Accepted Manuscript

Title: Dimethyl Ether: A Review of Technologies and

Production Challenges

Author: Zoha Azizi Mohsen Rezaeimanesh Tahere Tohidian

Mohammad Reza Rahimpour

PII: S0255-2701(14)00126-3
DOI: http://dx.doi.org/doi:10.1016/j.cep.2014.06.007
Reference: CEP 6436

To appear in: Chemical Engineering and Processing

Received date: 15-1-2014

Revised date: 24-4-2014
Accepted date: 15-6-2014

Please cite this article as: Z. Azizi, M. Rezaeimanesh, T. Tohidian, M.R. Rahimpour,
Dimethyl Ether: A Review of Technologies and Production Challenges, Chemical
Engineering and Processing (2014), http://dx.doi.org/10.1016/j.cep.2014.06.007

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Dimethyl Ether: A Review of Technologies and Production Challenges

Zoha Azizia,b, Mohsen Rezaeimaneshb, Tahere Tohidiana, Mohammad Reza Rahimpoura,*

a
School of Chemical and Petroleum Engineering, Department of Chemical Engineering, Shiraz University, Shiraz 71348, Iran

b
Department of Chemical Engineering, College of Chemical Engineering, Mahshahr Branch, Islamic Azad University, Mahshahr, Iran

t
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Abstract

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Dimethyl ether (DME) is a well-known propellant and coolant, an alternative clean fuel for

diesel engines which simultaneously is capable of achieving high performance and low emission

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of CO, NOx and particulates in its combustion. It can be produced from a variety of feed-stocks

an
such as natural gas, coal or biomass; and also can be processed into valuable co-products such as

hydrogen as a sustainable future energy. This review, which also can be counted as an extensive,
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pioneer review paper on this topic, presents recent developments in synthesis methods of

dimethyl ether as an alternative energy while focuses on conventional processes and innovative
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technologies in reactor design and employed catalysts. In this context, synthesis methods are
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classified according to their use of raw material type as direct and indirect methods as well as
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other routes, since different methods need their own operating condition. Also, the available data
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for the selectivity to dimethyl ether (DME) and yield of DME as a function of H2/CO and CO2

content of the feed is discussed.

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Keywords:

Dimethyl Ether, Direct Synthesis, Indirect Synthesis, Syngas Conversion, Catalyst, Reactor.

*
Corresponding author. Tel: +98 711 2303071; fax: +98 711 6287294;
E-mail address: rahimpor@shirazu.ac.ir (Prof. M. R. Rahimpour).

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Page 1 of 93
Table of Contents:

1. Introduction ............................................................................................................................. 3

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1.1. Scope of the current review .............................................................................................. 4

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2. Synthesis methods .................................................................................................................... 5

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2.1. Indirect synthesis method ................................................................................................. 6
2.2. Direct synthesis method.................................................................................................... 7

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2.3. Comparison of methods and seeking other routes............................................................ 9
3. Different types of DME reactors ........................................................................................... 10

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3.1. Conventional types ......................................................................................................... 11
3.1.1. Fixed-beds ................................................................................................................... 11
3.1.2. Slurry phase reactors................................................................................................... 11
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3.1.3. Fluidized-bed reactors................................................................................................. 12
3.2. Innovative technologies .................................................................................................. 13
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3.2.1. Coupled and dual type reactors................................................................................... 13

3.2.2. Coupling the reactor and separation units.................................................................. 15
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3.2.3. Micro-reactors............................................................................................................. 16
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3.2.4. Membrane reactors...................................................................................................... 17

3.2.4.1. Spherical reactors..................................................................................................... 19
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3.3. Comparison of different reactors ................................................................................... 20

4. Catalysts ................................................................................................................................ 20
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4.1. Catalyst diversity ............................................................................................................ 20

4.2. Catalyst preparation....................................................................................................... 27
4.3. Surface acidity of methanol dehydration catalysts......................................................... 28
4.4. Catalyst deactivation ...................................................................................................... 29
4.5. Comparison of different catalysts................................................................................... 32
4.5.1. Activity ......................................................................................................................... 32
4.5.2. Yield and selectivity ..................................................................................................... 34
4.5.3 Deactivation.................................................................................................................. 35

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Page 2 of 93
5. Essential factors affecting the performance of DME production.......................................... 36
5.1. Water removal ................................................................................................................ 36
5.2. H2/CO ratio and CO2 content of the feed ....................................................................... 38
5.3. Operational temperature ................................................................................................ 39
5.4. Operational pressure...................................................................................................... 41

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5.5. Space velocity ................................................................................................................. 41

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6. Process intensification (PI) ................................................................................................... 43

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7. Conclusions and future perspectives ..................................................................................... 45

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1. Introduction

The inordinate use of oil-based fuels for transportation purposes is one of the major reasons

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of the rapid depletion of petroleum which causes major environmental problems. These issues

have necessitated the development of clean non-petroleum based alternative transportation fuels.
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In recent years, the application of dimethyl ether (DME) as a potential diesel substitute used in
d

compression ignition engines has attracted considerable attention [1, 2]. DME is a volatile
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substance which forms a liquid phase when pressurized above 0.5 MPa; therefore, it is

commonly handled and stored as liquid (see the physical property of DME in Table 1). Burning
p

with a visible blue flame and with similar properties as propane and butane, DME may hence be
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used as liquefied petroleum gas (LPG) for heating and home cooking [3]. For many reasons,
Ac

DME is known to be a clean fuel: 1) Unlike other homologous ethers, it has a safe storage and

handling as it does not form explosive peroxides [4]. 2) Since DME only has C-H and C-O bond,

but no C-C bond, and since it contains about 35% oxygen, its combustion products such as

carbon monoxide and unburned hydrocarbon emissions are less than those of natural gas. 3)

Owing to its high cetane number, DME is considered to be an excellent alternative to the present

transportation fuel with no emission of particulate matter and toxic gases such as NOx at burning

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Page 3 of 93
[1, 2, 4, 5]. 4) Moreover, it has a similar vapor pressure to that of LPG, and hence can be used in

the existing infrastructures for transportation and storage [6]. Thus, the significant future

perspective of DME can be counted as an alternative energy.

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Furthermore, DME is widely recommended as environmentally friendly aerosol and green

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refrigerant since it has zero ozone depletion potential and lower global warming potentials

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compared with traditional chlorofluorocarbons (CFCs, Freon) and R-134a (HFC-134a) [7]. In

addition, DME can be used as pesticide, polishing agent, and anti-rust agent. It can also be

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considered as an attractive material for producing alkyl-aromatics, a suitable source for the

an
hydrogen used in fuel cells, as well as a key intermediate for producing dimethyl sulfate, methyl

acetate, light olefins, and so many other important chemicals [8, 9, 10].
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Table 1
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DME can be produced from a variety of feed-stock including natural gas, crude oil, residual
te

oil, coal and waste products [11]. Among potentially interesting raw materials, natural gas

appears to be the most promising one due to its wide availability and the fact that producing
p

DME from natural gas allows production costs to be independent of the swings in the oil price
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[12].
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1.1. Scope of the current review

Extensive works have been undertaken to improve DME synthesis methods and the

employed catalysts, but the DME subject is still suffering from the lack of a critical review. The

underlying goal of this paper is to present an extensive review considering the valuable works

accomplished over the years 1965-2013 on dimethyl ether synthesis. The trend of related

publications on DME over the years 1996-2013 is shown in Fig.1. As obvious, the number of
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Page 4 of 93
published papers rises gradually and has a peak within the year of 2011. It is concluded from

Fig.2 that these publications are drastically concentrated on catalyst and then on reactor

technology. Thus, taking these results into account, this paper focuses on production methods

and a discussion on their wide variety of reactors and catalyst configurations while also

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investigating effective parameters including water removal, H2/CO ratio, CO2 content of the

feed, temperature, pressure, and space velocity.

