Advanced reactor and catalyst design for conventional and sorption enhanced direct dimethyl ether synthesis

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PhD Thesis of Simone Guffanti

Matriculation Roll 880803

Department of Chemistry, Materials and Chemical Engineering “Giulio Natta” PhD in Industrial Chemistry and Chemical Engineering

XXXII Cycle 2016-2019

Advanced reactor and catalyst design for conventional

and sorption enhanced direct dimethyl ether synthesis

Coordinator: Prof. Alessio Frassoldati

Tutor: Prof. Enrico Tronconi

Supervisor: Prof. Gianpiero Groppi

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Abstract

Dimethyl ether (DME) is a commodity mostly used as propellant for aerosol spray and a promising alternative fuel that can be synthetized from syngas obtained from renewable sources as biomass. DME can be produced through two routes: a two-step process consisting in the methanol synthesis followed by the methanol dehydration to DME (indirect synthesis), or a one-step process in which DME is obtained from the syngas coupling the two steps in a single reactor (direct synthesis). Differently from the indirect synthesis, the DME direct synthesis is not a fully established industrial process and it is a subject of research with focus on catalyst development and reactor and process design. Different issues are indeed associated to the direct DME synthesis, including: i) the difficult thermal control; ii) the coupling of methanol synthesis and dehydration catalyst; iii) the thermodynamic and kinetics limitations with rich CO2 syngas feeds.

i) The direct synthesis of DME from syngas is an exothermic process: a correct reactor design is fundamental for a proper thermal control in order to prevent excessive hot-spot temperature with consequent catalyst deactivation. ii) The direct DME synthesis requires to properly coupling two different catalyst functions in the same reactor: methanol synthesis Cu/ZnO/Al2O3 catalyst (CZA) and methanol dehydration to DME catalyst

(γ-alumina or zeolites). iii) The presence of large amounts of CO2 in the feed (as in the case

of biomass gasification) leads to a large production of water, that hinders both the kinetic and the thermodynamic of the process. Sorption enhanced DME synthesis (SEDMES) is a promising concept to solve some of these issues. SEDMES consists in the coupling of direct DME synthesis with in-situ water adsorption, that shifts the thermodynamic equilibrium limitations with a consequent improvement of reactant conversion and DME yield.

The aim of this thesis is the is the development of mathematical models of chemical reactors for the direct synthesis of DME from biomass gasification syngas, to be used for the simulation and design of industrial scale equipment. The models provide an accurate description of all the significative chemical and physical phenomena occurring inside of the reactors, with particular focus on the thermal behavior, which plays a key role in a correct reactor design. The thesis work includes both a model of a conventional multitubular externally cooled fixed bed reactor and a model of an advanced SEDMES reactor. After a preliminary thermodynamic analysis of the SEDMES process, the

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conventional reactor model is used for the simulation and design of the converter unit. The effect of active phase distribution in catalyst pellets is analyzed: the two catalyst functions can be intimately mixed in hybrid pellets, located on separated pellets or coupled in core@shell engineered pellets. It is shown that the different spatial distribution of the active phases has a drastic impact on reactor performances and temperature profiles.

Finally, the SEDMES reactor model, validated by comparison with experimental bench scale data, is used to describe the dynamic profiles of composition and temperature inside the reactor during the reaction/adsorption step, allowing for a rational design of the unit.

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List of publications

• S. Guffanti, C.G. Visconti, G. Groppi, On the effect of active phase distribution at the catalyst pellet scale on the direct synthesis of dimethyl ether – Submitted to Industrial & Engineering Chemistry Research

• S. Guffanti, C.G. Visconti, J. van Kampen, J. Boon, G. Groppi, Reactor modelling and design for sorption enhanced dimethyl ether synthesis – Submitted to Chemical Engineering Journal

• Techno-economic analysis of flexible Power&Biomass-to-Methanol plants – In preparation

• S. Guffanti, C.G. Visconti, G. Groppi, GRICU PhD National School 2017, Palermo (Italy). “Sorption enhanced dimethyl ether synthesis: reactor modelling and simulation” – Poster presentation

• S. Guffanti, C.G. Visconti, G. Groppi, 25th_{International Conference on Chemical}

Reaction Engineering (ISCRE25) (May 2018) Firenze (Italy): “Sorption enhanced dimethyl ether synthesis: reactor modelling and simulation” – Poster presentation • S. Guffanti, C.G. Visconti, G. Groppi, XX Congresso Nazionale di Catalisi – XX

Congresso Nazionale della Divisione di Chimica Iindustriale (GIC – DiChIn2018) (September 2018) Milano (Italy): “The effects of intraparticle diffusion phenomena on dimethyl ether direct synthesis” – Poster presentation

• S. Guffanti, C.G. Visconti, G. Groppi, XXIII International Conference on Chemical Reactors (CHEMREACTOR-23) (November 2018) Ghent (Belgium): “The effects of intraparticle diffusion phenomena on dimethyl ether direct synthesis” – Oral presentation

• S. Guffanti, C.G. Visconti, G. Groppi, 12th_{Natural Gas Conversion Symposium}

(NGCS12) (June 2019) San Antonio (Texas – USA): “The effects of intraparticle diffusion phenomena on dimethyl ether direct synthesis” – Oral presentation • S. Guffanti, C.G. Visconti, J. van Kampen, J. Boon, G. Groppi XXIV International

Conference on Chemical Reactors (CHEMREACTOR-24) (September 2020) Milano (Italy): “Reactor modelling and design for sorption enhanced dimethyl ether synthesis” – Accepted oral presentation

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List of figures

Figure 1. Final use of methanol in year 2015 [14]. ... 4 Figure 2. DME indirect synthesis process. ... 5 Figure 3. DME direct synthesis process. ... 8 Figure 4. dimethyl ether synthesis publications: a) subject of the publication; b) synthesis method [21]. ... 10 Figure 5. Graphical representation of the different catalyst configurations for DME direct synthesis. Colors: auburn - methanol catalyst; grey - dehydration catalyst. ... 12 Figure 6. SEDMES process. ... 15 Figure 7. FLEDGED project logo. ... 17 Figure 8. Graphical representation of the different catalyst configurations for DME direct synthesis. Colors: auburn - methanol catalyst; grey - dehydration catalyst. ... 24 Figure 9. Axial profiles of the average specific molar flow rate of a) reactants (CO – full line; CO2 – dashed line) and b) products (MeOH – full line, DME – dashed line); (Hybrid

pellet – blue lines; mechanical mixture – red lines; MeOH@DME – green lines;

DME@MeOH – orange lines). ... 25 Figure 10. DME carbon yield profiles; (Hybrid pellet – blue line; mechanical mixture – red line; MeOH@DME – green line; DME@MeOH – orange line; equilibrium yield – dashed line). ... 27 Figure 11. Catalyst temperature profile on the reactor centerline; (Hybrid pellet – blue line; mechanical mixture – red line; MeOH@DME – green line; DME@MeOH – orange line; coolant – dashed line). ... 28 Figure 12. Effect of MeOH/DME catalyst ratio on DME yield using different catalyst configurations (Hybrid pellet – blue lines; mechanical mixture – red lines; MeOH@DME – green lines; DME@MeOH – orange lines). ... 29 Figure 13. Effect of MeOH/DME catalyst ratio on catalyst centerline temperature hot-spot using different catalyst configurations (Hybrid pellet – blue lines; mechanical mixture – red lines; MeOH@DME – green lines; DME@MeOH – orange lines). ... 29 Figure 14. Time evolution of outlet molar fraction experimental (dots) vs. model (lines). ... 32 Figure 15. Envelope of local maximum temperatures. Experimental (dots) vs. model (line). ... 32 Figure 16. Time evolution of outlet DME flowrate normalized with respect to inlet carbon flow rate at different feed CO/CO2 ratios. ... 34

