Perspective of dimethyl ether as fuel: Part II- analysis of reactor systems and industrial processes

Perspective of dimethyl ether as fuel: Part II- analysis of reactor systems and industrial processes

Journal of CO₂ Utilization 32 (2019) 321–338 Contents lists available at ScienceDirect Journal of CO2 Utilization journal homepage: www.elsevier.com...

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Journal of CO₂ Utilization 32 (2019) 321–338

Contents lists available at ScienceDirect

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

Review Article

Perspective of dimethyl ether as fuel: Part II- analysis of reactor systems and industrial processes

T

Ujjal Mondal, Ganapati D. Yadav⁎ Department of Chemical Engineering, Institute of Chemical Technology, Nathalal Parekh Marg, Mumbai, 400019, India

ARTICLE INFO

ABSTRACT

Keywords: Dimethyl ether (DME) Syngas CO2 Catalysis Fuel LPG Multiphase reactors Methanol

In Part I of this review article, the significance of DME as fuel and the various types of catalysts used for different feedstocks were considered. However, the industrial processes using a variety of reactor configurations affect the overall capex and opex. The production of syngas, irrespective of the source, is the first step in DME synthesis, which is then followed by conversion into DME using a battery of reactors and separators. A critical analysis is presented and future scope outlined.

1. Introduction Catalytic DME synthesis by using both indirect and direct methods was covered in Part I of this review [1]. Not only the nature of the catalyst but types of reactors and separators and their configuration will influence the process economics. In both the methods, either CO2 or syngas is the feedstock for making DME no matter whether coal, petroleum, natural gas or biomass is the primary source. CO2 comes from power plants and is the greenhouse gas. Syngas is the most industrially popular source for DME synthesis [2]. The reversible nature of the reactions in DME synthesis from carbonaceous gases makes the whole reaction thermodynamically limited. The unconverted reactants are separated downstream of the synthesis reactor and partially recycled into the DME synthesis reactor after recompressing. The reactor design is an important parameter for direct DME synthesis as the reaction temperature, feed composition, and pressure play a very important role in reactant conversion and product selectivity. In the past few decades, direct synthesis of DME by bifunctional catalysts has been studied. Direct DME synthesis process in a single reactor combines methanol synthesis and methanol dehydration reactions. Single pot synthesis of DME in a slurry reactor, introduced by JFE Corporation, Air Products and Tsinghua University [2] is among the most promising technologies for large scale DME production. Other than slurry reactor many researchers have evaluated the DME synthesis process from syngas in different types of reactors such as fixed bed reactor [3,4], internal



recycle type reactor [5], fluidized bed reactor [6], and back mixed slurry reactor [7]. Because of the importance of syngas in DME synthesis, a comprehensive review of different production technologies of syngas from various sources is also covered to put DME synthesis in perspective. Thereafter innovative reactor technologies for direct and indirect DME synthesis process are discussed. 2. Syngas production from different sources Syngas is a major requirement for the DME industry. Syngas can be synthesized from a number of natural resources, which include coal, methane, and biomass. Various gasification technologies have been used to produce syngas from biomass and coal. This gasification process involves reforming techniques (steam and/or oxygen), syngas clean-up process to remove various contaminants such as sulfur, ash, mercury, arsenic, and tar for the biomass gasification process. 2.1. Coke from coal and petroleum The purpose of the gasification process is to convert maximum chemical energy stored in the fuel into chemical energy in gases such as CO, hydrogen, and light hydrocarbons. Coal gasification is an established technology for the production of syngas and it has matured over the years regarding the handling of feed, energy efficiency and the introduction of entrained flow gasifier for rapid feed conversion and high

DOI of original article: https://doi.org/10.1016/j.jcou.2019.02.003 Corresponding author. E-mail addresses: [email protected], [email protected] (G.D. Yadav).

https://doi.org/10.1016/j.jcou.2019.02.006

Available online 23 April 2019 2212-9820/ © 2019 Elsevier Ltd. All rights reserved.

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throughput. Bituminous coal is the most suitable source for gasification but due to the abundance of low-grade coal such as sub-bituminous coal and lignite, most of the gasification units are operating on low-grade coal. Petroleum refineries around the world are producing a huge amount of petroleum coke which can be an alternative cheap source for syngas generation. During the coal gasification process, crude coal is reacted with steam and oxygen/air to produce CO2 and CO. There are three well known main types of gasifiers to pyrolyze and gasify the carbon and hydrogen from coal in a continuous automated process.

2.2.2. Partial oxidation (POX) It is the second most popular process for producing syngas from methane. As this is an exothermic reaction, no external energy is required to produce hydrogen and CO gas mixture. A refractory lined pressure vessel is utilized and it is fed with preheated natural gas and oxygen. The gases are mixed in a burner at elevated pressure (40 bar) and immediately the partial oxidation reaction occurs in the combustion zone below the burner [19]. The produced syngas leaves the reactor at 1200–1300 °C. Various kinds of hydrocarbons (from natural gas to heavy hydrocarbons) can be utilized for syngas production by POX method and 100% conversion is achieved in this process. Catalytic POX method also has been developed which requires less volume and consumes less oxygen. Texaco and Shell Technologies are the most popular companies to build POX plant. Other companies which are also working in the same field are Koppers, Lurgi, Foster Wheeler, Starchem, and British Gas [20].

2.1.1. Fixed bed gasifier Fixed bed consists of a fuel bed, through which air or oxygen is blown for gasification purpose. Lurgi is well known for large-scale fixed bed gasifier in the industry [8] and most of the power plants are considering British Gas/Lurgi (BGL) fixed bed gasifier [9]. The BGL gasifier is superior in many areas such as higher gasification rate, better in terms of energy efficiency and the slag produced is a high strength material along with low leachability.

2.2.3. CO2 reforming In this type of reforming process syngas is produced with a lower H2/CO ratio (1:1) [21]. This process is also referred to as dry reforming as steam is not used in this process. The overall reaction mechanism and kinetics is similar to SMR but the rate of formation of carbon is much higher. Nickel and Ni/SiO2 catalysts are generally used for this reaction but noble metal catalyst such as rhodium and iridium also have been employed. After removing sulfur from methane gas, it is mixed with CO2 and purged into the catalyst tubes of reforming vessel, which is the same as SMR. The product gases are CO, H2, CO2 with less amount of CH4, but 100% conversion can be achieved with this process while changing the process parameters [22,23].

2.1.2. Fluidized bed gasifier Coal is grounded to the size of few millimeters to form fluidized bed which is suspended by blowing air or oxygen through the particles [10]. High-quality coal is required to achieve a high conversion rate for this type of gasifier [11]. Uhde-Prenflo have installed the world’s largest gasification plant in Spain based on fluidized bed technology and it operates with 50% mixture of each petroleum coke and coal. This type of gasifier is favored by power plants because of their low conversion efficiency. Char produced from this process is utilized by another unit of the power plant. British Coal and Westinghouse have been involved in the development of an improved version of this kind of gasifier [11]. Co-gasification of various types of biomass, and plastic waste with coal in the fluidized bed reactor has been investigated by many researchers [12–14].

2.2.4. Autothermal reforming (ATR) This process is a combination of partial oxidation and steam methane reforming process in a single vessel [24]. In this process, the energy required to convert hydrocarbons is supplied by POX reaction. From the top of the reactor, a preheated mixture of methane, steam, and oxygen is introduced. POX reaction proceeds at the upper portion of the reactor and then the gas mixture is traveled through the catalyst filled bed of the vessel, where final reforming occurs. The produced syngas is having a high mole ratio of H2/CO (1.6–2.6). Lurgi, Haldor Topᴓse, Foster Wheeler, and ICI are the license holders of these processes.