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Fig.1

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Fig.2

2. Synthesis methods
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M
As shown in Fig.3 and mentioned before, DME can be produced in two distinct ways: the

first called the indirect route uses the produced methanol to promote its dehydration; the second
d

way which is arguably more efficient is known as the direct route, in which DME is produced in
te

a single stage using bi-functional catalysts. The technology of this single step method belongs to
p

companies such as Haldor Topsoe, JFE Holdings, Korea Gas Corporation, Air Products, and
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NKK [13, 14]. Moreover, Toyo, MGC, Lurgi and Udhe have their own indirect processes for

DME production [13].

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Fig.3

One of the main steps in DME synthesis is the production of syngas, which is a mixture of

hydrogen and carbon monoxide and has been manufactured industrially from hydrocarbon fuels-

typically natural gas- either by steam reforming (SR) or gasification [15, 16, 17].

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Page 5 of 93
 In the following sections, essential information about two basic methods of DME synthesis

will be presented and other routes of DME production will then be pointed out.

2.1. Indirect synthesis method

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Traditionally, DME has been produced from syngas in a two-step process in which

methanol is produced from syngas, purified, and then converted to DME in another reactor. The

cr
schematic of this process is shown in Fig.4. The commercialized process reaction of DME

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production from methanol dehydration is shown in Eq.(1):

an

2CH3OH → CH3OCH3 + H2O H 298 K  23 .5kJ / mol (1)

Fig.4
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Many investigations on the kinetics of DME synthesis by dehydration of methanol on solid-
d

acid catalysts have been published. The majority of them agree that the mechanism follows
te

either Langmuir-Hinshelwood [18] or Eley-Rideal kinetic models [19], with water and DME

both acting as reaction inhibitors [20].

p
ce

Above 250°C, the rate equation related to Eq.(1) is given by Bondiera and Naccache [21] as:

 E 
Ac

 rmethanol  k 0 exp   a  p methanol (2)

 RT 
Where k0 = 1.21×106 (pressure of methanol (kPa).kmol)/(m3 reactor h kPa), Ea = 80.48

kJ/mol, and pmethanol is the partial pressure of methanol (kPa).

Theoretically, methanol dehydration is favored at lower temperatures because it is an

exothermic reaction and the formation of by-products such as ethylene, carbon monoxide,

hydrogen, and/or coke is significant at higher temperatures.

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Page 6 of 93
2.2. Direct synthesis method

More recently, a combined methanol synthesis and dehydration process has been developed

to synthesize DME directly from syngas in a single reactor [22].

t
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The direct synthesis of DME from syngas containing H2, CO and CO2 follows mainly two

overall reactions: Eq. (3) with water-gas shift reaction taken into account and Eq.(4) without it.

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3CO + 3H2 → CH3OCH3 + CO2 (3)

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2CO + 4H2 → CH3OCH3 + H2O (4)

The Eq. (3) involves four basic reactions including [23]:

Methanol synthesis from CO:

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M

CO + 2H2 ↔ CH3OH H 298 K  90 .4 kJ / mol (5)

Methanol synthesis from CO2:

d


CO2 + 3H2 ↔ CH3OH + H2O H 298 K  49 .4 kJ / mol (6)
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Water gas shift (WGS):

p


CO + H2O ↔ CO2 + H2 H 298 K  41 .0 kJ / mol (7)
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Methanol dehydration:


H 298 K  23 .0 kJ / mol
Ac

2CH3OH ↔ CH3OCH3 + H2O (8)

In each reaction, the equilibrium conversion reaches its maximum peak whenever the

H2/CO ratio in the feed stream corresponds to the stoichiometric value, that is 1.0 for Eq. (3) and

2.0 for both Eq.(4) and Eq.(5).

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Page 7 of 93
 According to the aforementioned reactions, both methanol synthesis and WGS reactions

occur in the syngas to DME (STD) process. The former is crucial for DME production and the

latter produces CO2 which is the main by-product [24]. CO2 can be used in methane reforming

unit to produce syngas based on the reaction stoichiometry of Eq.(9) with H2/CO molar ratio of

t
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1.

cr
2CH4 + O2 + CO2 → 3CO + 3H2 + H2O (9)

The overall reaction of STD process is highly exothermic and, therefore, the temperature of

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the process should be controlled properly in order to avoid run-away. Although the direct

an
synthesis method has minimum waste of natural gas, it is one of the most complicated chemical

reactions of methane conversion. Operational units of chemical engineering are commonly used
M
for the separation and purification of DME in the synthesis process. Through absorption, flash

and distillation [25, 26], H2, N2, CH4, and CO2 are removed; methanol is recovered and the final
d

DME product is obtained. Moreover, since the methanol synthesis is a thermodynamically

te

limited process, consumption of methanol in the consequent reaction to form DME will shift the
p

methanol synthesis equilibrium towards higher methanol conversion [27, 28]. The separation of
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DME and CO2 becomes more difficult when methanol is present in the system. Thus, in a

proposed process [29], the methanol and water resulted from the one-step reaction were first
Ac

condensed and then absorbed by water; finally the liquid stream containing DME was distilled

for final DME product. This separation process is demonstrated to be feasible to synthesize DME

with high purity. The schematic of this process is shown in Fig.5.

Fig.5

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Page 8 of 93
2.3. Comparison of methods and seeking other routes

Compared with the methanol dehydration process for DME synthesis, the direct process

allows for a higher CO conversion and a simple reactor design that results in much lower DME

t
production costs [30]. However, the separation process for high purity DME is relatively more

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complex due to the presence of unreacted syngas and produced CO2 in the one-step synthesis

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process.

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In addition, because of the water-gas shift reaction which consumes stoichiometric amount

of CO to form CO2 and hydrogen, DME synthesis directly from syngas is not much suitable for

an
commercialized purposes. Generally, the syngas process for DME synthesis is an energy

consuming and a greenhouse gas emitting process. More energy efficient and environmentally
M
friendly design solutions are desired to that end [31]. These processes are integrating the DME

synthesis line with hydrocarbon reforming units, recycling and valorizing CO2 byproduct.
d
te

In this regard, considerable attention has been given in literature to the use of CO2 as a raw

material in the synthesis of chemicals and liquid energy carriers in order to mitigate the
p
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accumulation of CO2 in the atmosphere [8, 32, 33]. However, the application of the conventional

process is not yet industrially acceptable owing to low CO2 conversion, low DME yield, and
Ac

selectivity. Combining the equilibrium limited reaction with the selective removal of H2O allows

for an increase in CO2 conversion and could be an interesting and effective way to by-pass

thermodynamic limitations of DME synthesis [34].

According to a novel method, DME can be synthesized from methane through a two-step

process in which a methyl halide (usually CH3Cl or CH3Br) is prepared from the oxidative

bromination reaction of methane in the presence of hydrogen halide and oxygen over a Rh-SiO2

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Page 9 of 93
catalyst. In the second step, methyl halide is further hydrolyzed to DME over a silica supported

metal chloride catalyst [31, 35]. The main by-product of this process is methanol while the major

problem associated with the hydrolysis of organic halides is corrosion. Surya Prakash et al. [36]

studied the hydrolysis of methyl bromide in the presence of PVP as a new catalyst in a batch

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reactor. Use of this catalyst as a potential reusable solid amine catalyst showed maximum

efficiency according to the capture of HBr by solid PVP. The major advantage of this process is

cr
that the polymer can be easily regenerated and reused without loss of activity.