Figure 17. Axial profile of envelope of maximum gas centerline local temperatures at different feed CO/CO2 ratios. ... 35

Figure 18. Axial profile of maximum gas centerline local temperatures with different tube diameters. ... 36 Figure 19. a) Carbon and b) Cold gas efficiency yield at different ratio α and module M. ... 40 Figure 20. SEDMES thermodynamic equilibrium yield at different pressure; a) carbon yield; b) cold gas efficiency yield. ... 41 Figure 21. SEDMES thermodynamic equilibrium yield at different temperature; a) cold gas efficiency yield; b) carbon yield. ... 42

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Figure 22. Dimethyl ether production plant designed in FLEDGED project. ... 43 Figure 23. Schematic representation of SEDMES reaction train. ... 44 Figure 24. DME obtained from the conventional reactor. ... 45 Figure 25. Heat managed by the conventional converter and water splitter in the

SEDMES reaction train at different temperatures. ... 46 Figure 26. Water removed by the water splitter in SEDMES reaction train at different equilibrium temperatures. ... 47 Figure 27. Effect of GHSV with mechanical mixture of different catalyst pellets on axial profile of a) DME carbon yield and b) centerline catalyst temperature. ... 51 Figure 28. Effect of GHSV on axial profile of pressure drops. ... 52 Figure 29. Effect of tube internal diameter with mechanical mixture of different catalyst pellets on axial profile of centerline catalyst temperature. ... 53 Figure 30. Effect of α with mechanical mixture of different catalyst pellets on axial

profile of a) DME carbon yield and b) centerline catalyst temperature. ... 54 Figure 31. Effect of tube internal diameter with mechanical mixture of different catalyst pellets on axial profile of centerline catalyst temperature with α=0.66. ... 54 Figure 32. Effect of GHSV with hybrid pellets on axial profile of a) DME carbon yield and b) centerline catalyst temperature. ... 55 Figure 33. Effect of tube internal diameter with hybrid pellets on axial profile of a)

centerline catalyst temperature and b) DME carbon yield. ... 56 Figure 34. Effect of α with hybrid pellets on axial profile of a) DME carbon yield and b) centerline catalyst temperature. ... 57 Figure 35. Hybrid (left) and mechanical mixture (right) configuration sketches. ... 70 Figure 36. Axial profiles of the average specific molar flow rate of a) reactants and b) products with hybrid pellet (h) and mechanical mixture (m) configurations. ... 72 Figure 37. DME carbon yield profiles with hybrid pellet and mechanical mixture

configurations. ... 73 Figure 38. a) Catalyst 2D temperature profile in hybrid pellet configuration; b) catalyst (hybrid, MeOH and DME catalyst in mechanical mixture) temperature profile on the tube centerline. ... 74 Figure 39. Catalyst pellet effectiveness factor centreline profiles for a) methanol synthesis b) methanol dehydration with hybrid pellet and mechanical mixture configurations. ... 75 Figure 40. Reaction rate profiles inside catalyst pellets: a) methanol synthesis and b) methanol dehydration. Surface conditions: 50 bar, 543 K; surface composition: CO 20%, CO2 17.5%, H2 52.3%, H2O 0.5%, MeOH 0.8%, DME 0.6%, CH4 8.3%. ... 77

Figure 41. a) Methanol and b) water concentration profiles inside the catalyst pellet (hybrid, pure dehydration catalyst and diluted dehydration catalyst). Surface conditions: 50 bar, 543 K; surface composition: CO 20%, CO2 17.5%, H2 52.3%, H2O 0.5%, MeOH

0.8%, DME 0.6%, CH4 8.3%. ... 78

Figure 42. Core@shell pellets sketch (brown – MeOH catalyst; grey - DME catalyst). ... 79 Figure 43. Axial profiles of the average specific molar flow rate of a) reactants and b) products with different catalyst configurations. ... 80 Figure 44. DME carbon yield profiles with different catalyst configurations ... 80 Figure 45. Catalyst temperature profile on the tube centerline with different catalyst configurations. ... 81

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Figure 46. DME carbon yield profile along the reactor axial coordinate with different catalyst configuration using a feed with M=2 and CO/CO2=1. ... 83

Figure 47. Catalyst temperature profile on the tube centerline with different catalyst configuration using a feed with M=2 and CO/CO2=1. ... 84

Figure 48. Effect of pressure on a) DME yield; b) COx conversion; c) DME carbon

selectivity using different catalyst configurations. ... 85 Figure 49. Effect of MeOH/DME catalyst ratio on a) DME yield; b) COx conversion; c)

DME carbon selectivity using different catalyst configurations. ... 87 Figure 50. Effect of MeOH/DME catalyst ratio on catalyst centreline temperature hot-spot using different catalyst configurations. ... 88 Figure 51. Time evolution of outlet molar fraction experimental (dots) vs. model (lines). ... 113 Figure 52. Time evolution of outlet DME specific flow rate at different axial positions. ... 114 Figure 53. Time evolution adsorbed water load at different axial positions. ... 114 Figure 54. Time evolution of centerline gas temperature at different reactor coordinates. Experimental (dots) vs. model (lines). ... 115 Figure 55. Envelope of local maximum temperatures. Experimental (dots) vs. model (line). ... 116 Figure 56. Envelope of local maximum reaction/adsorption heat release. ... 116 Figure 57. Time evolution of the dimensionless flowrate for carbon containing reactants at different CO/CO2 feed ratios. a) CO (full lines) and CO2 (dashed lines) flow rate out/in

ratios; b) Total COx flow rate out/in ratio. ... 119

Figure 58. Axial profile of the average (on cross section area) adsorbent water load

profile at time 3600 s for different feed CO/CO2 ratios. ... 120

Figure 59. Time evolution of outlet product flowrate per unit area at different CO/CO2

ratios. a) DME outlet flowrate; b) methanol (full lines) and water (dashed lines) flowrates. ... 121 Figure 60. Time evolution of outlet DME flowrate normalized with respect to inlet carbon flow rate at different feed CO/CO2 ratios. ... 122

Figure 61. Axial profile of envelope of maximum gas centerline local temperatures at different feed CO/CO2 ratios. ... 123

Figure 62. Axial profile of maximum gas centerline local temperatures with different tube diameters. ... 125 Figure 63. Time evolution of outlet DME flow rate per unit area with different tube diameters. ... 126

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List of tables

Table 1. Physical and chemical properties of different fuels. ... 2

Table 2. Diffusion lengths (V/S [mm]) of methanol synthesis and methanol dehydration reaction in the different catalyst configurations. ... 26

Table 3. DME carbon yield at different CO/CO2 feed ratios. ... 34

Table 4. Base case input variables. ... 49

Table 5. Base case inlet feed composition. ... 50

Table 6. Inlet feed composition. ... 53

Table 7. Input variables used in the simulations. ... 71

Table 8. Inlet gas composition used as reference case. ... 71

Table 10. Diffusion lengths (V/S [mm]) of methanol synthesis and methanol dehydration reaction in the different catalyst configurations. ... 82