2.1.3. Entrained gasifier This type of gasifier utilizes pulverized coal which blows into the gasifier in a similar manner to a pulverized coal-fired system and then oxidized by reaction with oxygen or highly oxygen enriched air [15]. Other than coal, this technology can be utilized for petroleum coke and oil feedstock. Krupp Koppers, Dow, Texaco, and Shell are the leading companies which are working with entrained gasifier [11]. This type of gasifier is quite popular as it can handle a wide variety of solid or liquid fuels despite high installation cost and high wastage of energy in terms of heat [16].

2.2.5. Combined reforming Since the conversion efficiency of SMR process is not 100%, it is sometimes combined with a secondary reformer such as ATR or POX unit to reach complete methane conversion and improved H2/CO ratio. Thus this process is also known as “oxygen-enhanced reforming” or “two-step reforming”. This technology is offered by the main process developers of syngas industry like M. W. Kellogg and Haldor Topᴓse, etc.

2.2. Natural gas A number of technologies have been developed to convert natural gas into syngas such as steam methane reforming (SMR), partial oxidation (POX), CO2 reforming, auto-thermal reforming (ATR), and combined reforming [17]. 2.2.1. Steam methane reforming (SMR) Light hydrocarbons are reacted with steam in the presence of nickel catalyst to produce CO, H2, and CO2 via an endothermic reaction [17]. This process is operated in a direct–fired furnace where catalysts are loaded in tubes, through which preheated methane and steam mixture is blown. Higher hydrocarbons are not suitable to use as a feedstock in SMR due to the presence of high sulfur content, which results in catalyst deactivation, tar and coke deposition. The produced syngas leaves the reactor at high temperatures (800–900 °C). Many companies are developing new technologies to make this process more energy efficient and include United Technology Corporation, Haldor Topᴓse, and KTI [18]. Lurgi is the market leader in building SMR plants and a few engineering companies were or continue to build also building SMR such as Howe-Baker, Foster Wheeler, M. W. Kellogg, Kvaener, and the erstwhile ICI.

2.3. Biomass gasification Biomass is an important renewable energy source and thermochemical conversion also known as gasification of biomass produces syngas and a solid product, char. The gasification process involves partial oxidation of carbon present in biomass and the gasification is generally carried out in the presence of carrier gases such as oxygen, air, CO2 or steam (Fig. 1). The produced syngas is a mixture of CO, H2, CO2 and also includes light hydrocarbons. Some unwanted gases are also released with syngas such as sulfide (H2S) and chloridic (HCl) acids along with inert nitrogen gas. Most of the principal reaction of biomass gasification is endothermic and the required energy is provided by oxidation of biomass through either allo-thermal or auto-thermal process. The allo-thermal process requires external energy supply and autothermal process uses the energy produced by partial oxidation. Biomass 322

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Fig. 1. Schematic diagram of DME/methanol plant operated on biomass gasification. Modified from [27].

gasification system includes four main stages, i. e. (i) oxidation (exothermic stage), (ii) drying (endothermic stage), (iii) pyrolysis (endothermal stage), and (iv) reduction (endothermic stage). And finally, a tar decomposition stage is included to convert the big tar molecules into small light hydrocarbons [25]. A number of firms are setting up commercial bio-gas manufacturing plants which use air combined with oxygen, only for combined heat and power (CHP) application. Andritz/ Carbona fluidized-bed gasification technology of USA, The Energy Research Centre of Netherlands (ECN), Chemrec/Volvo/Haldor Topsoe and Uhde of Sweden, Guangzhou Energy Institute of China, and BioMCN of Netherlands are the leading firms which have already setup commercial plants to produce bio-gas from raw biomass, black liquor collected from pulp and paper plant, glycerol, etc. [26]. The foregoing analysis shows that production of syngas is very important for which different technologies have been introduced in the market. Now let us turn our attention to the reactors which influence over economics of the DME manufacture.

Fig. 2. Fixed bed reactor for DME synthesis. Modified from [31].

3. DME synthesis reactors

process, the direct process requires a single reactor and by continuously converting methanol to DME over a bifunctional catalyst, the thermodynamic limitation associated with methanol producing reaction is alleviated [4]. Gas-solid contactors are usually employed in some fixed bed reactors to enhance mass transfer rates. When the reactions are endothermic and exothermic in nature, this type of reactor is not suitable to carry out the reaction due to catalyst sintering. The most important factor which affects the performance of a fixed bed reactor is the formation of hot spots inside the reactor and these overheating zones can deactivate the catalyst irreversibly [28]. For the economic operation, a high recycle rate of syngas is needed to overcome low conversion due to these thermodynamic limitations and catalyst deactivation. CO2 is also used as an alternative to syngas as a starting material for DME synthesis and it requires an additional step, i.e. reverse water shift gas (r-WSG reaction) to hydrogenate CO2. The r-WSG step can influence the whole process significantly as the rates of methanol formation and methanol dehydration reaction are affected by in situ water formation. Many studies have been conducted for modeling and feasibility of DME synthesis by methanol dehydration in fixed bed reactor. Ghavipour and Behbahni [29] have done simulation study on industrial fixed bed reactor system at steady state conditions for DME synthesis by methanol dehydration. Farsi et al. have also developed a model of an industrial adiabatic fixed bed reactor for DME synthesis with preheated feed supply and simulated it at dynamic conditions [30].

DME synthesis from syngas or CO2 is a highly exothermic reaction and in order to avoid the runaway of the synthesis process, the temperature inside the reactor must be controlled. In recent times focus has been shifted to reactor technology to minimize the generated entropy of any chemical process. Initially, fluidized bed reactor was used for the highly exothermic reactions such as DME synthesis [6]. Many researchers have evaluated the DME synthesis process from syngas in different types of reactors such as fixed bed reactor [4,3], internal recycle type reactor [5], fluidized bed reactor [6], and back mixed slurry reactor [7]. Lu et al. [6] have done a comparative study among fluidized bed, fixed bed and slurry reactors for the direct DME synthesis process and found the slurry reactor to be the best choice. In the following section, different reactor systems are discussed which have been either modified or directly used for DME production from syngas or mixture of CO2 and H2. 3.1. Fixed bed reactor Adiabatic fixed bed reactors are quite popular in the refinery and chemical process industries for their ease of operation and maintenance (Fig. 2). Based on the number of DME production steps involved, the fixed bed reactor is designed accordingly. When the indirect process is adopted, then the separate stages of syngas hydrogenation and methanol dehydration are held in separate reactors. Unlike the indirect 323

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Fig. 3. Shell and tube type of reactor configuration for DME synthesis. Figure is taken from [34]. Fig. 4. A schematic diagram of slurry phase reactor. Modified from [47].

3.2. Isothermal adiabatic reactor

wettability of the catalyst because of the mismatch between hydrophilic catalyst surface and hydrophobic organic solvent. Several other problems are also encountered such as aggregation, misdistribution, and rheological deterioration of catalyst particles and therefore the catalyst is poisoned and deactivated. Gao et al. [39] have claimed a proper procedure and environment to prepare a catalyst for slurry reactor. Liu et al. [40] prepared an efficient slurry catalyst by following complete liquid-phase technology from AIP for production of DME from methanol in a slurry phase reactor. Syngas flows upward in the form of bubbles inside the slurry phase reactor, where catalyst particles are dispersed in the inert oil medium. First, the syngas has to get dissolved in the solvent to reach the surface of the catalyst and get converted. Methanol synthesis occurs first followed by methanol dehydration reaction. The water generated by dehydration reaction quickly reacts with CO and is removed from the system. Unlike fixed bed reactor there is no hotspot generation inside the reactor because heat conductivity of the slurry is very high and heat is homogeneously distributed inside the reactor. Heat transfer from the heat exchanger tubes to the slurry is very efficient and thus heat removal from the reactor for the highly exothermic reaction is very easy. Thus this type of reactor has several mass transfer limitations and the productivity of the reactor is affected greatly [41–44]. Papari et al. [45] have simulated a large scale slurry bubble column reactor for the production of DME from syngas. In their study, the catalyst deactivation is also taken into account for the transient conversion of syngas and DME yield. They evaluated the optimum temperature profile and under the uniform catalyst distribution inside the reactor, the positive effect of catalyst concentration and reaction pressure has been predicted. Aoki et al. [46] have established a pilot plant of 100 TPD for DME synthesis from syngas in a slurry phase reactor.