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Dimethyl ether can also be produced through oxidative carbonylation of methyl bromide.

an
The catalyst of this process can be either SbF5/graphite or metal oxide catalyst. The former

produces methyl acetate as by-product while the latter produces methyl alcohol [36, 37].
M
3. Different types of DME reactors
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Pinch analysis designs a process to minimize energy consumption and enhance energy
te

efficiency. In recent years, some works have been done to minimize generated entropy during a

chemical process by focusing on the reactors that are at the heart of chemical process
p
ce

technologies. Thus, reactors play a principal role in pinch analysis [38]. Hence, in the following

sections we will start our discussion on reactor design by considering the conventional fixed bed
Ac

and slurry phase reactors used for DME production, and then we will investigate the subsequent

innovative technologies developed to solve the problems that might arise with the conventional

techniques.

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Page 10 of 93
3.1. Conventional types

3.1.1. Fixed-beds

Owing to simplicity and lower costs, the reactors most commonly used either at laboratory

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ip
or pilot scale are fixed-beds [39]. For catalytic processes which have low or intermediate heat of

reactions adiabatic fixed-bed reactors illustrated in Fig.6 can be the first choice [40]. In such

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systems, diffusional restrictions between phases are eliminated by gas-solid contactors [41].

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Moreover, operation in a fixed bed reactor is an interesting alternative that allows for the use of

an optimum longitudinal profile of temperature from the inlet to the outlet of the reactor. Thus,

an
the reaction rate is high near the inlet (conversion is far from that limited by thermodynamics);

and by decreasing temperature along the reactor high conversion is attained at the outlet [41].
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However, in the case of highly endothermic or exothermic reactions, there is the problem of
d

reaction putting out or catalyst sintering [42]. Therefore, the challenges of thermodynamic
te

limitations and excessive catalyst deactivation in conventional fixed bed reactors have led the

DME reactor to operate at a high syngas recycle rate in order to avoid temperature rise that might
p

further result in a lower per-pass conversion as well as larger capital investment and operating
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costs [43, 44].

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Fig.6

3.1.2. Slurry phase reactors

Besides fixed beds, the other type of reactor commonly used in commercial direct DME

synthesis technology is slurry phase reactors [42]. In three-phase slurry reactors, synthesis gas is

dispersed as the bubble phase in a solvent used to suspend the catalyst. Having the merits of

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Page 11 of 93
lower investment and better heat transfer, one-step slurry phase DME synthesis has been known

as a potential process for large-scale DME production. For DME synthesis, syngas should be

transferred from gas bubbles to liquid phase solvent and then to catalyst particles. This process

causes severe limitations in mass transfer between phases and consequently decreases the overall

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reaction rate. While controlling the reactor temperature is much more manageable in slurry

reactors than in adiabatic fixed-bed reactors owing to the large heat capacity of the solvent, some

cr
disadvantages have been reported [45, 46, 47, 48, 49, 50]. For instance, the equipment required

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in slurry reactors is complicated because in addition to the main reactor body, a recycling system

and a gas-liquid separator are needed [41]. Moreover, the loss of catalyst particles from the

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reactor is another challenge that limits the reactor’s use in DME production [46, 51]. The

optimum values for temperature and space velocity can be obtained according to Papari et al.
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[52]. They also reported that rising pressure and catalyst concentration can enhance the reactor
d

performance.
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3.1.3. Fluidized-bed reactors

p

Fluidized-bed reactors have been suggested by some researchers as a perfect reactor

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configuration for DME synthesis [53, 54]. These reactors are at the initial stage of laboratory

testing, and their feasibility has not yet been established [13]. Fluidized bed reactors show better
Ac

heat removal characteristic owing to freely moving catalyst particles in the bed. Because of the

intensive mixing of catalyst particles gas-solid mass transfer resistance decreases, thereby

achieving an excellent temperature control. In addition, achieving high conversion without the

need for recirculation and under moderate operating pressure is another benefit [41]. However,

collision between catalyst particles and the reactor wall causes loss of catalyst [55].

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Page 12 of 93
3.2. Innovative technologies

To reduce both capital and operating costs and to increase energy efficiency, process

integration can be considered. The multifunctional reactor integration can be used, for example,

t
for coupling exothermic and endothermic reactions, or coupling reaction and separation units.

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3.2.1. Coupled and dual type reactors

cr
In this type of reactors, the exothermic reaction becomes the heat source for the endothermic

us
reaction(s) [56, 57, 58]. Some investigators suggest an industrial dual-type reactor for producing

an
DME directly from syngas [59, 60]. In this regard, Vakili et al. [61] investigated the design

parameters and operating conditions of their proposed dual-type reactor. In the dual-type reactor,
M
the cold feed entered the tube side of the second reactor and was preheated by the reacting gas

that flew in the shell side. The boiling water in the shell side of the first reactor absorbed the heat
d

of exothermic reactions and produced water vapor. The production capacity of the proposed
te

reactor configuration (see Fig.7) was estimated to be the same as that of the large-scale
p

commercial DME reactor based on the indirect method. Their results made it clear that the use of
ce

a counter-current configuration in the second reactor was better than a co-current mode owing to

more DME production rate. According to the simulation results, the process design based on the
Ac

proposed optimal reactor configuration could produce about 60 ton/day DME more than the

conventional DME plant. Furthermore, this new configuration reduced DME production costs by

eliminating a separate unit for methanol production and purification [61, 62, 63].

Fig.7

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Page 13 of 93
 Moreover, on the basis of simulation results, Vakili et al. [64] showed that changing the

molar flow rate of the exothermic side could control hot spots in the aforementioned thermally

coupled heat exchanger reactor.

t
In the work of Khademi et al. [65], optimal operation conditions for a thermally coupled

ip
reactor in which simultaneous DME synthesis and cyclohexane dehydration occurred have been

cr
evaluated. The reactor, as depicted in Fig.8 and Fig.9, consisted of two separate sides for

exothermic and endothermic reactions. Catalytic dehydrogenation of cyclohexane to benzene

us
took place in the shell side, whereas methanol dehydration occurred inside the tube with fixed

an
bed of different catalysts on both sides. Heat is transferred continuously from the exothermic to

the endothermic reaction zone. It was shown that suitable amount of initial molar flow rate and
M
inlet temperature of both sides could provide the necessary heat to heat up the mixtures and to

drive the endothermic process at the same time. In addition, the short distance of heat transfer
d

increased the efficiency of the process [65, 66, 67].

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Fig.8
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Fig.9
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Skinner et al. demonstrated a staged reactor for ethanol dehydration to ethylene which
Ac

achieved a 95% conversion of ethanol [68]. A similar reactor configuration was applied to

dehydrate methanol into DME using a staged reactor that coupled partial oxidation in the

upstream stage with methanol dehydration in the downstream stage. The staged partial oxidation

reactor is capable of integrating decomposition and deoxygenation to upgrade the energy density

of liquid fuels for transportation purposes [69].