Table 10. Physical properties of solid phases. ... 110

Table 11. Geometrical parameters and operating conditions of the reactor tube. ... 111

Table 12. Inlet feed composition in model validation. ... 111

Table 13. Inlet feed composition in industrial scale reactor analysis. ... 118

Table 14. DME carbon yield at different CO/CO2 feed ratios. ... 123

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1 Introduction

1.1 Dimethyl ether as alternative fuel

Dimethyl ether (DME or CH3OCH3) is the simplest ether, a molecule consisting of an

oxygen atom bonded to a couple of methyl groups. DME is an environmentally non-hazardous volatile compound since it is not toxic, carcinogenic, mutagenic or teratogenic. The atmospheric lifetime of DME is around 5 days and its Global Warming Potential (GWP) referred to CO2 is 1.2 on 20 years horizon, 0.3 on 100-years horizon and 0.1 on

500-years horizon [1] moreover, differently from many other Volatile Organic Compounds (VOCs), DME is not ozone-depleting: in reason of these environmentally benign properties, DME is a commodity mostly used as propellant for aerosol spray and refrigerant. Its main industrial application is the conversion to dimethyl sulfate by treatment with sulfur trioxide [2], furthermore DME, in reaction with carbon monoxide and water, can substitute methanol in the synthesis of acetic acid [2] and it is also considered as a possible chemical intermediate in the production of olefins (ethylene, propene and butenes) and gasoline [2,3]. In the last decades there has been a growing interest in the synthesis of DME for energetic uses, specifically: i) DME as alternative to Liquefied Petroleum Gas (LPG), in household and industrial applications [4–7]; ii) DME as alternative diesel fuel for compression-ignition (CI) engines [5,8]; iii) DME as fuel for power generation in gas turbines [4,5,9]; iv) DME as hydrogen carrier for fuel cells [4,5,10,11].

1.1.1 DME as alternative to LPG

Vapor-liquid equilibrium properties of DME are similar to LPG, whose main constituents are propane (C3H8) and butane (C4H10). A comparison of the physical properties of propane,

butane, methanol, diesel and DME is reported in Table 1. DME is gaseous in normal conditions (T = 20°C; P = 1 atm), but it is easily liquified at -24.8°C and atmospheric pressure or at 5.1 bar and 20°C temperature. Therefore, it is possible to blend DME and LPG and adopt the existent infrastructure used for LPG for the storage and delivery. Only a change in the sealing materials is required since DME has a solvent effect that can damage organic materials such as gum and plastics. Besides, some little changes are also necessary

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for the commercial household burners, as showed in the work of Marchionna et al. [6] that tested different DME/LPG mixtures, concluding that i) the nozzles diameter has to be increased, due to the DME smaller heat value per weight (⁓0.6 times of that of propane, see Table 1), and ii) the air/fuel premixing ratio must be changed in order to improve flame stability. They also demonstrated that DME/LPG mixtures with 15-20% DME offer better performances than pure DME, but they did not excluded that blends with >20% DME can be employed in the case of other application in which DME is used as substitute to LPG or make-up. Accordingly, although minor changes in the present infrastructure are required, DME has a huge potential as alternative to LPG in particular in non-petroleum producing Asian countries like China and Japan [12].

References [5,8,13] Table 1. Physical and chemical properties of different fuels.

1.1.2 DME as alternative to diesel

Methanol has a high octane number that makes it suitable for sparkle-ignition (SI) engines, DME instead has a cetane number of 55-60 and low auto-ignition temperature of 235°C, properties that make it an excellent diesel fuel for CI engines (Table 1) [5]. The low boiling point of DME is also a desired quality, allowing for fast evaporation when the liquid fuel is injected inside the engine. Moreover, DME showed better performances in terms of pollutant emissions with respect to the standard diesel fuel [8]. DME is an oxygenated compound (35% w/w oxygen content) without C-C bonds, that leads to a zero-soot combustion implying that a particulate matter filter is not required. Besides under

Propane (LPG) Butane (LPG) Methanol Diesel Dimethyl ether

Chemical formula C3H8 C4H10 CH3OH C10-C15 CH3OCH3

Boiling point [°C] -42.1 -0.4 64 125-240 -24.8 Ignition point [°C] 504 365 385 210 235 Liquid density [kg/m3_] ₅₀₁ ₅₇₃ ₇₉₂ ₈₃₁ ₆₆₇ Liquid viscosity [m2_/s] _0.3·10-6 _0.26·10-6 _0.55·10-6 _3·10-6 _0.22·10-6 LHV [MJ/kg] 46.4 45.8 20 42.5 27.6 Octane number 103 91 123 25 - Cetane number - - - 40-55 55-60

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optimized conditions, the combustion of DME in CI engines produces less NOx respect to

diesel, which was attributed to the lower peak of combustion temperature, while the emission levels of unburned hydrocarbons (HC) are aligned with the standard diesel ones and CO content in effluent gasses is equal or slightly higher. Finally, the combustion of DME is completely sulfur-free. However, DME presents some disadvantages with respect to diesel fuel: i) the small Lower Heating Value (LHV) (Table 1); ii) the high volatility of DME compared to diesel; iii) the low viscosity of DME. i) The most evident disadvantage is related to the low combustion enthalpy, DME is indeed an oxygenated molecule and a smaller amount of energy is stored in the C-O bonds respect to the C-C and C-H bonds. As consequence a larger injected volume of fuel and a longer time of injection is required when DME is employed; ii) the DME is gaseous in the same storage conditions of diesel, a modification of standard tank used in vehicles is therefore necessary; iii) DME has a lower viscosity than standard diesel, this characteristic may cause leakage from the fuel supply system. A low lubricity can cause intensified surface wear of moving parts within the fuel-injection system, requiring more maintenance Anyway, it is possible to overcame this issue using lubricity-enhancing additives [8]. In spite of these drawbacks DME is accounted as one of the most promising alternative to diesel as transportation fuel, thanks to the low emissions during the combustion and the possibility to use both fossils and renewable feed stocks for its production [8].

1.1.3 DME as fuel for power generation

The applicability of DME in low-NOx emission gas turbines for power generation has been

proved [9]. It has been shown that DME can burn in the same turbines using methane as fuel, with little modifications. The pollutant emissions are at the same level of methane or slightly higher in the case of CO.

1.1.4 DME as hydrogen carrier for fuel cells

Fuel cells efficiently convert chemical energy directly into electrical energy. The fuel usually employed in fuel cell anode is molecular hydrogen (H2), while air is fed to the

cathode. The storage of pure H2 represent one of the main challenges. The chemical storage

is one of the preferred options, DME is considered as promising candidates. Hydrogen can be obtained from the DME via steam reforming, partial oxidation or autothermal reforming. DME can be reformed at lower temperature than hydrocarbons but, as showed in the

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thermodynamic study of Semelsberger et al. [10], it can decompose to methane. The selectivity to H2 can be enhanced using a mixture of a solid acid catalyst and commercial

Cu/Zn/Al2O3 (CZA) catalyst [11], since DME decomposes passing through methanol to H2

in a two-step process. At present, methanol is the most studied hydrogen vector for fuel cells, however DME is a more than a valid alternative since it is a non-toxic substance.