Isothermal reactors are a type of fixed bed reactors which are employed for DME synthesis. The isothermal adiabatic reactor is a combination of heat exchanger and conventional reactor, in which reactions occur at the tube side and the heat of the reaction is removed with using a cooling medium from the shell side of the reactor (Fig. 3). In comparison with the conventional adiabatic reactor, the isothermal reactor is more controllable and higher yield generator. Reversible and exothermic reaction operated in plug flow or batch reactor leads to different optimal temperature distribution which is the main deciding factor for product selectivity among DME, methanol, and ammonia. Shell and tube fixed bed reactors have been employed for the DME synthesis from methanol [32], CO2 rich feedstock [33] and syngas [34]. Hu et al. [35] have built a mathematical model of the shell and tube reactor system based on the global kinetics of the direct DME production from syngas in the presence of a bifunctional catalyst and optimized the reactor structure and operating conditions for achieving the highest yield of DME at operational economic conditions. Farsi et al. [32] also have optimized and modeled indirect DME production from methanol in shell and tube fixed bed reactor based on the energy and mass conversion equations. They have reported maximum DME yield by adjusting the temperature profile along the length of the isothermal reactor by using a genetic algorithm. Methanol conversion of the modified reactor reached 85.75% from 81.9% of the conventional reactor and therefore DME production increased by 4.3%. Bai et al. [36] have simulated DME synthesis through methanol dehydration in a vapor phase adiabatic fixed bed reactor with a capacity of 1000 KTPY. They have compared their simulated results with industrial DME production plant with 2 MTPA capacity. In their follow-up paper, a fivebed industrial adiabatic reactor battery fitted with heat exchangers (for inter-stage cooling) is reported [37]. They also simulated a three-bed adiabatic reactor with interstage cooling, which is provided by injecting a fraction of fresh cold reaction feed.

3.4. Fluidized- bed reactor According to many studies, fluidized bed reactor is the ideal reactor for DME synthesis [48]. Use of fine catalyst particles makes this kind of reactor free from mass transfer resistance in comparison with fixed bed and slurry reactors. Other advantages of fluidized bed reactors (FBR) are: (a) high mixing enables superior temperature control and very fast to attain isothermal condition, (b) less pressure drop, (c) catalyst regeneration and recycling process is very easy, and (e) more quantity of catalyst can be used leading to higher rates. When the three type of reactors are compared based on different parameters such as CO conversion, DME selectivity and DME productivity, respectively, the fluidized bed reactor shows the highest performance (48.5%,97%,0.45 g.gcat−1.h−1) followed by slurry reactor (7%, 70%, 0.2 gg-cat−1.h−1) and FBR (10.7,91.9,0.5 gg-cat−1.h−1) [17,6,49,50]. Despite these advantages and potentials, the application of FBR for DME synthesis has

3.3. Slurry phase reactor Slurry bubble column reactor is a type of multiphase pneumatic contactor. Other than adiabatic fixed bed reactor, it is one of the most promising reactor systems for DME synthesis. There are several reasons for the better performance of slurry bubble phase reactor than fixed bed reactor such as simple construction, uniform temperature distribution inside the reactor for highly exothermic reaction, easy catalyst addition, and separation process, good control of the reactor temperature which helps to avoid catalyst sintering and low energy requirement for adequate mass transfer between phases [38]. In the traditional slurry reactors catalyst particles are introduced by dispersing them in inert solvent inside the reactor (Fig. 4). This creates problems such as non324

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not been properly investigated. Abashar et al. [51] have explored the potential of multiphase fluidized bed catalytic reactor for the use of DME synthesis through simulation studies. Their simulation studies indicated that dual–phase fluidized bed reactor showed a substantial increase in DME yield with almost 100% selectivity. They concluded that catalyst bed composition of the second bed controls the performance of the reactor and multistage fluidized bed reactor can be a better option to employ for DME synthesis in industry (Fig. 5). Randhava et al. [52] have patented a technology to produce DME from syngas in the presence of catalyst in a fluid pluralized bed reactor operated in the gas phase. As the DME synthesis from syngas is a highly exothermic reaction, the removal of reaction heat generated inside the reactor is necessary for higher yield. Uniform temperature profile was achieved inside the reactor by removing reaction heat from the reactor which also resulted in the reduction of the flow rate of recycling syngas within the DME synthesis loop. Elimination of the diffusion resistance and shift of the equilibrium to more favorable reaction conditions, i.e. product diffusion from concentrated region to bubble phase, are among the other reasons behind the higher efficiency of fluid pluralized bed reactor. Catalysts which are used in fluidized bed reactor are very prone to attrition because of the collision between the wall of reactor and catalyst particles, which is the prime reason for deactivation and loss of catalyst [53]. Teng [54] has developed a suitable attrition resistant catalyst for DME synthesis in fluidized bed reactor and introduced silica sol as a binder for active components (methanol synthesis and dehydration catalyst) followed by spray drying to achieve low attrition with a minimum decrease in catalytic activity.

reaction rate and then at a later stage of the reaction, the temperature should be decreased to get higher equilibrium conversion [55]. The implication of multifunctional reactor is an efficient and compact way to reduce the operating and production cost of a conventional reactor, operating with a reversible endothermic reaction and it is widely used as a process intensifier in many industries. In this reactor, the heat is generated by an exothermic reaction and used as a heat source for other endothermic reaction [56,57]. In the last few years, there are several reports published in the literature regarding multifunctional reactors [58]. Hunter and Mcguire [59] are one of the first researchers to investigate this type of reactor system. They have utilized heat exchangers to harness the heat produced in catalytic combustion or any other exothermic reactions to sustain other endothermic reactions. Similarly many other reports are available on multifunctional reactors [60–63]. Glockler et al. [64] have proposed a simple theory to design and analyze the process of combining endothermic and exothermic reactions in an adiabatic fixed bed reactor operated in a reverse flow mode. Farsi et al. [65] have done simulation study of DME formation from methanol dehydration and cyclohexane dehydrogenation in a dual reactor which is a combination of two fixed beds configured as a heat exchanger. Vakili et al. [66] have built a steady-state one dimensional model of direct DME synthesis from syngas coupled with cyclohexane dehydrogenation reaction in a coupled reactor. The results showed several improvements over conventional reactor systems such as reduction in reactor size, wherein hydrogen and benzene had been produced as additional products and a low outlet temperature of the product stream was achieved. This reactor configuration appears to have a very high potential in the industry as 15,000 and 19,000 TPA of benzene and hydrogen can be produced along with 600 TPA of DME production. Yusefi et al. [31] have studied the direct DME synthesis process in an innovative system composed of a couple of shell and tube reactors, where the standard commercially utilized DME synthesis catalyst has been placed in the tube side of the first reactor and the shell side of the second reactor (Fig. 6). In order to the break the thermodynamic barrier of direct DME synthesis, the first reactor is cooled by a cold water present in the shell side and the second reactor is gas cooled. Preheated syngas will enter in the tube side of the first reactor and the cold water present in the shell side of the reactor, will absorb the heat produced in the tube side and therefore, heat energy is recovered in the form of water vapor from the first reactor. The reacting gaseous stream will enter the shell side of the second reactor and fresh cold feed will enter the tube side of the second reactor. This flow of two gases at different temperatures causes a continuous reduction of temperature along the length of the second reactor. Co-current and counter-current are the two possible flow patterns of the cold feed gas and the hot reacting gas that have been studied. The product is collected from the downstream of the second reactor [31]. Vakili et al. [66] have concluded that the counter-current flow pattern in the second reactor offers better DME synthesis rate. According to the simulation results, the reactors which are operated according to the proposed optimum conditions, can produce 60 TPA DME.