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Page 14 of 93
3.2.2. Coupling the reactor and separation units

Catalytic distillation (CD), also known as reactive distillation (RD), is another example of

an integrated process where the distillation column and the reactor are combined to form a single

t
unit. The advantages of using CD for methanol dehydration include higher selectivity towards

ip
DME, higher conversion, and lower operational costs compared to a single reactor [70, 71]. CD

cr
requires moderate operational temperature and pressure (40-180°C and 800-1200 kPa

us
respectively), but most of the catalysts previously studied for this reaction were solid-acid

catalysts (e.g. zeolites) which tended to be active at higher temperatures (about 250°C). Hence,

an
an extensive study is required in order to make this process operate at milder conditions [20]. A

typical RD column followed by an ordinary distillation column to recover methanol is shown in

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Fig.10.
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Fig.10
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Other innovative solutions to overcome the drawback of high-energy consumption in

distillation units are thermally coupled distillation columns [72], dividing-wall columns (DWC)
p
ce

[71], heat-integrated distillation or cyclic distillation [73]. The Petlyuk configuration, consisting

of two fully thermally coupled distillation columns [74], evolved to the practical implementation
Ac

in a DWC that split the middle section of a single tower into two sections by inserting a vertical

wall in the vessel at an appropriate position. DWC found great appeal in the chemical process

industry as it could separate more components in a single distillation unit, thereby reducing the

cost of building two columns and cutting the operating cost by using a single condenser and

reboiler. In fact, using DWC can save up to 30% in capital investment and up to 40% in

operating costs and also up to 30% in energy saving [72, 75].

15
Page 15 of 93
 On the basis of the integration concept, some researchers proposed a novel process for DME

production by methanol dehydration based on a reactive dividing-wall column (R-DWC) [76].

Fig.11 simply illustrates the path from CD to R-DWC. Kiss and Suszwalak concluded that the

RD process alone might not justify investments in revamping existing plants but that in the case

t
ip
of building a new plant the RD alternative would be a preferable choice because of its lower

footprint and milder operating conditions. They also claimed that the innovative reactive DWC

cr
process had better performance compared to the conventional or the reactive distillation process:

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significant energy saving of 12-58%, up to 60% reduced CO2 emissions and up to 30% lower

total annual costs. Consequently, the novel R-DWC process can be considered as a serious

an
alternative for the production of high-purity DME in new as well as revamped industrial plants

[76].
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Fig.11
d

3.2.3. Micro-reactors
te

Micro-structured reactors, where chemical reactions take place in channels or slit-like

p
ce

arrangements of sub-millimeter range of dimensions, are new alternatives to the synthesis of

dimethyl ether from syngas. A typical micro-reactor provides a high surface-to-volume ratio and
Ac

a short distance to the wall, thus enhancing heat and mass transfer rates greatly [14, 23]. Hence,

they are suited for both highly exothermic and endothermic reactions [5]. The microstructured

reactors also offer high controllability of the reaction conditions owing to a small holdup value

which allows for partial or total elimination of hot spots by avoiding thermal runaway, laminar

flow behavior, compactness, and parallel processibility [77].

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Page 16 of 93
3.2.4. Membrane reactors

Membranes may act as permselective barriers or as an integral part of catalytically active

surfaces [78]. In DME synthesis indirectly from methanol, if water vapor generated by the

t
catalytic reaction can be selectively removed from the reaction zone, decrease in catalytic

ip
activity can be prevented and hence a good reaction yield can be obtained even in a mild

cr
temperature condition. Further, no additional steps of dimethyl ether separation and purification

us
are required [79]. The removal of H2O during methanol dehydration by means of the membrane

concept can reduce H2O promoted catalyst deactivation and enhance DME productivity [80].

an
However, the enhancement of H2O flux through the membrane (i.e. by using higher sweep flow

rates) would also produce undesired HC [81]. Consequently, an optimum value for H2O flux
M
through the membrane should be selected.

Presently available membranes are amorphous silica, F-4SF, ZSM-5, MOR, SIL, and
d

polymeric membranes [81, 82]. However, they still suffer from pore-blockage, thermal, and
te

mechanical stability; and the dilution caused by the need for sweep gases have limited the
p

usefulness of the membrane reactor systems [83]. However, the benefits of the membrane
ce

systems have been demonstrated through a wide number of experimental and theoretical studies

[78, 84].
Ac

Another way to increase the yield of reaction is the controlled addition of reactants.

Rahimpour and colleagues showed in their work that proper addition of H2 could enhance DME

production [85, 86, 87, 88]. In light of this new concept, Mardanpour et al. have developed a

model for a shell and tube fluidized bed membrane reactor in order to synthesize dimethyl ether.

17
Page 17 of 93
In their proposed model, as depicted by Fig.12, hydrogen permeates along the reactor leading to

a better performance and efficiency [89].

Fig.12

t
Polymer/ceramic catalytic membranes in the work of Volkov et al. were prepared by

ip
deposition of a polymeric solid-acid catalyst, namely F-4SF resin (the Russian analog of Nafion),

cr
onto the internal surface of the ceramic ultrafiltration tubular membrane with an intermediate

us
selective layer of titanium dioxide (see Fig.13). They studied vapor phase dehydration of

methanol to DME on the F-4SF/ceramic catalytic membrane reactor as depicted in Fig.14.

an
Although at the temperature range of 70-120°C the F-4SF sample showed high initial activity in

the conversion of methanol to DME, the reaction was accompanied by the fast deactivation of
M
the catalyst. The authors have related this behavior to the strong adsorption of methanol and the

produced water within this temperature interval. Hence, to prevent deactivation of the catalyst
d

and to improve methanol conversion, it was proposed that the rate of alcohol desorption should
te

be increased and also selective removal of water from the reaction zone must be provided [82].
p

Fig.13
ce

Fig.14
Ac

Optimal coupling of exothermic and endothermic reactions can be feasible and beneficial

[89]. A distributed mathematical model followed by optimization for products in a thermally

coupled membrane reactor composed of three sides has been developed for methanol and

benzene synthesis by Khademi and colleagues [90, 91, 92, 93, 94]. In their proposed

configuration, methanol synthesis took place in the exothermic side which supplied the necessary

heat for the endothermic dehydrogenation of cyclohexane reaction. By co-current flow of sweep

18
Page 18 of 93
gas through the permeation side, selective permeation of hydrogen through a Pd/Ag membrane

was achieved.

Iliuta et al. evaluated the use of glycerol, a source of syngas, to produce DME in a dual bed

t
membrane reactor involving the endothermic catalytic glycerol reforming and exothermic DME

ip
synthesis process. Not only did this technology enhance thermal efficiency but also it could

cr
reduce the cost of syngas production in DME production unit. The schematic figure of the dual

bed membrane reactor is shown in Fig.15. In order to enhance the productivity of the DME

us
synthesis process operated at high CO2 feed conditions, the reactor had the capability to remove

an
water from the system by means of a hydrophilic membrane tube placed in the center of the

reactor [95].
M
Fig.15
d

3.2.4.1. Spherical reactors

te

Due to some reported disadvantages of tubular packed-bed reactors such as high pressure
p

drop along the reactor, high manufacturing cost and low production capacity, the feasibility of
ce

spherical packed-bed reactor as a novel reactor configuration was studied by Samimi et al. They

developed an axial-flow spherical packed-bed membrane reactor (AF-SPMR) for the dehydration
Ac

of methanol. During the reaction, water vapor was withdrawn by Hydroxy-sodalite (H-SOD)

membrane, a zeolite-like material, in order to shift thermodynamic equilibrium towards the

production side. The results indicated that in AF-SPMR, not only the pressure drop decreased,

but also more production of DME was obtained [80, 97].

19
Page 19 of 93
3.3. Comparison of different reactors

A comparison of aforementioned reactors is provided in Table 2. It is worth mentioning that

further engineering studies are necessary to identify design solutions for the direct DME

t
synthesis reactor, providing maximum process intensity, advanced recovery of heat generated in

ip
the process and preservation of catalyst activity. These should consider the most efficient spatial

cr
distribution of the two catalyst components, along with good temperature and composition

us
control in the space of catalyst bed.

