beside the high CO

2 content of the stream, a portion of the separated CO2 can be utilized in the process as external feed to adjust the seeping gas flow rate to award the process high flexibility; having demonstrated this makes it feasible a process that uses CO2 as a sweeping gas and obtains the desired product DME and the remaining CO2 by-product as a high-pressure liquid.

• Double recycling loop makes the actual whole plant very flexible

•

Fig. 10. Effect of the H2/COx in the MR (Stream 2) on the production yield of DME in both the plant and the mere reactor. Feedstock composition yH2 = 0.8; yCO2 = 0.15, feedstock flow rate (Stream 1) = 1kmol/s, GHSV=6000 h−1, the values of the other process parameters reported in Tables 2 and 3.

synthesis performed in a packed bed catalytic membrane reactor incorporating zeolite membrane for H2O removal, enhance the process performances of the conventional process, enables the CO2 valorisation and furthermore is practically feasible in the described DLR plant (Fig. 10).

and enhance the process performances with respect to the MR alone. This feature should be further investigated by extending the variation of the feedstock composition and crossing it with the variation of the sweeping gas flow rate. In this first paper the feedstock composition has been chosen to have a large variability of H2/COx in the MR at fixed SW. The interconnected effect between different process parameters fixed in this paper (e.g. sweeping gas flow rate and inlet MR temperature, permeability of each component) will be investigated in next works to further explore the process feasibility. From the energetic point of view, it appears that the process is selfsufficient (from the thermal point of view) because the energy content of the produced steam is higher than the heat required by the distillation units. Apart from the head losses (not considered in this ideal simulation), all the equipment works at high pressure and the compression costs are minimized. The process shows a low energy footprint, the greatest impact on production costs is given by the equipment costs that includes expensive materials and some complicated processes (PSA) and will be object of future investigations (e.g. implementation of separation units alternative to the PSA). Further research should also include the optimization of the plant configuration with different feedstock and a careful economic evaluation by including the estimation of the capital and operative costs.

References

5. Conclusions

[1] R. Rohde, R.A. Muller, R. Jacobsen, E. Muller, S. Perlmutter, A. Rosenfeld, J. Wurtele, D. Groom, C. Wickham, A new estimate of the average earth surface land temperature spanning 1753 to 2011, Geoinfor Geostat: An Overview 1 (2013) 1–7. [2] https://www.epa.gov/cleanpowerplan/carbon-pollution-standards-new-modifiedand-reconstructed-power-plants. [3] P.J. Cook, Clean Energy, Climate and Carbon, CRC Press, 2012.

In this paper, the proof of concept of the innovative Double Recycling Loop for DME production has been presented. This process requires the use of unreacted recirculation and an external CO2 current to control the input composition. This external input of pure carbon 42