1.2 State of the art of dimethyl ether production technology

The world production capacity of DME in 2012 was estimated to be 10 million metric tons per year with an actual global market of 3 million metric tons per year [12]. The global demand of DME has rapidly grown since the early 2000s: the market of its traditional applications (aerosol spray propellant, refrigerant and chemical intermediate) has not undergone drastic changes, while the demand as LPG supplement and chemical intermediate in olefins production has significantly increased, primarily in China. In the present situation the synthesis of DME is strictly correlated to the production of methanol: in 2015 the 8% of the methanol industrial output has been used for the synthesis of DME (Figure 1) [14].

Figure 1. Final use of methanol in year 2015 [14].

Almost all the DME is currently produced in a process known as DME indirect synthesis, that consists of two consecutive steps (a schematic representation is reported in Figure 2). The first step of the process consists in the conversion of the synthesis gasses (CO, H2 and

a variable amount of CO2) into methanol. The methanol synthesis is a well-established

7% 2% 2% 3% 3% 4% 8% 9% 9% 18% 27%

8%

DME Formaldehyde MTO Acetic Acid Gasoline Blending MTBE/TAME Solvents Biodiesel Methylamines Chloromethanes MMA Others

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industrial process [2], that is thermodynamically limited and requires a catalyst. The process was originally performed at high pressure (250-350 bar), in order to overcome the thermodynamic limitations, and with a sulfur-resistant ZnO/Cr2O3 catalyst active at high

temperature (320-450°C). Since the 1960s, the traditional catalyst has been substituted with the more active copper based Cu/ZnO/Al2O3 (CZA) [15] catalyst, allowing to operate the

process in milder conditions (200-300°C and 50-100 bar). On the CZA catalyst three main reactions selectively occur: the methanol synthesis from CO (1), the reverse water gas shift (rWGS) (2) and the methanol synthesis from CO2 (3).

CO + 2 H2 ↔ CH3OH ∆𝐻𝑟0= −90.5 kJ/mol (1) CO2+ H2 ↔ CO + H2O ∆𝐻𝑟0 = +41.1 kJ/mol (2) CO2+ 3 H2 ↔ CH3OH + H2O ∆𝐻𝑟0= −49.4 kJ/mol (3) CO + 2H2 CH3OH CO2 + H2 CO + H2O CO2 + 3H2 CH3OH + H2O Feed CO, CO2, H2 Methanol reactor Recycle CO, CO2, H2 CH3OH, H2O 2CH3OH CH3OCH3 + H2O H2O CH3OH DME reactor H2O CH3OCH3 To recycle CH3OH Recycle CH3OH CH3OH, H2O Purge

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The feed composition is determined by the stoichiometric module usually known as ‘M module’ (4):

𝑀 = H2− CO2 CO + 𝐶𝑂2

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The module for a perfectly stoichiometric process is 2, however a slightly higher M (with a value between 2.06-2.1) are used in the syngas feed in the industrial practice [16]. The effective module M at the reactor inlet is significantly higher, considering that the process is thermodynamically limited and a recycle is needed in order to increase the process yield [17].

The methanol synthesis (as shown in stoichiometries (1-3)) is an exothermic process, therefore temperature control is fundamental in order to maximize syngas conversion and to prevent the sintering of Cu clusters on CZA catalyst, which is the main deactivation mechanism [18]. Different reactor solutions have been selected in order to minimize the investment costs, based on: i) the high cooling duty; ii) the minimization of pressure drops; iii) the positive economy of scale [17]. The favored reactor configurations are: i) Multistage adiabatic fixed bed reactors with intermediate cooling by external heat exchange or by fresh feed quenching; ii) Multitubular fixed bed reactors externally cooled with boiling water or with cold reactants. Multitubular fixed bed reactors are more efficient in the heat removal but have a more complex design and are usually more expensive than multistage adiabatic beds. The choice of the reactor configuration is commonly made on the basis of an economical and operational optimum of the whole process.

Whatever the type of reactor, a syngas recycle loop is needed in order to increase the methanol yield and reduce the waste of unconverted feed. The recompression of recycled gasses has a considerable impact on operating costs, consequently the pressure drops must be as lower as possible and the recycle ratio must be correctly tuned. The synthesis products, methanol and water, are separated and then the methanol is fed to the second step of DME indirect synthesis. The separation requires additional costs but clearly improves significantly the efficiency of methanol conversion to DME. Anyway, this operation is necessary since methanol is usually produced in large plants and only a part of it is employed for the synthesis of DME (Figure 1).

The second step of the indirect synthesis consists in the conversion by dehydration of methanol to DME (5).

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2 CH3OH ↔ CH3OCH3+ H2O ∆𝐻𝑟0= −23.0 kJ/molDME (5)

The process operates in the range of 300-320°C and 10-20 bar [13]. High pressure is not thermodynamically advantageous, as in the case of the methanol synthesis, since the reaction is equimolar. The dehydration reaction requires an acid catalyst function: γ-alumina, zeolites or silicoaluminophosphates (SAPO) have been employed [19]. The choice of the catalyst is related to the operating conditions: at high temperature in the presence of strongly acid catalyst (zeolites, SAPO) the selectivity to DME decreases due to the formation of olefins, paraffins and aromatics. The consequent coke deposition becomes also a possible issue [19]. On the other hand, γ-alumina is easily deactivated by water: it has been reported that γ-alumina activity decreases 12.5 times faster in presence of a methanol-water feed, with respect to a feed containing only methanol [20]. Fixed bed reactor, adiabatic or multitubular externally cooled, are used in the industrial practice. The methanol dehydration is a mildly exothermic reaction, so that the heat duty to manage is smaller than in the methanol synthesis, therefore adiabatic reactors are often preferred for their simplicity and lower costs [21].

The second possible process route is the DME direct synthesis (Figure 3). In this process, DME is directly obtained from syngas in a single reactor. The rest of the process is similar to the indirect route and consists in the simple separation of products and unconverted reactants, which are recycled to the reactor together with the methanol.

The DME direct synthesis is a simultaneous combination of methanol synthesis and dehydration, which requires to couple the two different catalyst functions in the same reactor [22]. The reactions that selectively take place on the catalysts are the same reported in equation (1-3) and (5), with the only exception of the rWGS reaction (2) that, due to the larger quantity of water present in the system (that is produced by reaction (3) and (5)) proceeds in the opposite direction as water gas shift (WGS). One of the main advantages of coupling the methanol synthesis and dehydration processes in a single process is the thermodynamic synergy between the reactions [23]: the methanol produced by (1) and (3) is consumed by the reaction (5) to form DME, meanwhile the water produced in (3) and (5) is consumed by the WGS (2). This potentially leads to an overall higher syngas per pass conversion and DME yield with respect to the indirect process, with a possible reduction of the recycle ratio and consequently lower operating costs.

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8 CO + 2H2 CH3OH CO2 + H2 CO + H2O CO2 + 3H2 CH3OH + H2O 2CH3OH CH3OCH3 + H2O Feed CO, CO2, H2 DME direct synthesis reactor Recycle CO, CO2, H2, CH3OH CH3OH, CH3OCH3, H2O Recycle CH3OH Purge H2O CH3OCH3 To recycle CH3OH CH3OH, H2O

Figure 3. DME direct synthesis process.