Table 2

4. Catalysts
an
M
4.1. Catalyst diversity
d

Extensive research is conducted on finding better catalysts that have higher selectivity
te

towards DME formation and a lower tendency to generate hydrocarbons and coke. With regard
p

to DME synthesis by indirect method, the commonly employed catalysts are solid-acid types.
ce

Catalysts for the STD process are bi-functional catalysts composed of a metallic function for

methanol synthesis and a solid-acid function for the transformation of methanol into DME [41,
Ac

96]. It should be noted that the heat conduction of the bi-functional catalysts is poor; hence, the

applied working temperature of the bi-functional catalysts is in a temperature range of 523-673 K

and pressures up to 10 bar [7, 23, 98]. The metallic function is mainly composed of such oxides

as CuO, ZnO, Al2O3 and Cr2O3 [24, 41, 99]. Moreover, a myriad of solid-acid catalysts have

been explored including γ-Al2O3, modified alumina with silica, TiO2-ZrO2, clays, ion exchange

resins, Boehmite (AlOOH) and zeolites such as H-ZSM-5, HY, mordenites, SAPO, MCM,

20
Page 20 of 93
Ferrierite, chabazite and H-beta [2, 7, 10, 23, 30, 41, 100, 101, 102, 103, 104, 105, 106, 107, 108,

109]. Meanwhile, the solid-acid catalysts can be modified with sulfate, zirconium, iron, silica,

phosphorus, B2O3, and rare metals to obtain moderate acidity for higher CO conversion and

minimal by-product (light olefins and heavy hydrocarbons) formation [7, 23, 30]. For instance,

t
ip
Jin et al. prepared a series of zeolite Y modified with La, Ce, Pr, Nd through ion-exchange. It

was found that these rare earth metals resulted in Y enhanced acidity and thus exhibited higher

cr
activity and stability than did pure HY for methanol dehydration to DME [100].

us
In CuO-ZnO-Al2O3 (CZA) catalysts, metallic copper clusters are the active sites for both

an
methanol synthesis and WGS reactions and conversion of syngas to methanol depends on the

copper metal surface area [8, 23, 110, 111]. ZnO plays a pivotal role in maintaining the active
M
copper metal in optimal dispersion, thus providing a high number of active sites exposed to

gaseous reactants [2]. However, Hadipour and Sohrabi observed that excess of ZnO in CZA had
d

a negative effect on the activity [96]. The purpose of the addition of M3+ ions (e.g. Al3+) into
te

CuO-ZnO-based catalysts is to increase both surface area and copper dispersion. This trivalent
p

ion has also an inhibiting effect on the sintering of Cu particles at on-stream conditions [2]. In
ce

contrast to ZnO, Hadipour and Sohrabi expressed that the presence of CuO and Al2O3 in excess

amount enhanced the catalyst activity by increasing the dispersion of active sites and hence
Ac

promoted the surface area of catalyst [96].

Copper particle size and its dispersion are found to be affected by preparation conditions

such as Cu/Zn molar ratio, type of precipitant and calcination temperature [110, 112, 113, 114].

In a series of experiments conducted by Wang et al., the effect of different Cu/Zn molar ratios of

Cu-Zn-based catalyst was investigated. The observation showed that low Cu/Zn ratio was more

beneficial to the WGS reaction, because more WGS active centers might be represented at those

21
Page 21 of 93
ratios. However, catalysts with higher Cu/Zn ratio exhibited greater activity for methanol

synthesis which might be ascribed to the stronger interaction between CuO and ZnO molecules

and the atomic dispersion between them [24]. Another important factor for CZA-based catalysts

in STD process is the ratio between CZA and the solid-acid function which was studied by Abu-

t
ip
Dahrieh et al. [25, 115]. Under the conditions they used, it was observed that the most suitable

ratio between the metal and acid function was 1:1 for CZA/γ-Al2O3 and 3:1 for CZA/HZSM-5

cr
admixed catalyst. In another work, a special core-shell structured bi-functional catalyst (an H-

us
ZSM-5 zeolite core enwrapped by one layer of CuO-ZnO shell) was prepared for the STD

process. Characterizations disclosed that these bi-functional catalysts possessed high Cu surface

an
area as well as high Cu dispersion, and consequently they would display excellent catalytic

performance [116].
M
As mentioned before, γ-Al2O3 is a methanol dehydration catalyst. It is very attractive since it
d

is cost effective and exhibits high surface area, excellent thermal and mechanical stability, high
te

mechanical resistance, and high selectivity towards DME [117, 118]. Furthermore, it has high
p

catalytic activity towards DME formation due to its low content of highly acidic sites which are
ce

mostly of the Lewis type [101]. Although γ-Al2O3 is active, it tends to strongly adsorb water

thereby losing activity [119].

Ac

Zeolites are crystalline aluminosilicates with periodic arrangement of cages and channels

which were found to have extensive industrial use as catalyst, adsorbent, and ion exchanger. It

can be concluded from literature that zeolite materials in a temperature range of 250-400°C and

pressures up to 18 bar can be a proper candidate to play the role of a solid-acid catalyst in

methanol dehydration process [119]. In comparison to other catalysts, zeolites in general possess

high surface area which comes from their microporous crystalline interface [10]. However, the

22
Page 22 of 93
zeolites' narrow and slender microporous structure may restrain DME from quickly diffusing

through the pores. As a result, zeolites may lose their catalytic activity and selectivity quickly

owing to the formation of by-products and deposition of carbonaceous compounds [30]. To

overcome this disadvantage, researchers have used some modifications to zeolite catalysts. For

t
ip
example, Tang et al. employed a ZSM-5/MCM-41 composite molecular sieve as the methanol

dehydration catalyst. The results exhibited high activity, selectivity and stability in the process of

cr
methanol dehydration to DME according to the combination of the channel advantage of the

us
mesoporous molecular sieve and the acidity advantage of ZSM-5 [30].

an
Among zeolite-type solid-acid catalysts used for the dehydration of methanol to DME, H-

ZSM-5 which exhibits more activity and stability than γ-Al2O3 catalyst [2, 119] is reported to be
M
the most promising for DME synthesis from syngas [120]. According to Qi et al. the activity and

selectivity of ZSM-5 can be increased by the use of an H-ZSM-5 supported Cu-Mo oxide
d

catalyst for direct synthesis of DME [121].

te

Another effective methanol dehydration catalyst is BFZ, with Beta zeolite cores and Y
p

zeolite polycrystalline shells. BFZ in the H-form (HBFZ) exhibits moderate acid strength and
ce

meso-porosity which is responsible for its high activity for CO hydrogenation [30]. In

comparison to CZA/HY bi-functional catalyst, CZA/HBFZ shows higher activity and stability
Ac

for the direct synthesis of DME from CO hydrogenation [30].

H-mordenite that can be a very attractive methanol processing catalyst is another zeolite that

is of interest owing to its high catalytic activity in etherification in conversion to olefins (MTO).

This property allows for the possibilities of performing both DME synthesis and MTO

technologies in a single reactor only by temperature adjustment [10]. Moradi et al. showed that

23
Page 23 of 93
acidic mordenite zeolites in H-form, produced from Na-form through ion-exchange process, had

higher surface area than the Na-form one. They explained that new pores were generated by ion-

exchange treatment, thereby increasing the surface area [10].

t
Stiefel et al. studied various dehydration catalysts in the synthesis of DME directly from

ip
carbon monoxide rich syngas. They showed that pore volumes and specific areas of zeolites

cr
would decrease in the following order:

us
H-MOR 90 > H-MFI 400 > H-MFI 90.