3.5. Dual type reactor

3.6. Catalytic distillation

DME synthesis from the hydrogenation of CO and/or CO2 is a reversible exothermic reaction, and strong effect of temperature profile variation has been observed on the final conversion of this type of reaction. Initially, the reaction is kinetically controlled, and as the reaction progresses, the reaction rate also increases with the surge in temperature. On the contrary, increasing temperature can harm the equilibrium conversion of the reversible exothermic reaction. Thus, in order to achieve higher conversion involving this reaction, the reaction temperature should be decreased as the reaction proceeds towards the equilibrium. So, in the case of single step DME reaction, it is necessary to increase the reaction temperature at the beginning to achieve higher

As DME production in a conventional reactor system neither achieves complete conversion nor high purity, a number of improvements of reactor system has been proposed by researchers. Current industrial approach for the DME synthesis involves an FBR system followed by two distillation columns to achieve ultra-pure DME (>99%) [67]. A number of research studies have been conducted to investigate the possibility of using reactive distillation (RD) as an alternative to methanol dehydration reaction for DME production [68,69] (Fig. 7). Reactive distillation is a process which intensifies a reaction by combining reaction and product separation process (distillation) in one unit to felicitate continuous product synthesis followed

Fig. 5. A schematic diagram of fluidized-bed reactor. Reproduced from [51].

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Fig. 6. A schematic diagram of dual reactor system. Modified from [31].

by separation. It is also known as catalytic distillation (CD) when the solid catalyst is used [70]. Dehydration-of-methanol like equilibrium limited reactions are already implemented with reactive distillation reactor system in the industry and it was found to be more economically beneficial, process intensifier, and energy efficient with low investment cost [70–72]. One CD tower can replace the dehydration reactor and distillation columns in conventional DME process. Catalytic distillation reaction conditions are mild in comparison to other methods such as moderate reaction temperature (40–180 °C) and pressure (800–1200 kPa) [66,73,74]. The mechanism of methanol dehydration in CD tower is very simple. The reactant methanol has an intermediate boiling point between water and DME. Thus methanol is concentrated at the middle of the reactor, where the catalyst is loaded. Water and DME can be withdrawn from the bottom and top of the CD tower, respectively. The design pressure is low in the tower for methanol dehydration reaction because of the high boiling point difference betwewater and DME. Thus tower operation becomes cheap as low-cost heat source and sink can be utilized in the CD tower [75]. There are also other alternative technologies available for the DME synthesis such as simultaneous synthesis and separation of DME in fully thermally coupled distillation column, dividing wall column (DWC) replacing the direct distillation sequence, heat-integrated distillation or cyclic distillation, self-heat recuperation, and reactive dividing-wall column [76–80]. Petlyuk et al. [76] have designed a DWC which consists of two thermally coupled distillation columns and it was implemented dividing a single tower into two sections by inserting a single wall at proper position. This type of reactor holds many advantages over conventional systems since it can separate more compounds in a single unit and thus saving the cost of the additional unit along with reboiler and condenser. Kiss et al. [77] have proposed an innovative DME synthesis method based on reactive (R)- DWC tower, which combines reactive distillation with DWC in a single vessel. They were able to produce DME at lower cost and with a lower carbon footprint within a single unit. They have compared their improved R-DWC system with conventional and RD systems and reported that their optimized system could save energy up to 12–58%, and CO2 emission could be reduced to 60%; and all of these was possible with 30% lower operating cost. Lei et al. [68] have investigated DME synthesis from methanol dehydration by an improved synthesis process consisting of FBR and catalytic distillation column. A kinetic model based on the Eley–Rideal mechanism and the empirical power-rate law have been proposed for DME

synthesis from methanol dehydration reaction over macroporous sulphonic acid ion exchange resin as catalyst. Su et al. [75] have explored the designing parameters of a CD tower for DME synthesis and designed a dual catalytic system to minimize the tower temperature, and catalyst deactivation rate. They have modeled the CD tower and validated the model based on the VLE data and pilot test data. They have designed the CD tower in such a way that the low-temperature catalyst and hightemperature catalyst are kept in upper and lower portion of the tower, respectively and the produced DME is condensed through normally available cooling water. Bîldea et al. [81] have done a rigorous design simulation study of DME synthesis in a reactive distillation reactor and proposed an optimum plant design based on RD which can significantly improve existing reaction-distillation-recycle approach of DME production. The new 100 KTPY capacity plant is promised to operate with lowest opex, and capex, though the energy consumption of the plant is 2.43 MJ/kg DME) only. Tong et al. [82] have explored the possibility of combining methyl acetate hydrolysis reaction with methanol hydration reaction in an RD tower and the effect of different operating parameters such as methyl acetate conversion, methanol conversion and DME mole fraction have been optimized. It was realized that the rate of methanol hydration in the RD column was the rate-limiting step. Kaewwisetkul et al. [83] have designed an RD column for DME synthesis from dehydration of methanol in crude glycerol, which is a by-product in a biodiesel plant. They have used RADFRAC module in Aspen Plus to study the reaction parameters on RD column such as reboiler duty, reflux ratio, the number of stages required in the RD column for high yield and RD column conjugation is also optimized using a total annual operation cost. 3.7. Membrane reactor In recent years, process intensification of DME synthesis has been achieved by adding membrane sieve for water removal with the catalysis process within the reactor. One mole of water is produced along with one mole of DME from two moles of methanol. Therefore, if the produced water can be selectively withdrawn from the system, then according to the Le Chatelier’s principle the reaction equilibrium will be shifted towards selective DME synthesis. The inclusion of selective membranes in the catalytic reactor system has shown promising results [84–87] and the reactor configuration is called Membrane Reactor (MR). Hydrophilic selective membranes are of many types depending 326

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properties (such as selective H2O permeation and high temperature tolerance) to be used in MR for DME synthesis. To overcome this issue, a solid-acid catalyst F-4SF resin has been utilized in catalytic membrane reactor to convert methanol into DME, which attains 36% methanol conversion with 100% DME selectivity [33,89]. Fedosov et al. [90] have synthesized a thermally resistant metal-ceramic supported layered zeolite membrane which achieves similar conversion and selectivity. De Falco and coworkers [91] have studied all aspects of DME synthesis by CO2 valorization process, such as thermodynamic analysis, simulation of the performance of the direct process of the MR, which is integrated with microporous zeolite membrane. Lee et al. [92] have claimed that DME yield increased to 83% from 63% when the silica-alumina membrane was used in the reactor. Another report by Farsi et al. [30] mentions that methanol conversion is increased from 80 to 86% when the silica-alumina membrane is used for in situ water removal. Iliuta et al. [93] have done a numerical analysis of the effect of in situ water removal, from the fixed bed membrane reactor under DME synthesis conditions, on the rate of DME synthesis from the mixture of CO, CO2, and H2. It was found that a fixed bed membrane reactor with water removal process through the hydrophilic membrane leads to a higher yield than conventionally fixed bed reactor. Zhou et al. [94] have employed a catalytic dual membrane system for DME synthesis, which is a sandwich of a reactive membrane (zeolite H-FAU) for methanol conversion and a water-penetrable membrane (zeolite Na-LTA) for selective water removal (Fig. 8). Methanol is converted to DME on the catalytically active zeolite H-FAU membrane. Thereafter water produced in the reaction permeates through the next layer of hydrophilic zeolite Na-LTA membrane. The zeolite FAU–LTA double-layer membrane was prepared on α-Al2O3 support. The pore size of the Na-LTA membrane is smaller than the kinetic diameter of DME (0.43 nm), and thus selectively allows passing of water molecules. This double layer membrane was able to achieve 91% methanol conversion with 100% DME selectivity. Diban et al. [95] have demonstrated a DME synthesis process combined with CO2 capture and recovery process in a packed bed catalytic membrane reactor (PBMR) and a “non-ideal” zeolite membrane was incorporated in the reactor for in situ water removal from the reactor. A mathematical model has also been developed to investigate the influence of flow rate, sweep gas recirculation rate to enhance the CO2 conversion and DME yield. The PBMR configuration has shown higher performance than conventional PBR; for instance, as CO2 conversion increased to 85%, DME yield increased to 55% from 16.3% (PBR), and the recirculation of sweep gas stream is responsible for the