The resulting overall stoichiometry is a combination of the reaction in equation (1-3) and (5) that also depends on the inlet syngas composition: in the case of a feed consisting mainly of CO and H2 the process stoichiometry produce DME and CO2 according to

equation (6), while in presence of a large amount of feed CO2 the stoichiometry shift to the

the one reported in equation (7), due to the inhibited role of the WGS equilibrium.

3 CO + 3 H2 ↔ CH3OCH3+ CO2 ∆𝐻𝑟0= −246 kJ/molDME (6)

2 CO2+ 6 H2↔ CH3OCH3+ 3 H2O ∆𝐻𝑟0= −122 kJ/molDME (7)

Therefore, in the case of a syngas containing only CO the optimal H2/CO feed ratio is 1,

while in presence of feed CO2 the optimal stoichiometry H2:CO:CO2 changes according to

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combination of methanol synthesis, methanol dehydration and, for stoichiometry (6), WGS. The management of the heat duty is more complex than in the indirect process and an adequate reactor configuration is required.

At present the DME direct synthesis, differently from the indirect route, is not fully established in the industrial practice. In the open literature, only in a couple of projects, held respectively by the JFE Holdings Inc. [25–27] and by the Korea Gas Corporation (KOGAS) [28–32] the technology has been implemented on large scale.

In the JFE project a slurry phase reactor has been developed in order to control the temperature. The experimental tests have been performed at three different scale level: a bench scale reactor (50 kg/day), a pilot scale reactor (5 Tons/day) and in a demonstration plant (100 Tons/day) [26]. The operating conditions in the pilot scale reactor have been explored in the range between 240-280°C and 30-70 bar, while 260°C and 50 bar have been finally selected as optimal temperature and pressure. The reactor inlet syngas, obtained with the autothermal reforming of natural gas, has been fed with the stoichiometric ratio of H2/CO=1 [25]. The catalyst, designed for the slurry phase, has not been specified.

Multitubular fixed bed externally cooled reactors have been instead adopted in the KOGAS project. The experimental and modeling study has been performed on a pilot scale reactor (50 kg/day) [30] and on a demonstration plant (10 Tons/day) [32]. On the basis of the data obtained, a 3000 Tons/day full scale plant has been designed [32], which operates in the temperature range of 200-300°C and at 50 bar in presence of a mixture of CZA and γ-alumina catalyst. The syngas employed has an over-stochiometric H2/CO ratio of 1.4-1.5

and has been obtained in a tri-reforming reactor fed with natural gas and recycled CO2.

1.3 Dimethyl ether production challenges

The growing interest in DME as potential alternative fuel is driving the research to find new solutions. In the work of Azizi et al. [21] is reported a survey (Figure 4) on the focus of the open literature papers related to the DME synthesis and applications. The DME catalyst development and the reactor technologies are the main research subjects (Figure 4a) while the literature papers about the direct DME synthesis are nearly twice those related to the indirect synthesis route (Figure 4b).

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Figure 4. dimethyl ether synthesis publications: a) subject of the publication; b) synthesis method [21].

The direct DME synthesis is indeed considered more promising in reason of the possibility of reducing significantly the production costs and make DME more competitive as fuel. However, the direct route poses new challenges and more research work is necessary to make the process really feasible. The main issues related to the direct DME synthesis are i) the thermal management of the reactors,ii) the difficult catalyst design and iii) the thermodynamic and kinetic limitations in presence of syngas rich in CO2.

1.3.1 Thermal management

The thermal management of the reactors represents a major issue in direct DME synthesis as consequence of the strong exothermicity of the process (as shown in equation (6) and (7)). The temperature control is fundamental because i) the process is thermodynamically limited; ii) the sintering at high temperature of copper clusters on the zinc-alumina support is the main deactivation mechanism of the CZA methanol synthesis catalyst, that is also employed in DME direct synthesis [18]; iii) temperatures higher than 300°C favours the formation of hydrocarbons (olefin, paraffins, aromatics) with an eventual coke deposition [19]. A proper reactor configuration and design is required in order to carefully control the temperature. As stated in the previous section, slurry phase reactors and multitubular externally cooled fixed bed reactors have been adopted for industrial scale testing of DME direct synthesis. Slurry phase reactors ensure an excellent temperature control thanks to the

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large heat capacity of the solvent. However, these reactors present some disadvantages: i) there are severe mass transfer limitations between the phases that decrease the apparent reaction rates [33]; ii) the auxiliary equipment of the slurry phase reactors is complicated because in addition to the reactor, a recycling system and a separator gas-liquid are required [33]; iii) the loss of catalyst particles is an additional issue for this reactor type [21,34]. Fixed-bed reactors are often preferred for their simplicity and lower costs [21] but differently from slurry phase reactors, do not present an uniform internal temperature profile: in the case of an exothermic process marked hot-spot typically occurs close to the inlet of the externally cooled reactor, while the temperature progressively approaches the coolant one along the axial coordinate. This profile can be advantageous for thermodynamically limited processes since the high temperature at reactor inlet favours the kinetics, while the lower temperature near the reactor outlet enhances the equilibrium conversion. However, in presence of highly exothermic reactions the hot-spot temperature can be excessive, a proper reactor design is needed to effectively remove the reaction heat. In the literature, the majority of multitubular fixed bed reactors for DME direct synthesis have been numerically simulated with single tube 1D models [27–29,34,35] while few work has been made with 2D models [37] which also describe the radial gradients within the tubes. The radial temperature profile is often significant and its proper numerical description is important in the reactor design.

Fluidized bed is another conventional reactor type that has been suggested for the direct synthesis of DME. These reactors have excellent gas-solid mass and heat transfer performances thanks to the intense catalyst particle mixing. However, the collision between the solid particles and the reactor wall can causes loss of catalyst [38]. Fluidized bed reactors, for the direct DME synthesis, have been tested only on a laboratory scale and their applicability on large scale reactors has not been assessed yet [39]. Micro-packed beds are instead an advanced reactor solution that has been proposed also for the direct DME synthesis [40–43]. The small catalyst particles in these reactors provide a high surface/volume ratio with consequent high mass and heat transfer rates. Moreover, the excellent heat dissipation due to the reactor channels smaller than 1 mm, make these reactors suitable for highly exothermic processes.

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The main challenge in the DME synthesis is represented by the catalyst design because two different catalyst functions, the hydrogenation catalyst for the methanol synthesis and the acid catalyst for dehydration, must be coupled in the same reactor.

There are two conventional approaches: combining the two catalyst in a single hybrid/bi-functional pellet or mechanically mixing together the separated pellets of methanol and dehydration catalyst (configuration referred to as “mechanical mixture” in the following). Combining the two catalyst in a hybrid/bi-functional pellet is accounted as the most convenient solution thanks to the close contact between the methanol synthesis and methanol dehydration active sites that eliminates the intraparticle diffusion limitations [29]. The two functions can be intimately coupled in a single hybrid catalyst obtained either by pelletizing a mixture of fine powders of the two formulations or by synthetizing bifunctional material by co-precipitation or by impregnation [22,44]. Unfortunately, it is reported in literature that both the hydrogenation and the dehydration function in hybrid catalyst can suffer from deactivation due to (i) the migration of Cu and Zn to the acidic sites [45,46]; (ii) the migration of Si to the methanol synthesis catalyst[45]; (iii) the pore blockage caused by carbonaceous species deposition [47,48] or (iv) the sintering of Cu nanoclusters due to the interaction of allumino-silicates and Cu sites [49–51]. The mechanical mixture instead minimizes the contact between the methanol and acidic active phases, thus preventing their detrimental interactions, but with a loss in efficiency due to the intraparticle diffusional limitations [29].