Moreover, their experiments introduced H-MOR 90 as a zeolite with the highest total

an
number of acidic sites among those studied and their investigation showed a decreasing order of

catalyst acidity as follows [117]:

M
H-MOR 90 > H-MFI 90 > γ-Al2O3 > H-MFI 400
d

The strongly increasing CO conversion in the case of H-MFI 90 is achieved at the expense
te

of a significant decrease in DME selectivity and an increase in the formation of hydrocarbons. In

p

addition, the CO2 concentration in the product mixture is also considerably increased. This can
ce

be traced back to an enhanced WGS reaction activity stimulated by a higher water concentration

in the reaction system leading to a pronounced production of CO2 and H2. In consequence of
Ac

increase in H2 concentration, the conversion of CO to methanol is favored [117].

Recently, polymeric heterogeneous catalysts, namely Nafion resin, have attracted a lot of

attention in the conversion of methanol to DME [122, 123]. Experiments were carried out in a

vapor phase flow reactor using Nafion resin beds [122] or Nafion/silica nanocomposites of

different compositions [123]. The Nafion catalysts provide 40% methanol conversion. In this

24
Page 24 of 93
case, no catalyst activity loss and coke formation were observed. Thus, Nafion is proved to be an

advantageous catalyst for the synthesis of DME from methanol [122, 123].

Aluminum phosphate (AlPO4) is also a promising catalyst in DME synthesis owing to its

t
lower amount of coke deposition, by-product formation, and its better water resistant property

ip
[124, 125]. The catalytic activity of AlPO4 in methanol dehydration is found to be dependent

cr
upon the preparation method, chemical composition (Al/P molar ratio), and activation

temperature [124, 126].

us
In recent years, multi-walled carbon nanotubes (MWCNTs) as a novel nano-carbon support

an
or promoter have drawn lots of attention [127, 128, 129]. MWCNTs possess several unique

features such as graphitized tube-wall, nanometer-size channel, high thermal conductivity and
M
excellent surface area. MWCNTs can be used as a catalyst support where metal particles with

catalytic activity may decorate along the external walls or be filled in the interior of the
d

MWCNTs. A type of bi-functional hybrid catalyst of Pd-decorated CNT-promoted Cu-ZrO

te

admixed with H-ZSM-5 zeolite has been developed. Its application to direct synthesis of DME
p

from CO2/H2 has been studied. The catalyst displayed excellent performance for the direct DME
ce

synthesis from CO2/H2 in heterogeneous “one-pot” reactions [130]. From literature, the

MWCNT-supported CZA/H-ZSM-5 catalyst is another appropriate option exhibited higher

Ac

catalytic activity and higher DME yield than unsupported one [99].

For the synthesis of DME by CO2 hydrogenation, the methanol synthesis component of the

bi-functional catalysts is usually CuO/ZrO2 catalyst, besides the traditional CZA one. Although

CuO/ZrO2 catalyst has been reported to be an effective catalyst, there are still disadvantages of

using zirconia support. In addition to the low specific surface area provided by CuO/ZrO2

25
Page 25 of 93
catalyst, the performance of the catalyst is under the influence of phase transformation. Zirconia

has three phases including m-ZrO2, t-ZrO2 and c-ZrO2 in which the transition between them

might change the properties of the catalyst. For a fixed Cu surface area, CuO/m-ZrO2 is more

active for methanol synthesis than CuO/t-ZrO2 [33]. Furthermore, compared with single oxide,

t
ip
mixed oxides have higher surface area, better thermal stability, mechanical strength, and stronger

surface acidity. Therefore, not only do the mixed oxides of titania and zirconia have the

cr
specialties of both oxides, but also they improve their disadvantages [77].

us
Recently, several studies have focused on the design of newer and more complex catalyst to

an
overcome the catalyst deactivation [131, 132]. However, this induces the addition of several

preparation steps resulted in extra costs and more waste production during the preparation of an
M
effective catalyst. Besides, it is difficult to limit the dehydration of methanol to the sole

formation of DME. Indeed, the presence of an acid catalyst leads to the consecutive formation of
d

hydrocarbons (methanol to hydrocarbons, MTH) [133, 134, 135]. More precisely, light olefins
te

(methanol to olefins, MTO) or alkanes (methanol to gasoline, MTG) can be obtained as a

p

function of temperature and pressure [136, 137, 138, 139, 140, 141, 142, 143, 144,].
ce

Nowadays, the MTO reaction is considered to be a valuable option for the improvement of

stranded gas reserves. Therefore, several studies are devoted to either the reaction mechanism or
Ac

the applied technology [133, 139, 140, 141, 142, 143, 144]. Song et al. showed that a pool of

adsorbed polymethyl benzenes played a leading role in catalytic cycles of MTH process [141,

142, 143, 144, 145]. Moreover, the methylation of hexa-methyl-benzene to form hepta methyl

benzeniumation has been shown to be the first step in the carbon pool mechanism [145, 146].

This is highly dependent on the acidity of the zeolite [146, 147]. In order to limit/inhibit the

MTH reaction, Ivanova et al. investigated SiC-supported ZSM-5 zeolite catalysts. The authors

26
Page 26 of 93
reported that the foam support allowed for the formation of small zeolite crystals which favored

the diffusion of produced DME throughout the porous network and thus seriously prohibited the

formation of consecutive hydrocarbons. The other possible way to artificially deactivate the

“hydrocarbon pool” is to carry out catalytic tests under the air atmosphere [132, 148]. Indeed, Fu

t
ip
and co-workers demonstrated that over SAPO-34 and ZSM-5 zeolites the presence of 20% of air

in the feed led to a negative impact on the production of olefins at 350°C [149]. As seen, the

cr
preparation of the catalyst remains quite complex and costly.

us
4.2. Catalyst preparation

an
Researchers are trying to modify the catalyst structure and/or formulation in order to

optimize the DME production as well as catalyst stability improvement. The preparation method
M
of the bi-functional catalyst systems in direct DME synthesis has a significant effect on the

performance of the process [124, 126, 131, 150, 151, 152]. Hybrid catalysts used for
d

synthesizing DME directly from syngas are prepared in different ways including physical mixing
te

of methanol synthesis catalyst and solid-acid catalyst, co-precipitation (sol-gel), impregnation

p

,and combined co-precipitation-ultrasound [30, 153]. In the case of the admixed catalyst, each
ce

function is prepared separately and then the powders of both functions are mechanically blended

[154, 155, 156, 157, 158, 159, 160]. The activity of the catalysts prepared by physical mixing is
Ac

higher than the activity of ones prepared by co-precipitation and impregnation [30]. In a study

conducted by Hosseini et al. nanocrystalline γ-Al2O3 catalyst was prepared by sol-gel and

precipitation methods. The obtained results showed that the catalysts prepared by the sol-gel

method have higher activity than catalysts prepared by the precipitation method. Furthermore,

non-aqueous sol-gel method offered higher activity in comparison with aqueous sol-gel one. The

advantages of the sol-gel method include the ability of maintaining a high degree of purity, the

27
Page 27 of 93
possibility of preparing samples at low temperatures, and changing physical characteristics such

as pore size distribution and pore volume [101].