Fig. 7. Schematic representation of various improved DME synthesis processes: (a) fixed-bed reactor (FBR) is followed by a catalytic distillation column; (b) a fixed-bed reactor (FBR) is followed by two ordinary distillation columns; and (c) a catalytic distillation (CD) column followed by an ordinary distillation column. Adapted from [68].

upon their constituent material and pore structure, such as (i) microporous zeolite membrane, (ii) polymeric membranes, and (iii) amorphous microporous membranes. It was found that the microporous zeolite membranes can withstand temperature of more than 200 °C and give satisfactory water molecule permeation (the range of H2O permeance was within 10−7−10−6 mols−1 m−2 Pa−1) with an H2O/H2 selectivity more than 10 [88]. Despite the advantages of zeolite membranes, some studies have found that when the CO2 rich feedstock is used for DME synthesis, then lower DME yield has been observed due to the permeation of reactants through the membrane such as methanol [87]. But till now no zeolite membrane is developed with suitable

Fig. 8. A schematic representation of membrane reactor (MR). Reproduced from [94]. 327

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U. Mondal and G.D. Yadav

Fig. 9. A schematic diagram of thermally double-coupled two-membrane reactor (TDCTMR). Reproduced from [97].

Fig. 10. A schematic diagram of thermally double coupled double membrane heat exchanger reactor (TDMHR). Reproduced from [99].

328

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U. Mondal and G.D. Yadav

Fig. 11. A schematic diagram of axial-flow spherical packed-bed (AF-PSR) reactor for DME production. Reproduced from [100].

Fig. 12. A computational element in the axial flow sperical packed bed membrane reactor (AF-SPMR). Reproduced from [100].

reduction methanol loss which permeates through zeolite membrane and increases in DME yield. Falco et al. [96] demonstrated a basic design of a DME plant which works upon a membrane reactor to convert CO2 at higher conversion. A unique DME synthesis process configuration, named Double Recycling Loop DME (DRL-DME) was adopted, which enables pure CO2 gas to be used as a sweeping gas in the permeation site and a double recycling loop to reintroduce unconverted syngas as well as the permeation side gases. A mathematical model of the DME plant was also developed to study the effects of different process parameters on the plant productivity and it was observed that with this DRL-DME process configuration, 60% conversion of CO2 along with 60% DME yield was achieved. Farniaei et al. [97] have studied a thermally double-coupled twomembrane reactor (TDCTMR) with H-SOD (eliminating water from methanol side) and Pd/Ag (transferring unreacted H2 to DME side) membranes for the simultaneous synthesis of hydrogen, DME and methanol via different cycle systems (Fig. 9). TDCTMR is a multi-tubular reactor where dehydrogenation of cyclohexane takes place in the central tube (endothermic reaction) and exothermic reactions such as methanol, DME generation takes place in the inner and outer tubes, respectively. Three different types of recycled steams have been tried and when a mixture of DME and methanol outlet steam was recycled into DME side, the highest yield of hydrogen, methanol, and DME was observed. A steady-state one-dimensional heterogeneous catalytic model was also developed to investigate this innovative reaction process. Samini et al. [98] also investigated the thermally coupled dual membrane reactor to investigate the concept of combining exothermic reaction (methanol synthesis and methanol dehydration) with

endothermic reaction (cyclohexane dehydrogenation) for energy efficient operation. The multi-tubular reactor was combined with two water permeable membranes and the membranes were fitted at the inner and outer tubes for water removal from the exothermic side of the reactor. Bakhtyari et al. [99] studied the simultaneous synthesis of DME and methanol from syngas along with methyl formate from methanol in a double membrane reactor with a catalytic heat exchanger reaction setup (Fig. 10). In the reaction setup, three catalytic beds were placed in the concentric tubes and two water permeable membranes were fitted at the inner and outer surfaces. Heat produced in DME formation reaction was used as heat source for the methyl formate formation from the methanol. A generic algorithm was proposed to optimize the reaction conditions to get the highest yield. Methanol conversion to DME and methyl formate enhanced respectively to 97% from 76% and 66% from 62% in the optimized TDMHR (Thermally Double Coupled Double Membrane Heat Exchanger Reactor) compared to THR (conventional Heat Exchanger Reactor). 3.8. Spherical reactors Samini et al. [100] have proposed use of a spherical membrane reactor instead of the conventional tubular reactor to address a few issues of DME synthesis from methanol such as undesirable pressure drop, high reactor fabrication cost, water generation, low production capacity, low diffusion through the catalyst bed due to large catalyst particle size. Spherical reactor with lower pressure drop configuration has the potential to overcome all above problems. Spherical reactors are 329

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of two types, radial-flow packed bed reactor (RF-SPR) and axial flowpacked bed reactor (AF-PSR). AF-PSR is superior and more efficient than the former because of better feed distribution and application of membrane inside the reaction. A unique axial–flow spherical packed bed membrane reactor (AF-SPMR) has been fabricated for methanol dehydration to DME synthesis (Figs. 11 and 12). The AF-SPMR consists of two concentric spheres and the inner sphere surface is coated with a water-permeable selective membrane. Methanol is introduced axially into the inner sphere through the packed bed. Water vapor permeated through the H-SOD membrane is removed by the sweep gas introduced in the outer sphere. This type of reactor configuration develops a lower pressure drop and generates higher yield. The operational parameters and length per radius of AF-PSR for methanol dehydration have been optimized by a differential evaluation method for achieving the highest DME yield. Total eleven process parameters have been considered for optimization including molar flow rates, inlet temperatures, inlet pressure of the reaction side, length per radius (L/R) of the reactor, composition of all components in the permeation and reaction sides. 3.9. Temperature gradient reactor A number of reactor configurations and designs have been investigated to overcome the equilibrium limitation of methanol dehydration reaction for DME synthesis. In order to restrict the produced methanol from decomposing into syngas, Berty et al. [101] proposed solvent methanol process (SMP), where the produced methanol is instantly quenched into the inert solvent in order to increase conversion and eliminate gas recycle. Selective methanol permeable membranes (Li-Nafion) have been utilized by many researchers to remove methanol from the reaction mixture. A condenser is also fitted inside the multifunctional reactor to drive the equilibrium composition towards the desired product [102]. The same strategy was opted by Li et al. [103] for DME synthesis, which showed a 30% higher DME yield. Normally in the quench type catalytic reactor, the catalyst bed temperature varies with time and the increasing temperature profile of the catalyst bed of DME synthesis reactor is optimized by injection of cool reactant gases

Fig. 13. The configuration of temperature gradient reactor (TGR) reactor. Modified from [105].

Fig. 14. Graphical illustration of (a) the stacking of reaction and heat exchange oil slits in the three different micro packed bed reactor-heat exchangers, (b) internal pillar structure inside the reaction slits of reactors A and B, and (c) pillar structure inside the slits of reactor C. Reproduced from [112].