Figure 5. Graphical representation of the different catalyst configurations for DME direct synthesis. Colors: auburn - methanol catalyst; grey - dehydration catalyst.

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A recently proposed trade off solution in order to avoid detrimental interaction and, at the same time, provide a close contact between the two catalysts consists in the manufacturing of core-shell (core@shell) pellets. Core@shell are constituted by a pellet core made by a catalyst material covered by a shell of another catalyst material. The graphical representation of all the cited catalyst configurations is reported in Figure 5. In the case of DME direct synthesis, both the configuration with methanol catalyst in the core and DME catalyst in the shell (MeOH@DME) [52–60] and the reverse one with DME synthesis function in the core and methanol synthesis function in the shell (DME@MeOH) [61] have been manufactured and tested. The experimental results show that core@shell catalysts suffer less from issues related to the catalyst deactivation [59] and have better performances in term of DME selectivity respect to the hybrid pellets [52]. Moreover, in the core@shell pellets an intermediate inactive layer can be used, in order to avoid the direct contact between the two active phases and the consequent deactivation [53].

1.3.3 Thermodynamic limitations

The thermodynamic advantage of direct DME synthesis becomes progressively weaker when a gas rich in CO2 is fed to the reactor due to the negative mass action effect of H2O,

which is increasingly produced according to the stoichiometry in equation (7)) [35]. CO2

rich feeds are of interest either when syngas is obtained by biomass gasification or when the focus is on utilization of CO2 . Moreover, it has been found that excess amount of H2O

are detrimental for the activity of the CZA methanol synthesis catalyst [23]. Besides water it is strongly adsorbed on γ-alumina with a consequent loss of activity [20,21]. A possible solution to these issues is given by the in-situ separation of water by membranes or in-situ adsorption. In the next section it is presented a general outlook of DME direct synthesis with steam separation with a specific focus on Sorption Enhanced DME direct Synthesis (SEDMES).

1.4 Sorption enhanced dimethyl ether synthesis

The potential of reactive separation in order to overcome chemical equilibrium limitations is well known. Reactive distillation has been already established in industrial practice in several processes as methyl acetate synthesis, etherifications and ethylene glycol synthesis [62]. However, reactive distillation can be used only in applications where reactants and products are present in separate phases (liquid and gas), while in most of the processes all

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the components remain at the same state. The processes for CO2 utilization belong to this

second category. The large production of gaseous water represents indeed one of the main issues of the processes for the CO2 utilization. In-situ steam separation processes are a

promising solutions for thermodynamically limited processes in which water is the reactions by-product such as rWGS, methanol synthesis, methanation, dimethyl carbonate direct synthesis and DME direct synthesis and for processes, like the Fischer-Tropsch synthesis, in which water removal can mitigate catalyst deactivation [63]. The in-situ removal of water can be performed with two different methods: membrane separation and reactive adsorption. In both cases, the selectivity of the separation material used is the first aspect to be taken into account: an unselective separation is detrimental resulting in additional costs. Once the applicability of the material used for separation is assessed, the choice of the proper method of separation depends on the characteristic of the target process. Reactive permeation with membranes is a continuous process preferred when very low water partial pressures are not required, as for the Fischer-Tropsch synthesis [64,65]. The driving force of permeation is indeed the water partial pressure difference between the permeate and the concentrate, which is strongly affected by transport limitations. For instance in the case of Fischer-Tropsch synthesis is reported that at least 2.5 bar of partial pressure difference is required to make the steam separation feasible [64]. Reactive adsorption is not applicable if water is generated in large amounts, since the main limit of the adsorbent material is represented by its capacity. On the other hand, the advantage of reactive adsorption is represented by the possibility of removing water down to very low levels of partial pressure, thanks to the strong thermodynamic driver of the adsorption process and the low transport resistance. Therefore, the general guideline is that reactive permeation is used for processes with water partial pressure >1 bar, while reactive adsorption is adopted with a required water partial pressure <1 bar [63].

The DME direct synthesis is a process suitable for reactive adsorption (SEDMES), even if the membrane reactive separation has been also reported in the literature [66]. The SEDMES process take place in presence of the same catalysts used in DME direct synthesis and of LTA zeolites sorbents, thanks to their high selectivity to water adsorption [67]. Few works specifically focused on SEDMES are present in literature. The first experimental investigation, on liquid phase SEDMES has been performed by Kim et al. [68]. Without considering the adsorbent regeneration, they observed an effective, but relatively short,

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enhancement given by the water adsorption. The theoretical validity of the SEDMES process has been firstly explored by Iliuta et al. [69] through a numerical reactor simulation. Boon et al. [67] tested the activity and stability of γ-alumina in DME direct synthesis conditions, with relatively large amount of water, showing that the steam deactivation of alumina, possible in this conditions, is reversible and that, potentially, this catalyst is more stable in SEDMES conditions where less water is present in the reacting environment. A paper modelling and design of the adsorption/regeneration cycle has been published by van Kampen et al. [70]. They haveanalyzed the SEDMES cycle with a 1D heterogeneous model, validated with experimental data, addressing the effects on the process performances of different operating conditions and pointing out the critical role of the regeneration method (TSA, PSA, TPSA). Moreover, the importance of heat management has been pointed out, showing that a strong loss in DME yield is observed when operating under adiabatic conditions instead of isothermal ones.

CO + 2H2 CH3OH CO2 + H2 CO + H2O CO2 + 3H2 CH3OH + H2O 2CH3OH CH3OCH3 + H2O H2O H2O ads Recycle CO, CO2, H2, CH3OH, CH3OCH3, H2O CH3OCH3 CH3OH, H2O H2O ads H2O Outlet purge SEDMES Inlet purge Feed CO, CO2, H2

Figure 6. SEDMES process.

A schematic representation of a possible SEDMES process is reported in Figure 6. The process is conceptually similar to the DME direct synthesis (Figure 3) with some substantial differences: i) there are multiple reactors, one working in adsorption/reaction

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step and the others in depressurization-regeneration-repressurization; ii) the recycle of reactants is optional because high conversion with high carbon selectivity to DME (over 80% with any CO/CO2 and stoichiometric H2/COx at 275°C and 40 bar) is potentially

reached in SEDMES [63].

The overall process stoichiometries, respectively from CO and CO2, are reported in

equation (8) and (9).

2 CO + 4 H2↔ CH3OCH3+ H2O ↓𝑎𝑑𝑠 ∆𝐻𝑟0= −250.0 kJ/molDME (8)

2 CO2+ 6 H2↔ CH3OCH3+ 3 H2O ↓𝑎𝑑𝑠 ∆𝐻𝑟0 = −259.7 kJ/molDME (9)

In the SEDMES all the reactions (equation (1-3) and (5)) of conventional DME direct synthesis occur together with water adsorption, but the WGS reaction (2), which proceeds reversely as rWGS, as in the case of methanol synthesis, since water is removed shifting the thermodynamic equilibrium. The effective exothermicity of the process is higher than the conventional process one since also the enthalpy of adsorption needs to be accounted. Considering the isosteric heat of adsorption -45.95 kJ/mol (per mole of adsorbed water) used by Gabrus et al. [71] for zeolite 3A, the overall heat per mole of DME produced is reported in equation (8) and (9). In view of the high exothermicity of the process a rational reactor design is required for a careful control of the temperature.