4.3. Surface acidity of methanol dehydration catalysts

t
ip
Direct conversion of carbon-monoxide-rich synthesis gas to DME essentially depends upon

the acidity of the dehydration component in the catalyst system. If acidity is too low, the amount

cr
of methanol formed cannot be dehydrated with sufficient efficiency. If the acidity of the catalyst

us
is too high, it also catalyzes further the conversion of DME to hydrocarbons [117]. Hence,

according to several studies conducted, DME formation is related to sites with weak and medium

an
acidity and those catalysts with strong acid sites may be preferable for coke deposition [100].
M
In order to attain optimal condition for DME production, strong acid sites must be diluted in

order to achieve a high stability against coke formation [131]. Accordingly, Yaripour et al.
d

modified γ-Al2O3 with silica to improve its surface acidity. Their results evidenced that by
te

modifying alumina with silica, the surface acidity of the aluminosilicate catalyst increased with

increasing silica loading and reached its maximum peak at silica loading of 6 wt%. Then the
p
ce

surface acidity gradually decreased and showed almost similar performance in comparison to

unmodified γ-Al2O3 at 15 wt% silica loading [161]. Hosseini and coworkers studied the effect of
Ac

crystal size on the acidity of nanocrystalline γ-Al2O3 catalyst for the synthesis of DME through

the dehydration of methanol. The results showed that samples with smaller crystallite size

possessed higher concentration of medium acidic sites and consequently higher catalytic activity

[101].

28
Page 28 of 93
 It is well known that the acidic sites on the surface of solid-acid catalysts are of either

Bronsted or Lewis acid type [161]. Methanol is supposed to be dehydrated over both Lewis acid–

base pair and Bronsted acid–Lewis base pair sites [39, 162].

t
Concerning the acidity of γ-Al2O3, according to several investigations, it is mainly due to

ip
Lewis-acid sites whilst the acidity of MFI zeolites and AlPO4 in many cases is dominated by

cr
Bronsted-acid sites [102, 105, 163]. Acidity of H-MOR zeolites can also be dominated by

Bronsted-acid centers; but H-MOR systems with an approximately equal number of Bronsted

us
and Lewis acid centers have been also described. Both types of acid sites exhibited high acid

an
strength [117]. High Bronsted acidity of mesoporous aluminosilicate namely alumina

impregnated SBA-15 has also facilitated the conversion of methanol to values close to
M
equilibrium with 100% dimethyl ether selectivity at temperatures over 300oC [1]. In fact, authors

have proposed that H-ZSM-5 which is currently the best available methanol dehydration catalyst
d

for the STD process has large number of moderate strength Bronsted acid sites which is
te

responsible for its perfect behavior [102, 105, 164]. However, some reports did not find a
p

relationship between MeOH dehydration rates and the number of Bronsted acid sites in HZSM-5
ce

[165]; hence, more research is needed to prove the validity of this hypothesis.

4.4. Catalyst deactivation

Ac

The catalyst systems described in the preceding sections are generally prone to become

deactivated by the sintering of active copper sites, to coke deposition due to the presence of

strong acid sites, to poisoning because of the contaminants that might be present in the syngas,

and thus to the blockage of acidic sites [100, 117].

29
Page 29 of 93
 For hydrocarbon reactions over zeolites, deactivation is mainly attributed to two main

mechanisms: The acid site coverage which deactivates the catalyst by coke adsorption; and pore

blockage which is the deposition of carbonaceous compounds in cavities or channel intersections

that make a pore inaccessible and consequently prohibits access to the active sites inside the

t
ip
pores for the reactants [10]. In addition, it is well known that coke formation on zeolites is a

shape-selective process. Under comparable conditions, large-pore zeolites are more susceptible

cr
to deactivation by coke deposition than medium-pore zeolites [164].

us
Although H-ZSM-5 is not sensitive to water [2, 33], it shows high activity for the

an
transformation of DME into hydrocarbon byproducts. These hydrocarbons can further evolve

into heavy structures (coke) and consequently can block the zeolite pores and cause its
M
deactivation. However, this deactivation is slow due to the high partial pressure of hydrogen that

attenuates the mechanism of coke formation [41]. This phenomenon can be controlled by
d

employing a suitable concentration of Na in the zeolite in order to moderate the number of

te

Bronsted sites and to reduce the acid strength of the H-ZSM-5 zeolite [41]. The addition of
p

silicalite shell to the ZSM-5 zeolite is also considered as an efficient method for improving the
ce

resistance towards carbon formation [9].

A comparison between H-form mordenite and Na-form mordenite catalysts carried out by
Ac

Moradi et al. showed that H-MOR catalysts possessed higher initial dehydration activity owing

to the strong acidic properties of hydrogenation. However, the catalytic activity decreased

rapidly with time on stream in the reaction that would be attributed to higher concentration of

strong acid sites leading to the formation of coking materials. Moreover, they reported that the

super cage of H-MOR was likely to provide enough space for complete coking which resulted in

entrance blockage of the super cage. In this case, since water had no opportunity to eliminate the

30
Page 30 of 93
carbon deposited on the active sites and to regenerate mordenite catalysts, the deactivation of H-

MOR by carbon formation was irreversible [10].

Raoof et al. performed the indirect process of DME synthesis in an adiabatic fixed bed

t
heterogeneous reactor by using acidic γ-alumina in order to investigate the effect of water on the

ip
deactivation of γ-alumina. The mixture of methanol-water feed showed a catalyst activity loss of

cr
about 12.5 times larger than that of the pure methanol feed [4]. Decreasing the catalyst activity in

methanol-water feed was reported to be related to the blockage of the active sites through

us
competitive adsorption of methanol on the catalyst surface [117].

an
For the system of CuO-ZnO-Al2O3/γ-Al2O3 bi-functional catalyst, results evidenced that the

deactivation level was much lower for “H2 + CO2” feeds than for “H2 + CO” feeds. This result
M
was explained by the fact that higher concentration of water in the reaction medium formed

through the reverse water gas shift reaction in the case of CO2 rich feed limited coke deposition
d

owing to competitive adsorption between water and coke precursors on the active sites [8].
te

As previously mentioned, two types of reactors mostly used in the production of DME are
p

slurry reactors and fixed bed ones with CZA/γ-Al2O3 as DME synthesis catalyst. These catalysts
ce

deactivate more quickly in the slurry reactor than in the fixed-bed reactor. The deactivation of
Ac

hybrid catalyst for DME synthesis is caused by the deactivation of Cu-based methanol synthesis

catalyst rather than methanol dehydration catalyst [166].Compared with fixed-bed reactor, it is

more difficult to remove H2O from the surface of Cu-based catalyst in the slurry reactor because

liquid paraffin pose an additional resistance [8]. Thus, the morphology of the catalyst might

change owing to the existence of water. A part of Cu changes into Cu2(OH)2CO3 due to higher

partial pressure of water in DME synthesis, consequently leads to a decrease in the number of

31
Page 31 of 93
active sites of the Cu-based catalyst. In addition, under DME synthesis condition in slurry

reactors, some ZnO converts into Zn5(OH)6(CO3)2 which weakens the synergistic effect between

Cu and ZnO. Metal loss of Zn and Al, caused by hydrothermal leaching, is also another aspect of

methanol catalyst deactivation in such systems [50].