330

Micro reactor

Multistage fluidized bed Slurry Bed

High pressure high temperature cell Dual-type catalytic fluidized bed reactor A fixed bed and plasma reactor Fixed bed reactor

Direct method

Direct method

Direct method

331

Two stage reactor system

Slurry reactior

Fixed bed reactor

Three phase slurry reactor (Autoclave)

Fixed bed micro reactor

Direct synthesis

Direct synthesis

Direct synthesis

Indirect synthesis

stainless steel micro packed bed reactorheat exchanger Micro packed bed reactor fixed-bed microreactor

autoclave reactor (Parr) Pilot scale plant with fixed bed tubular reactor vertical reactor (Pyrex vessel)

Direct synthesis

Direct synthesis

Direct synthesis

Direct synthesis

Indirect synthesis

Direct synthesis

Direct synthesis

Indirect synthesis

Direct synthesis

Direct synthesis

Direct method

Direct synthesis

Indirect method

Multitubular fixed bed reactor High-pressure isothermal fixed-bed reactor Fixed bed micro reactor

Fixed bed

Direct method

Direct method

Types of reactor

Synthesis method

260

Cu-ZnO colloidal catalysts + ℽ-Al2O3

Nano size γ-Al2O3

Cu–ZnO–ZrO2/Al-modified H-mordenite

Multi-walled carbon nanotubes promoted CuO–ZnO–Al2O3/HZSM-5.

First stage = Cu–Zn–Al Second stage = Cu–Zn–Al mixed with HB Zeolite Bi-function catalyst (C301/ P-γ-Al2O3)

Core-shell catalyst(CuO/ZnO/ [email protected]) Pd-CeO2-nMxOy/γ- Al2O3 (M = Zr2+, Ca2+, or Zn2+)

Cu/ZnO/Al2O3 hybrid with γ-Al2O3 or NaZSM5 or HZSM-5

Zeolite (BEA, MFI, FER)

320

220-270

0.1

5

3

4.3

270 262

0.98

3

4

1-5

0.1

4-4.2

5

0.2

3

2.2–2.5

3

2-4

2-4

3.3

4

5

3

3

pressure, MPa

300

300

250-280

210-270

180-300

200-230

250-400

Nano-sized γ-Al2O3

Cu/Zn/Al/HZSM-5

270-350

CuO-ZnO-MnO/SAPO-18

220-260

250/270

Cu.ZnO/ Al2O3 or Nb2O5

Cu-Zn-A1/γ-Al2O3

240-280

277

250

280

Cu–Fe–Ce/HZSM-5 with 3.0 wt% CeO2

CuO/ZnO/ Al2O3 & ℽ-Al2O3

Cu-infiltrated Zr doped SBA-15

Cu/Zn/Al slurry catalysts

250-260

250

Cu-ZnO-Al2O3- Li2O/HZSM-5

CuO–ZnO–Al2O3/HZSM-5

275

temp, °C

CuO-ZnO-MnO / SAPO-18

catalyst

Table 1 Effect of different reactors, catalysts and process parameters on selectivity for DME synthesis.

−1

Syngas(H2/CO = 2) Cat = 3 g Feed rate = 100-400 ml/min Methanol = 0.01-9.99 ml/min; GHSV = 15 h−1

Syngas(H2/CO = 2) SV = 600 mL/h·gcat H2/ CO2 = 3/1 GHSV = 1800 ml gcat−1 h−1

H2/ CO2 = 3/1

syn-gas (70% H2/24% CO/6% CO2) SV = 7.4 l syngas/gCuZn/h Bio syngas= H2 26-28 vol%ˈN2 16-21 vol%ˈCO 2631 vol%ˈCH4 3- 4 vol%ˈCO2 19-24 vol%ˈ C2˜3<0.7 vol% GHSV = 1000 h−1 mixture of methanol (0.056 mol/mol) and nitrogen (60 N mL/min) WHSV = 2 gMetOH/(gcat h) Syngas1 (H2:CO:CO2:CH4:N2) = 42:42:5:6:5 Syngas2 (H2:CO:CO2:CH4:N2) = 56:28:5:6:5 SV = 150–800 N cm3/min/g catalyst (H2: CO : CO2 = 57.6 % : 28.8 % : 3.6 %) SV = 900 mLSTPg cat–1 h–1 H2:CO:N2:CO2:H2S = 50:25:19.95:5:0.5 SV = 1600 L·(kg·h)−1

MeOH = 0.84 ml/min; WHSV 20

H2/CO = 0.75-2.15; Gas superficial velocity = 0.0624 m/s Plasma catalytic reactor = 11.0–15.0 kV, H2: C02 = 4:1 GHSV = 1500–3500 mL.gcat−1 h−1 H2/CO2 = 3 SV = 10 h−1 Biomass to Syngas GHSV 1200−h H2/CO = 3; space time (2.5–20 gcat h (molC)−1)

H2:CO2 = 3:1

H2:CO = 1:1, Syngas

H2:COx = 3; CO2:COx = 0-0.5; Space time = 5.2 gcat. h (molC)−1 H2/CO/CO2/N2 = 36/36/18/10 (vol%) SV = 2400 ml/gcat/h H2:CO2 = 1; Sv = 3000 ml/gcat/h;

Others essential factor

CMeOH = 84% SDME = 82%

CCO = 93% SDME = 60% CCO2 = 46.2% SDME = 45.2% YDME = 20.9% CCO = 68% SDME = 82%

CCOx = 63.2% SDME = 68.5% Time space yield = 30.92 mmol.(ml.h)

CMeOH = 82.6% SDME = 99.9% YDME = 82.5% CCOx = 66% SDME = 68%

CCOx = 36% YDME = 13%

CCO2 = 69.5% YDME = 16.9% CCO2 = 9% S DME = 31% CCOx = 45%

CCO = 61%

CCOx = 47% YDME = 21. CCOx = 61.1% YDME = 38.7% CCOx = 48.6% YDME = 32.67% CCO = 54.32% SDME = 69.74% CCO2 = 15%

Conversion & yield

Rahmanpour et al. [134]

Khandan et al. [133]

Zha et al. [132]

Yin et al. [131]

Fujiwara et al. [130]

(continued on next page)

–1

Meng et al. [129]

Ding et al. [128]

Dadgar et al. [127]

Catizzone et al. [126]

García-Trenco et al. [124] Yuping et al. [125]

Hosseini et al. [123]

Ateka et al. [122]

Wang et al. [121]

Silva et al. [120]

Su et al. [119]

Atakan et al. [118] Yousefi et al. [31]

Abashar et al. [51] Sun et al. [117]

Zuo et al. [116]

Ateka et al. [115]

Authors

U. Mondal and G.D. Yadav

Journal of CO₂ Utilization 32 (2019) 321–338

three integrated micro packed bed reactorheat exchangers Down flow Fixed bed reactor Fixed bed reactor

Direct synthesis

Slurry reactor

Fixed bed reactor with a catalytic distillation column Fixed bed reactor

Two neck continuous stir tank reactor

Slurry reactor with a mechanical magnetic stirrer Fixed bed micro reactor Slurry reactor

Indirect Synthesis

Indirect synthesis

Indirect synthesis

Direct

332

Fluidized bed reactor

Fixed bed reactor

shell-and-tube type fixed bed reactor shell-and-tube type fixed bed reactor Autoclave reactor

Temperature gradient reactor Autoclave reactor

Autoclave reactor

Slurry autoclave reactor Fluidized bed reactor

Direct synthesis

Indirect synthesis

Direct synthesis

Direct synthesis

Direct synthesis

Direct synthesis

Direct synthesis

Direct synthesis

Direct synthesis

Direct synthesis

Direct synthesis

high-pressure tubular stainless steel fixedbed reactor Slurry phase autoclave reactor

Direct synthesis

Direct synthesis

Direct synthesis

Direct synthesis

Slurry reactor

Indirect synthesis

Direct synthesis

Direct synthesis

Types of reactor

Synthesis method

Table 1 (continued)

(CuO-ZnO-Al2O3-Ga2O3-MgO) and (γ-Al2O3 or ZrO2·Al2O3)

C301(Cu/ZnO/Al2O3) and γ-Al2O3 Cu − ZnO − Al2O3/HZSM-5

Cu-Zn-X oxide and γ-alumina (X = Al-Ti-Nb-V-Cr) Mn was added by impregnating co-precipitated CuZnAl and γ-Al2O3 Cu/ZnO/ Al2O3 and γ-Al2O3

CuO − ZnO − Al2O3/γ-Al2O3

CuO/ZnO/Al2O3 for methanol forming, γAl2O3 for methanol dehydration Cu–Zn–Cr/ γ- Al2O3

230-270

4

3

5

7

250 260

5

1

1-4

4.13

5

1.6

3

2-5

5.