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1.5 Aim of the work

The goal of this PhD work is the development of mathematical models of chemical reactors for the direct synthesis of DME, to be used for the simulation and design of full-scale equipment. The models must provide an accurate description of all the significative chemical and physical phenomena occurring inside of the reactors, with particular focus on the thermal behavior, which plays a key role in a correct reactor design. The work includes the modelling of both a conventional multitubular externally cooled fixed bed reactor and an advanced SEDMES reactor. The sorption enhancement is indeed a promising concept for solving some of the issues related to the direct DME synthesis.

The work is also part of the Horizon 2020 project “FLEDGED” [72] (FLexible Dimethyl ether production from biomass Gasification with sorption-EnhanceD processes) whose purpose is the development of a process for the large scale production of DME to be used as alternative fuel, starting from different type of biomasses and/or organic waste feedstocks (project logo in Figure 7).

Figure 7. FLEDGED project logo.

The syngas obtained by biomass gasification contains a larger and variable content of CO2

compared with traditional feedstock (e.g. syngas from natural gas) [73]. As a consequence, the DME synthesis train is supposed to operate under flexible conditions, depending on the different possible feeds. The SEDMES concept is particularly advantageous in this case, since it allows to obtain high DME yields independently of the CO/CO2 content in the feed

[63,70].

The thesis work starts with a preliminary thermodynamic analysis of SEDMES (Section 4.1), in which are explored the performances of the process in different operating conditions

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(considering temperature, pressure and inlet composition). The analysis includes also the investigation of the ‘SEDMES reactor train’, a process configuration in which a SEDMES unit is coupled with a conventional converter and a water separation unit, positioned upstream, in order to reduce its heat and water duties.

The next part (Section 4.2) consists of the analysis and design of the conventional direct DME synthesis reactor that operates in the range of conditions of the FLEDGED project. A dedicated paper (Section 5) focus on the effects of the active phase distribution at the catalyst pellet scale on the reactor performances. Different active phase distributions (hybrid pellets, mechanical mixture of methanol and dehydration catalyst pellets, core@shell pellets) are analyzed, interpreting the results through the analysis of the reaction-diffusion phenomena inside the catalyst pellets, highlighting the importance of the intraparticle diffusion limitations on the consecutive reaction scheme involved in the direct DME synthesis process.

The last section (Section 6) is related to the development of a dynamic model used for the analysis and design of the SEDMES reactor unit. The model, validated by comparison with experimental bench scale data, describes the dynamic profiles of composition and temperature inside the reactor during the reaction/adsorption step, allowing for a rational design of the unit.

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2 Methods

2.1 Thermodynamic analysis

The thermodynamic analysis is performed by calculating the equilibrium composition at the reactor outlet at imposed temperature, pressure and inlet composition using the ‘extent of reaction’ ξj method (j-reactions considered). The species actively involved in the DME

direct synthesis are CO, CO2, H2, H2O, CH3OH, CH3OCH3, therefore three independent

reaction stoichiometries are considered: the methanol synthesis from CO (1), rWGS (2) and methanol dehydration to DME (3).

CO + 2 H2 𝜉1 → CH3OH (1) CO2+ H2 𝜉₂ → CO + H2O (2) 2 CH3OH 𝜉3 → CH3OCH3+ H2O (3)

The equilibrium composition is obtained by solving the algebraic system of j-equations (5) in the variables ξj. 𝑛𝑖 = 𝑛𝑖𝑜+ ∑ 𝜈𝑖𝑗 𝑗 𝜉𝑗 (4) 𝐾𝑒𝑞,𝑗 = 𝑒𝑥𝑝 (− ∆𝐺_𝑗𝑜 𝑅𝑇) = ∏ 𝑓𝑖 𝜈𝑖𝑗 (𝑇, 𝑃, 𝜉𝑗) 𝑁𝐶 𝑖 (5)

The equilibrium constants (6), (7) and (8), taken from literature references [39,74], are used for the calculation.

log Keq,1= 5139 T − 12.621 (6) log Keq,2= −2073 T + 2.029 (7) ln Keq,3= 4019 T + 3.707 ln T − 2.783 x 10 −3_{T + 3.8 x 10}−7_T2₋6.561x 10 4 T3 − 26.64 (8)

In the case of SEDMES, also adsorption, removing H2O from the gas phase, must be taken

in account, therefore an additional reaction is added (9).

H2O 𝜉4

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In order to calculate the extent of the additional rection the value of the outlet water partial pressure is imposed. The equations are implemented in gPROMS® ModelBuilder 5.0.2 commercial software for the numerical solution. Fugacity coefficients to be used in equilibrium equations are calculated using the gPROMS® Multiflash 4.3 utility tool, which implements the Redlick-Kwong-Soave (RKS) equation of state.

2.2 Conventional reactor model

The conventional multitubular externally cooled fixed-bed reactor for the direct synthesis of DME is mathematically described by a steady state 2D heterogeneous model of a single catalytic tube. The model consists of i-species mass (i= CO, CO2, H2, H2O, CH3OH, DME,

CH4, N2, Ar), energy 2D balances for the gas-phase, written in cylindrical coordinates,

coupled with i-species continuity equations for the catalyst phases. The model evaluates pressure drops along the axial coordinate by a 1D momentum balance In order to estimate the composition gradients inside the catalyst pellets and calculate the catalyst effectiveness factors, the 2D model is also coupled with 1D i-species mass balances of the catalyst isothermal pellets. Different pellet models are formulated considering the catalyst active phase distribution (hybrid pellets, mechanical mixture of different catalyst pellets, core@shell pellets). In the case of a mechanical mixture of different pellets two separate balances are required, one for the methanol synthesis catalyst and one for the dehydration catalyst. In the cases of hybrid and core@shell pellet configurations instead there is only one type of catalyst particles. In the hybrid catalyst pellet, since the two solid phases are intimately dispersed perfect mixing approximation is adopted. In the case of a core@shell the catalyst pellet consists of two layers (the core and the shell) corresponding to the one of the catalysts that activate different reactions and are separated by an internal interface. The kinetic equations, diffusivity and transport correlation are taken from literature references, while the physico-chemical mixture properties are calculated with the gPROMS® Multiflash 4.3 utility tool. All the equations are implemented in gPROMS® ModelBuilder 5.0.2 for the numerical solution of the boundary values problem. A detailed description of the model is reported in Section 5.2.

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2.3 SEDMES reactor model

The SEDMES multitubular externally cooled fixed-bed reactor is mathematically described by a 2D heterogeneous single tube model. The model is dynamic, since SEDMES intrinsically operates in unsteady state conditions. The model is able to describe the time evolution of concentration and temperature radial and axial profiles of the adsorption/reaction step of a SEDMES cycle. The model includes 2D total mass balance for gas phase, 2D i-species mass (i= CO, CO2, H2, H2O, CH3OH, DME, N2) and energy

balances for the gas-phase, catalyst and adsorbent solid phases. Two separate k-catalyst phases, methanol catalyst (MeOH) and dimethyl ether synthesis catalyst (DME), are considered. The internal diffusion limitations in the catalyst particles are accounted by coupling the dynamic 2D balances in the gas phase with pseudo-stationary 1D mass balances of i-species in isothermal catalyst pellets. The pressure drops in the reactor, evaluated with a 1D momentum balance, are found to be negligible in the investigated range of operating condition. Therefore, isobaric conditions are taken in the simulations. The kinetic equations, diffusivity and transport correlation are taken from literature references, while the physico-chemical mixture properties are calculated with the gPROMS®

Multiflash 4.3 utility tool. All the equations are implemented in gPROMS®_ModelBuilder

5.0.2 for the numerical solution of the boundary values problem. A detailed description of the model and is reported in Section 6.2.