Journal of CO₂ Utilization 22 (2017) 33–43

M. De Falco et al. [4] K. Michael, A. Golab, V. Shulakova, J. Ennis-King, G. Allinson, S. Sharma, T. Aiken, Geological storage of CO2 in saline aquifers − A review of the experience from existing storage operations, Int. J. Greenhouse Gas Control 4 (2010) 659–667. [5] C. Gough, State of the art in carbon dioxide capture and storage in the UK: An experts’ review, Int. J. Greenhouse Gas Control 2 (2008) 155–168. [6] K. Goto, K. Yogo, T. Higashii, A review of efficiency penalty in a coal-fired power plant with post-combustion CO2 capture, Appl. Energy 111 (2013) 710–720. [7] Hyungwoong Ahn, Mauro Luberti, Zhengyi Liu, Stefano Brandani, Process configuration studies of the amine capture process for coal-fired power plantsI nternational, J. Greenhouse Gas Control 16 (2013) 29–40. [8] D. Jose ́ Figueroa, Timothy Fout, Sean Plasynski, Howard McIlvried, D. Rameshwar Srivastava, Advances in CO2capture technologýThe U. S. department of energy—s carbon sequestration program, Int. J. Greenhouse Gas Control 2 (2008) 9–20. [9] Centi Siglinda Perathoner Opportunities and prospects in the chemical recycling of carbon dioxide to fuels, Catal. Today 148 (2009) 191–205. [10] G.A. Olah, A. Goeppert, G.K. Surya Prakash, Chemical recycling of carbon dioxide to methanol and dimethyl ether: from greenhouse gas to renewable, environmentally carbon neutral fuels and synthetic hydrocarbons, J. Org. Chem. 74 (2009) 487–498. [11] J. Ma, N. Sun, X. Zhang, N. Zhao, F. Xiao, W. Wei, Y. Sun, A short review of catalysis for CO2conversion, Catal. Today 148 (2009) 221–231. [12] G. Centi, M. De Falco, G. Iaquaniello, S. Perathoner, Strategy and drivers for CO2 (Re)use, in: M. De Falco, G. Iaquaniello, G. Centi (Eds.), CO2: A Valuable Source of Carbon, Green Energy and Technology, Springer-Verlag, London, 2013. [13] J. Guo, H. Lou, H. Zhao, D. Chai, X. Zheng, Dry reforming of methane over nickel catalysts supported on magnesium aluminate spinels, Appl. Catal. A Gen. 273 (2004) 75–82. [14] A.G. Haiduc, M. Brandenberger, S. Suquet, F. Vogel, R. Bernier-Latmani, C. Ludwig, SunCHem: an integrated process for the hydrothermal production of methane from microalgae and CO2 mitigation, J. Appl. Phys. 21 (2009) 529–541. [15] V. Barbarossa, C. Bassano, P. Deiana, G. Vanga, CO2 Conversion to CH4, in M. De Falco, G., Iaquaniello, G., Centi, CO2: A Valuable Source of Carbon, Green Energy and Technology, Springer-Verlag, London, 2013. [16] L. Mennicken, A. Janz, S. Roth, The german R & D program for CO2 Utilization—Innovations for a green economy, Environ. Sci. Pollut. Res. 23 (2016) 11386–11392. [17] S.W. Kang, J.W. Bae, K.W. Jun, H.S. Potdar, Dimethyl ether synthesis from syngasover the composite catalysts of Cu–ZnO–Al2O3/Zr-modified zeolites, Catal. Commun. 9 (2008) 2035–2039. [18] T. Chmielniak, M. Sciazko, Co-gasification of biomass and coal for methanol synthesis, Appl. Energy 74 (2003) 393–403. [19] C. Arcoumanis, C. Bae, R. Crookes, E. Kinoshita, The potential of di-methyl ether (DME) as an alternative fuel for compression-ignition engines: a review, Fuel 87 (2008) 1014–1030. [20] W. Ying, Z. Longbao, W. Hewu, Diesel emission improvements by the use of oxygenated DME/diesel blend fuels Atmosph, Environ 40 (2006) 2313–2320. [21] G.A. Olah, G.K. Goeppert, Surya Prakash Beyond Oil and Gas: The Methanol Economy, (2017). [22] Q. Li, G. Wu, C.M. Johnston, P. Zelenay, Zelenay direct dimethyl ether fuel cell with much improved performance, Electrocatalysis 5 (July (3)) (2014) 310–317. [23] P. Pérez-Uriarte, M. Gamero, A. Ateka, M. Díaz, A.T. Aguayo, J. Bilbao, Effect of the Acidity of HZSM5 Zeolite and the Binder in the DMETransformation to Olefins, Ind. Eng. Chem. Res. 55 (6) (2016). [24] P. Pérez-Uriarte, A. Ateka, A.T. Aguayo, J. Bilbao, Comparison of HZSM-5 zeolite and SAPO (-18 and –34) based catalysts for the production of light olefins from DME, Catal. Lett. (2016), http://dx.doi.org/10.1007/s10562-016-1829-z. [25] Z. Azizi, Mo. Rezaeimanesh, T. Tohidian, M. Reza Rahimpour, Dimethyl ether: a review of technologies and production challenges, Chem. Eng. Process. 82 (2014) 150–172. [26] F.S. Ramos, A.M. Duarte de Farias, L.E.P. Borges, J.L. Monteiro, M.A. Fraga,

[27]

[28]

[29]

[30]

[31]

[32]

[33]

[34]

[35]

[36]

[37]

[38]

[39]

[40]

[41]

[42]

[43]

[44]

[45]