5

4.13

4

0.1

4

2

0.1

3

5

3

5-7

pressure, MPa

260

100-289

225-325

240-270

280

300

260

Attrition resistance CuO/ZnO/Al2O3- HZSM-5

Commercial γ-Al2O3

200-240

260

260

280

280

180,200, 220

CuO/ZnO/ Al2O3 + H-ZSM-5

a Cu-based methanol synthesis catalyst (C301 MSC)/ γ-Al2O3methanol dehydration CuO − ZnO − Al2O3 / SiO2−Al2O3

Hydrotalcite derived Cu-Zn-Cr + γ- Al2O3

Ionic liquids incorporating an alkane sulfonic acid as a part of the cation, a complex acidic anion, [A2H]−, or both CuZnAl slurry catalyst

260

118-150

Macro porous sulphonic acid ion exchange resign

CuO–ZnO–Al2O3 /HZSM-5 with various contents of antimony oxide (0–30 wt.%)

230-290

250

250

250

220-320

temp, °C

Slurry catalyst (AlOOH)

Highly porous nanocrystalline alumina

CuO-ZnO-Al2O3-La2O3/HZSM-5 with various La loadings Pd/CNT-promoted Cu-ZrO2 / HZSM-5

physical mixtures of CuO–ZnO–Al2O3 and γAl2O3

catalyst

H2:CO:CO2:CH4 = 37.4 : 46.3 : 7.7 : 8.6 Flow rate = 1 SLPM H2/CO = 1 SV = 10,000 h−1 H2/CO = 1 GHSV = 3000 C GHSV = 6000 h−1

H2/CO/CO2/N2) 60/30/ 5/5) GHSV = 500 mmol h-1 g of cat.-1 syngas (H2/CO = 2.1)

(H2 + CO) and (H2 + CO2) = 4:1

H2/CO ≈ 1-2 GHSV = 3000 ml/(g·h) MeOH flow rate = 0.55 mL/min WHSV = 26.07 h−1 H2/CO ratio = 1.5:1 GHSV = 2000 h−1 H2–CO ratio of 1.5, GHSV of 6000 h-1

H2/CO ≈ 1-2 SV = 500-1300 mLn (g of catalyst)−1 h−1

H2/CO = 1.5 GHSV = 6000 h−1 H2/CO ≈ 2 GHSV = 4000 h−1 Syngas comp. = 47% CO, 47% H2, and 6%CO2. GHSV = 5000 ml gcat−1 h−1

H2/CO = 1:1 GHSV = 250 mL/gcat h

The syngas (premixed) H2:CO:CO2:N2 = 61.4:28.5:2.8:7.3 (vol.%) GHSV = 1500 ml h−1 gcat−1

MeOH = 0.20 ml min−1

MeOH flow rate = 0.1 ml/min

Syngas = (H2:CO:CO2:N2:CH4 = 56:28:5:5:6 (mol%) GHSV = 7500 Nml/gcat/h 5 vol.% N2, 23.7 vol.% CO2 and 71.2 vol.% H2 GHSV = 3000 h−1 H2 / CO2/N2 = 69/23/8 GHSV = 25,000 mLSTP/(h g-hydr. catal.) MeOH = 1.2 mol

Others essential factor

CCO = 65%

CCO = 67.5% SDME = 74.1% SDME = 95%

CCO = 70%

CCO = 89.1% SDME = 57.6% YDME = 50.4% CCO = 68% SDME = 46% CMeOH = 90% SDME = 100% CCO = 39.4% SDME = 50.2% CCO = 50% SDME = 70% SDME= >60%

CCO = 20% SDME = 96%

CCO = 63% YDME = 43% CCO = 52.6%

CCO = 33% SDME = 93%

CCO = 70% YDME = 95%

CCO2 = 43.8% SDME = 71.2% CCO2 = 18.9% SDME = 51.8% CMeOH = 92.3% SDME = 66.4% CMeOH = 84.4% SDME = 100% CMeOH = 35%

CCOx = 70-80% YDME = 25%

Conversion & yield

(continued on next page)

Lu et al. [151]

Jia et al. [150]

Lee et al. [149]

Venugopal et al. [113] Aguayo et al. [147] Omata et al. [106] Tan et al. [148]

Mollavali et al. [146] Song et al. [4]

Teng et al. [145]

Moradi et al. [144]

Naik et al. [143]

Venugopal et al. [113] Wang et al. [142]

Gao et al. [141]

Atkins et al. [140]

Mao et al. [139]

Lei et al. [68]

Zaherian et al. [137] liu et al. [138]

Wengui et al. [135] Zhang et al. [136]

Hayer et al. [112]

Authors

U. Mondal and G.D. Yadav

Journal of CO₂ Utilization 32 (2019) 321–338

Journal of CO₂ Utilization 32 (2019) 321–338

Shikada et al. [154]

Ng et al. [5]

Moreno-Castilla et al. [153]

for better conversion. Omata et al. [104] have designed and applied a temperature-gradient reactor (TGR) to eliminate the equilibrium conversion limit for DME synthesis at a higher temperature and the low activity of catalyst at a low temperature. In TGR temperature of the catalyst bed gradually decreases along the length of the reactor with downflow of the reaction gas mixture. Equilibrium conversion is achieved fast at the high-temperature zone near the inlet and then the conversion increases gradually along the temperature-equilibrium curve as the reaction mixture flows through the low-temperature zone. The catalyst bed is divided into a series of five zones and all the zones are optimized for the highest syngas conversion. At 3 MPa of operational pressure and fixed temperature profile of TGR, they have simultaneously achieved high one pass conversion of syngas and high space-time yield of DME of 1.1 kg-MeOH-equiv.kg-cat−1h−1. For improving the CO conversion at lower pressure, the reactor temperature needed to be optimized by adopting an artificial neuron network (ANN) and genetic algorithm (GA). The five zones of the catalyst bed were encoded as “gene,” and the fitness of each gene was evaluated by the CO conversion achieved according to the temperature set to the gene. After adopting the GA and ANN, the temperature profile was optimized and accelerated to achieve 71% conversion at 1 M P pressure [105]. Omata et al. [106] adopted an improved catalyst and optimized TGR; one pass conversion of CO at 82% was attained at 1 MPa, W/F = 50 gcat·h/mol (Fig. 13). 3.10. Micro reactor

5

There are multiple advantages of miniaturization of chemical reactors, such as high mass transfer and heat transfer rates due to increase in surface area to volume ratio and very short characteristic dimensions [68]. Therefore microchannel reactors are often used for the study of both exothermic (Fischer-Tropsch reaction) and endothermic reactions (methane reforming) [107,108]. Another advantage of the microchannel reactor is the better control of reaction process even under failed operation (leak, run away), as microreactor occupies a small volume and a small number of chemicals inside the microchannels of the rector. These safety features enable the use of the microchannel reactor for high-risk processes [109,110]. Microchannel reactor has become a unique solution for those processes which were previously operated in batch mode. The continuous mode of operation and ease of scale-up from laboratory scale to plant scale is easy. Catalytic bed and packed bed micro-reactors are two types of microchannel reactors, which have been investigated for gas-solid heterogeneous reactions [111]. Hayer et al. [112] have done a detailed study of differently packed bed micro-reactors connected with heat exchangers for the synthesis of DME and the effect of a cross-section of slits, the number of reaction slits and the slit geometry on the conversion and selectivity have been studied (Fig. 14). They have achieved 95% of CO conversion and 55% of DME yield. Venugopal et al. [113,114] have investigated DME synthesis in a microreactor with different catalyst compositions and achieved total 70–80% of CO conversion and 25% of DME yield with Cu-Zn-Cr/γ-Al2O3 and 70 of CO conversion and 45% of DME yield with Cu-Zn-Al-Y/γ-Al2O3.