The SEDMES reactor model is validated by comparison with the experimental dynamic behavior during the adsorption/reaction step of a bench scale SEDMES tubular reactor operated at the TNO test facilities in Petten.

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2.4 gPROMS ModelBuilder

The simulation and calculations of this thesis work are performed with the commercial software gPROMS® ModelBuilder 5.0.2 (general PROcess Modelling System), developed by Process System Enterprise. The software is specifically developed for the custom modelling and simulation of process engineering problems. gPROMS® is particularly oriented to solve dynamic problems, having a specific setting for the time variable. Nevertheless, it has an Integro-Partial Differential Algebraic Equation (IPDAE) formulation that allows to solve both steady-state and dynamic problems simply by setting the appropriate specifications. Moreover, it is possible to handle any number of spatial dimensions (e.g. axial and radial distribution in a tubular fixed bed reactor) or property distributions.

The software offers the possibility of using different standard solver procedures and discretization methods: the standard solver ‘DASOLV’ for differential-algebraic equations systems based on variable time step with variable order Backward Finite Difference Method (BFDM) is used for time integration. ‘BDNLSOL’ (Block Decomposition NonLinear SOLver) block decomposition and triangularization method is used to solve non-linear algebraic equation systems, the single blocks are solved with ‘SPARSE’ based on Newtonian-type method. In the reactor simulations, a first order BFDM is used for the discretization of the axial reactor coordinate, instead, third order Orthogonal Collocations on Finite Elements Method (OCFEM) are used for the radial and the pellet coordinates. The number of discretization points used in BFDM and the collocation elements in OCFEM is checked by convergence analysis. In the case of conventional DME direct synthesis reactor (steady state conditions) the discretization points along the axial coordinate in a non-uniform grid: the points are spaced using a logarithm transformation implemented in gPROMS® with a transformation parameter α=15. This allows to have a denser discretization grind in the first part of the reactor, giving a better description of the sharp profiles generated close to the inlet. Conversely, in the case of SEDMES reactor a uniform axial grind is choose since the shape of the variable profiles evolve in time, making the use of a non-uniform grind counterproductive.

The numerical convergence is reached acting both on first guess solutions (in the case of steady state simulations) and algorithm parameters. The first guesses are obtained generating a trial solution of the whole equation systems with ‘easy solution’ conditions

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(e.g. with chemical kinetics off). The algorithm parameters changed in the simulations are in ‘DASOLV’ the AbsoluteTolerance and RelativeTolerance (both set at 1e-5 in the final solution) and in ‘SPARSE’ the ConvergenceTolerance (set at 1e-5 in the final solution) and the SLRfactor ‘Step Length Restriction factor’. This parameter improves the convergence from poor initial guesses, limiting the step in the iterations so that the magnitude of the change in any variable does not exceed. A large value of SLRfactor is employed in the cases of first guesses far from the expected solution, while smaller SLRfactor is used for reasonably close solutions.

gPROMS®_{can also import external objects: in this work the utility tool Multiflash 4.3,}

included with the license, is used to generate external files for the calculation of the gas mixture properties.

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3 Resume of paper results

3.1 A model analysis of the effects of active phase distribution

at the pellet scale in catalytic reactors for the direct dimethyl

ether synthesis

Direct synthesis of DME requires to combine two catalyst functions, methanol synthesis and methanol dehydration, in the same reactor. At the pellet scale such two catalyst functions can be spatially arranged in different ways; intimately coupled in a single hybrid pellet; in a mechanical mixture of different catalyst pellets; in core@shell pellets where the methanol catalyst is placed either in the core with the dehydration catalyst in the shell (MeOH@DME) or reversely (DME@MeOH) (see Figure 5).

Figure 8. Graphical representation of the different catalyst configurations for DME direct synthesis. Colors: auburn - methanol catalyst; grey - dehydration catalyst.

A 2D+1D heterogeneous model of a single tube of a multitubular conventional reactor model is used for an accurate analysis of the above catalyst configurations (Hybrid, mechanical mixture, MeOH@DME and DME@MeOH), considering a benchmark direct DME synthesis catalyst formulation, CZA+γ-Al2O3. Tubes with 25.65 mm internal

diameter and 6 m length are taken as a typical geometry of full scale reactors and on the basis of the tube diameter analysis reported in the Section 4.2. Spherical catalyst particles with a 4.86 mm diameter is in line with data on commercial methanol catalyst [16]. A typical feed composition in direct synthesis DME reactors with recycle loop is considered

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with a H2/CO= 2, a CO/CO2 ratio of 2/1 and an inert content (CH4) equal to about 8% [24].

A weight ratio of MeOH/DME catalyst of 2 is assumed according to [75], which corresponds to a 1.5 volume ratio considering the different densities [16,76]. For the core-shell configurations these specifications in the case of MeOH@DME correspond to an internal radius of 2.04 mm of the methanol catalyst core is with an external dehydration catalyst shell 0.39 mm thick. Due to the lower volume fraction of the dehydration catalyst, the core in the DME@MeOH case has a radius 1.80 mm and the methanol catalyst shell is 0.63 mm thick. Kinetics of methanol synthesis and methanol dehydration reactions are taken from the literature [75].

0 1 2 3 4 5 6 0 2 4 6 8 10 12 14

MeOH hyb. MeOH MeOH@DME

MeOH mech. MeOH DME@MeOH

DME hyb. DME MeOH@DME

DME mech. DME DME@MeOH

Spec ifi c mola r f lo w ra te [mol /m 2 /s ] Axial coordinate [m] b 0 1 2 3 4 5 6 0 10 20 30 40 CO hyb. CO MeOH@DME CO mech. CO DME@MeOH CO₂ hyb. CO₂ MeOH@DME CO₂ mech. CO₂ DME@MeOH Spec ifi c mola r f lo w ra te [mol /m 2 /s ] Axial coordinate [m] a

Figure 9. Axial profiles of the average specific molar flow rate of a) reactants and b) products with different catalyst configurations..

In Figure 9 the specific molar flow rates per unit area of reactants (CO and CO2) and

products (MeOH and DME) are reported. The ranking in the consumption rate of CO is as follows: Hybrid>DME@MeOH>MeOH@DME>MM; the conversion of CO is fastest in the hybrid pellet configuration thanks to the synergy between the involved reactions, i.e. methanol synthesis and dehydration. Figure 9a also shows that CO2 is not a reactant, but

mainly a product of the process, in fact its flow rate slightly increases along the axial coordinate [29]. Wide differences are observed in the distribution of the products; i.e. methanol and DME. Figure 9b shows that DME@MeOH pellets exhibit the highest methanol net production rate, with DME production close to the one obtained with the mechanical mixture. The production of DME is highest in the hybrid configuration thanks to the synergy between methanol synthesis and dehydration. However, the MeOH@DME pellets show the DME formation rate closest to the hybrid pellet one, widely outperforming