43

E.F. Sousa-Aguiar, L.G. Appel, Role of dehydration catalyst acid properties on onestep DME synthesis over physical mixtures, Cat. Today 101 (2005) 39–44. W.Z. Lu, L.H. Teng, W.D. Xiao, Simulation and experiment studyof dimethyl ether synthesis from syngas in a fluidized-bedreactor, Chem. Eng. Sci. 59 (5455) (2004) e64. L. Wang, D. Fang, X. Huang, S. Zhang, Y. Qi, Z. Liu, Influence of reaction conditions on methanol synthesis and WGS reaction in the syngas-to-DME process, J. Nat. Gas Chem. 15 (2006) 38–44. M. De Falco, M. Capocelli, G. Centi, Dimethyl ether production from CO2 rich feedstocks in a one-step process: thermodynamic evaluation and reactor simulation, Chem. Eng. J. 294 (2016) 400–409. W.H. Chen, B.J. Lin, H.W. Lee, M.H. Huang, One-step synthesis of dimethyl ether from the gas mixture containing CO2 with high space velocity, Appl. Energy 98 (2012) 92–101. M.P. Rohde, G. Schaub, S. Khajavi, J.C. Jansen, F. Kapteijn, Fischer-Tropsch synthesis with in situ H2O removal − Directions of membrane development, Micropor. Mesopor. Mater. 115 (2008) 123–136. I. Iliuta, F. Larachi, P. Fongarland, Dimethyl ether synthesis with in situ H2O removal in fixed-Bed membrane reactor: model and simulations, Ind. Eng. Chem. Res. 49 (2010) 6870–6877. N. Diban, A.M. Urtiaga, I. Ortiz, J. Ereña, J. Bilbao, A.T. Aguayo, Influence of the membrane properties on the catalytic production of dimethyl ether with in situ water removal for the successful capture of CO2, Chem. Eng. J. 234 (2013) 140–148. N. Diban, A.M. Urtiaga, I. Ortiz, J. Ereña, J. Bilbao, A.T. Aguayo, Improved performance of a pbm reactor for simultaneous CO2 capture and DME synthesis, Ind. Eng. Chem. Res. 53 (2014) 19479–19487. C. Tellez, M. Menéndez, A. Basile, A.F. Gallucci, Zeolite membrane reactors, Membranes for Membrane Reactors: Preparation, Optimization and Selection, John Wiley & Sons, Singapore, 2011, pp. 243–273. V.V. Volkov, E.G. Novitskii, G.A. Dibrov, P.V. Samokhin, M.A. Kipnis, A.B. Yaroslavtsev, Catalytic conversion of methanol to dimethyl ether on polymer/ ceramic composite membranes, Catal. Today 193 (2012) 31–36. C. Zhou, N. Wang, Y. Qian, X. Liu, J. Caro, A. Huang, Efficient synthesis of dimethyl ether from methanol in a bifunctional zeolite membrane reactor Angew.Chem, Int. Ed. 55 (2016) 12678–12682. D.A. Fedosov, A.V. Smirnov, V.V. Shkirskiy, T. Voskoboynikov, I.I. Ivanova, Methanol dehydration in NaA zeolite membrane reactor, J. Membr. Sci. 486 (2015) 189–194. M. De Falco, M. Capocelli, A. Basile, Selective membrane application for the industrial one-step DME production process fed by CO2 rich streams: modeling and simulation, Int. J. Hy. En. (2017) (in press. K.M.V. Bussche, G.F. Froment, A steady-State kinetic model for methanol synthesis and the water gas shift reaction on a commercial Cu/ZnO/Al2O3 catalyst, J. Catal. 161 (1996) 1–10. A. Ateka, P. Perez-Uriarte, M. Gamero, J. Erena, A.T. Aguayo, J. Bilbao, A comparative thermodynamic study on the CO2 conversion in the synthesis of methanol and of DME, Energy 120 (2017) 796–804. R.G. Herman, K. Klier, G.W. Simmons, B.P. Finn, J.B. Bulko, T.P. Kobylinski, Catalytic synthesis of methanol from CO/H2: I. Phase composition, electronic properties, and activities of the Cu/ZnO/M2O3 catalysts, J. Catal. 56 (1979) 407–429. M.P. Bernal, E. Piera, J. Coronas, M. Menéndez, J. Santamarìa, Mordenite and ZSM5 hydrophilic tubular membranes for the separation of gas phase mixtures, Catal. Today 56 (2000) 221–227. H.J. Song, H.T. Zhang, W.Y. Ying, D.Y. Fang, Measurement and calculation of vapor-liquid equilibrium for dimethyl ether-methanol-water ternary system, J. Chem. Ind. Eng. 30 (2005) 67–71. J. Fontalvo, Using user models in Matlab® within the Aspen Plus® Interface with an Excel®, Ingenierìa e Invest. 4 (2) (2014) 39–43.