CuO-ZnO-Al2O3 and y-alumina-supported copper Slurry bubble column reactor Direct synthesis

260-280

5 250 Cu − ZnO − Al2O3/ γ-Alumina (2:1) Internal recycle reactor Direct synthesis

0.1 Plug flow micro reactor Indirect synthesis

180

MeOH and He gas mixture = 3.6 vol% of MeOH Flow = 63 ml/min SV = 0.41−1 Syngas composition COx : H2 ratio = 1 : 1 to 1 : 4 (CO = 16.2%, CO2 = 1.8%, H2 = 72% and He = 10%) SV = 27,500 h−1 H2/CO = 1

Hirano [152] CMeOH = 40% SDME = 40% 0.6

γ-Al2O3 modified with (ZrO2・Al2O3, SiO2・Al2O3, SiO2・TiO2, ZrO2・TiO2), and ZSM-5 Cu/ZnO/AI2O3 and γ-Al2O3 Activated carbon oxidised with (NH4)2S O8 and HNO3 1. Pulse Reactor 2. Micro flow reactor Indirect synthesis

250-300

CH3OH:H2O = 1:1 SV = 3000−1

Authors catalyst Types of reactor Synthesis method

Table 1 (continued)

temp, °C

pressure, MPa

Others essential factor

Conversion & yield

U. Mondal and G.D. Yadav

4. Comparison of different reactors Table 1 presents the types of catalysts and process parameters and reactors used for DME synthesis. Thereafter a detailed comparison of different reactor systems is presented in Table 2. A further better understanding of the direct DME synthesis process employed in different reactors is needed, which should focus more on heat recovery generated in the exothermic reaction, and catalyst performance, recovery and reusability. A graphical representation of the number of reactor systems that have been used for DME synthesis irrespective of the raw material is shown in Fig. 15. 333

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Table 2 Comparison of different reactors systems have been used for DME synthesis. Reactor Type

Operation

Advantages

Disadvantages

Fixed Bed reactor (FBR)

Catalytic heterogeneous gas phase reaction. Reactions with low or intermediate heat generation are best suited.

Very simple and low-cost operation, easy to recover catalyst and reuse it.

Fixed bed adiabatic reactor

It is an isothermal reactor and the exothermal reaction can be carried out at the optimal temperature depending on the thermodynamics and kinetics of the reaction. Combination of a conventional reactor and heat exchanger. It is a special type of multiphase pneumatic contactor.

Reactor operation is very simple and the operational cost is minimum. Extracted heat can be utilized. Reactor operation at optimum temperature can lead to higher yield and more selectivity.

Not suited for exogenous reactions, hot-spot generation inside the reactor. Catalyst sintering is a common phenomenon due to excessive heat generation. High syngas recycle ratio needed for the higher conversion. High initial investment. High-pressure drop across the catalyst bed.

Slurry phase reactor

Dual type reactor

Coupling of highly exothermic and endothermic reactions to intensify both the reactions.

Catalytic distillation reactor and dividing wall column reactor

Intensification of methanol dehydration reaction by combining reaction process with the distillation process (CD or RD). Dividing-wall column (DWC) reactor is a single tower divided into two section by inserting a single wall at the proper position. Reactive dividing-wall column (r-DWC) is a combination of reactive distillation with DWC in a single vessel

Membrane reactor

The hydrophilic water permeable membrane is added inside the reactor to continuously remove water molecules generated in the reactor from the direct and indirect method of DME synthesis.

Spherical reactors

Two types of spherical reactors are used for DME synthesis from methanol dehydration reaction, such as radial-flow packed bed reactor (RF-SPR) and axial flow-packed bed reactor (AF-PSR). The temperature of the catalyst bed gradually decreases along the length of the reactor with downflow of the reaction gas mixture. Used for the study of both exothermic and endothermic reactions.

Temperature gradient reactor Microreactor

Simple construction, uniform temperature distribution ever for a high exothermic reaction, easy catalyst addition and recovery, minimum catalyst sintering and low energy requirement for adequate mass transfer between phases Reduced operation and production cost, can be used as a process intensifier for many industries, highly energy efficient. More economically beneficial, process intensifier, energy efficient with low investment cost. DWC can separate more compounds in a single unit and thus saving the cost of the additional units along with reboiler and condenser. R-RWC makes the process more energy efficient, reduces CO2 emission, and reduces operation cost. Higher DME yield. Continuous removal of water molecules prevents the deactivation of the catalyst. Packed bed catalytic membrane reactor (PBMR) has shown higher conversion of CO2 and reduces methanol loss compared to normal PBR reactor system. Thermally Double-coupled two-membrane reactor (TDCTMR) is economically viable, energy efficient and generates the highest yield of H2, MeOH and DME. AF-SPMR generates lower pressure drop and higher yield of DME. Equilibrium conversion is attained very fast; overall conversion and DME yield are high. High mass transfer and heat transfer rates, high-risk reactions also can be operated; good control of the process

5. Conclusions

Usage of traditional solid catalyst leads to several problems such as aggregation, misdistribution and rheological deterioration of catalyst particles inside the reactor and thus catalyst also get deactivated and poisoned.

CD requires moderate reaction temperature but conventional methanol dehydration catalyst gets activated at a higher temperature.

Sometimes produced HCs can block the pores of the membranes. Thermal and mechanical stability is a minor issue.

Sometimes produced HCs can block the pores of the membranes. Thermal and mechanical stability is a minor issue. Complicated reactor setup; reactor operation and maintenance are high. Laminar flow of reactants inside the microchannels.

for high exothermic reaction, (iii) easy catalyst addition, and separation process, (iv) good control of the reactor temperature to avoid catalyst sintering, and (v) low energy requirement for adequate mass transfer between phases. There are some innovative reactor systems that have recently been adopted for DME synthesis such as catalytic distillation, spherical reactor, dual type reactor, reactive dividing-wall column (RDWC). Dual type reactors are very energy efficient as the exothermic DME synthesis reaction can be coupled with another endothermic reaction and eventually the DME yield will be enhanced. Microchannel catalytic reactors can be an ideal solution for downsizing and intensifying DME synthesis process. Improving the catalyst activity is also an important aspect of process intensification. It is envisioned that DME would be an ideal fuel and feedstock for some chemicals provided convenient and cost effective process technologies are provided. This review will encourage further research and innovation.

As the consumption of oil has increased by many folds, DME synthesis from coal, petroleum and biomass has become very important for energy needs. In the past few decades, significant advancements in DME synthesis have been achieved via a combination of reactor and catalyst development. Syngas is a major requirement of any DME industry, and the derivation of syngas from a variety of natural resources requires an efficient technology. Some reactor systems have been experimented for DME synthesis via direct or indirect route. The fixed bed reactor is the most used system for DME synthesis for its simplicity. For large-scale DME production in fixed bed reactor, a high investment cost and huge energy input are required. There are several criteria for a reactor system to be ideal for DME synthesis process such as: (i) simple construction, (ii) uniform temperature distribution inside the reactor

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Fig. 15. Usage of different reactor systems used for DME synthesis in industry.

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