t
ip
4.5. Comparison of different catalysts

cr
As described in the preceding sections, solid-acid catalysts are used in indirect method and

us
also as a component of bifunctional catalyst for methanol dehydration to DME. Some of them

that will be discussed comparatively are alumina, zeolites, Nafion resin, AlPO4 and CuO/ZrO2,

an
from the viewpoints of activity, yield or selectivity toward DME production, and their possible

deactivation.
M
4.5.1. Activity
d

It has been shown by Flores et al. [167] that metallic components influence the direct
te

synthesis of DME from syngas. The role of this component was related directly to the CO
p

conversion. The different precipitation conditions used to prepare the methanol synthesis catalyst
ce

influenced its textural and structural properties, but these changes did not influence the catalytic

activity. For the methanol synthesis, amongst the dozens of catalytic materials proposed, the
Ac

largest utilization has the classical methanol synthesis catalyst Cu-ZnO-Al2O3, sometimes

modified with ingredients contributing to the increase of the copper dispersion and stability. It is

commonly employed in the one-step DME synthesis and usually prepared by the conventional

co-precipitation method, the catalytic activity depending on Cu/Zn/Al ratio and the preparation

conditions [168]. Catalysts based in Zr as the promoter also presented changes on the textural

and structural properties and exhibited an increase in the metallic area and CO conversion. The

32
Page 32 of 93
activity of a solid acid on the methanol dehydration reaction was also found to be determined

mainly by the number of its more acidic sites. Good activity and selectivity for methanol

etherification have the solid acids with moderate acidity (γ-Al2O3, zeolites, mesoporous materials

etc.). A largely used etherification catalyst is γ-alumina. Due to its relatively low content of high

t
ip
acidity sites, this catalyst offers a good selectivity towards DME, while exhibiting reasonably

high activity and high chemical and thermal stability. Activity and stability performances of γ-

cr
alumina can also be improved by promoting with different metal oxides. The Nb2O5 modified γ-

us
alumina showed a higher catalytic activity in methanol etherification than the untreated one

[168]. The mixture containing methanol catalyst and HZSM-5 has also been found to be one of

an
the most effective amongst the systems evaluated so far. However, over CZA /HZSM-5 mixtures

the reaction is controlled by the methanol synthesis step, thus changing the HZSM-5 amount
M
cannot affect the reaction data [154]. The following notes are on activity of different catalyst
d

discussed earlier in a short review:

te

 CZA: Negative effect of excess ZnO on the activity of the catalyst [96]. A
p

comparison of two bifunctional catalysts with the same CZA component but
ce

different solid-acid catalyst from their activity and stability viewpoint: CZA/HBFZ >

CZA/HY. MWCNT supported CZA/HZSM-5 has high activity [30].

Ac

 Zeolites: ZSM-5/MCM-41 has higher activity than zeolite catalyst. HZSM-5 > γ-

alumina from their activity and stability viewpoint. Evaluation of the catalyst activity

with respect to CO conversion [117]:

33
Page 33 of 93
 H-MOR 90 < H-MFI 90 < H-MFI 400 < γ-Al2O3 for T < 240°C

H-MOR 90 < H-MFI 400 < γ-Al2O3 < H-MFI 90 for T > 240°C

 AlPO4: The catalytic activity of AlPO4 in methanol dehydration is found to be

t
ip
dependent upon the preparation method, chemical composition (Al/P molar ratio)

and activation temperature [124, 126].

cr
 CuO/ZrO2: For a fixed Cu surface area, CuO/m-ZrO2 is more active for methanol

us
synthesis than CuO/t-ZrO2 [33].

4.5.2. Yield and selectivity

an
M
By changing the hybrid catalyst ratio, the DME/MeOH ratio in the product mixture can be

controlled. If the same amount of methanol catalyst is used, reaction systems with higher loading
d

of methanol dehydration catalyst would lead to higher DME yield at the expense of lower
te

methanol yield [169].

p

For Cu-ZnO-Al2O3/γ-Al2O3 bi-functional catalyst, maximum values of DME selectivity

ce

(83.4%) were obtained for a molar ratio of H2/CO=6/1, at 275ºC, 40 bar and a space time of

33.33 (g of catalyst) h/(mol of reactants). Under these reaction conditions, the use of NaHZSM-5
Ac

zeolite as acid function allowed a DME selectivity of 77.6% with a lower H2/CO molar ratio

(2/1) [168].

The best results for direct conversion of synthesis gas to DME were obtained over bi-

functional catalysts including HZSM-5 and HSY as methanol dehydration components, prepared

34
Page 34 of 93
by coprecipitation method: 99% DME selectivity in organic products were reported at 290°C, 40

bar, space velocity = 1500 h-1, using syngas with molar H2/CO = 2 and 5% CO2 [168].

DME selectivity is lower over the less active mixtures while an equivalent production is

t
achieved over both HZSM-5 and sulfated-zirconia, which further confirms that methanol is not

ip
efficiently dehydrated over weak acidic solids. DME production may be effectively achieved by

cr
adding an optimized amount of a solid-acid catalyst [154]. The following notes are on yield and

selectivity to DME for different catalyst discussed earlier:

us
 CZA: MWCNT supported CZA/HZSM-5 has high DME yield [30].

an
 Zeolites: ZSM-5/MCM-41 has higher selectivity than zeolite catalyst [30]. The
M
selectivity of H-MFI 90 decreases when the temperature is above 240°C [117].

 Alumina: alumina impregnated SBA-15 has 100% selectivity towards DME at

d

temperature above 300°C [1].

te

4.5.3 Deactivation
p
ce

An important challenge in the formulation of the bi-functional catalyst and the reactor

design is the prevention or limitation of deactivating phenomena: copper sintering, coking of

Ac

acidic components and metal ions migration. Particularly, a good temperature control is

necessary, due to the important overall process exothermicity. Results show that the deactivation

behavior of the catalyst is mainly caused by the deactivation of the methanol synthesis catalyst,

which is deeply caused by the synergistic effect [170]. The following notes are on different DME

catalyst deactivation:

35
Page 35 of 93
  CZA: Experimental data obtained using CuOZnO-Al2O3/γ-Al2O3 catalyst show that

there is no significant sintering below 325°C [168]. Its deactivation is more quickly

in the slurry reactor than in the fixed bed. However, for CZA/γ-Al2O3 deactivation is

lower when the feed is H2+CO, rather than H2+CO2 [8].

t
ip
 Zeolites: HZSM-5 is not sensitive to water, but has high activity for transforming

cr
DME to HC [41]. However, it can be regenerated by the addition of water resulting

from removing carbon deposited on the catalysts [169]. Sic supported ZSM-5

us
prohibits the formation of HC [132, 148]. H-MFI 90 increases its HC formation

an
when the temperature is above 240°C [117]. The deactivation of H-MOR is

irreversible [10].
M
 Nafion resin: No coke formation [122, 123].
d

 AlPO4: Leave coke deposition [124, 125].

te

5. Essential factors affecting the performance of DME production

p
ce

5.1. Water removal

An important parameter in DME synthesis through one-step method is the CO/CO2 feed
Ac

composition ratio. A strong synergy is obtained with CO-rich feed owing to the effective

removal of methanol by the dehydration and elimination of produced water by means of the

water gas shift reaction [22]. Conversely, CO2-rich feeds favor high fractions of unconverted

methanol due to the large quantity of H2O produced in methanol synthesis and dehydration steps,

thus inhibiting methanol dehydration and lowering DME selectivity. Hence, it is anticipated that

H2O in situ removal during DME synthesis may bring some beneficial effects. Considering

36
Page 36 of 93
indirect DME production from methanol, water is the side product in the dehydration reaction.

So, the presence of excess water clearly shifts the equilibrium backward and reduces the initial

methanol dehydration activity of DME as well as selectivity. Under the equilibrium condition,

water competes with methanol for the same sites on the catalyst surface and consequently a

t
ip
higher reaction temperature is required in order to achieve the same level of conversion [106,

164].

cr
At high CO2 content, in situ H2O removal accelerates the reverse water gas shift reaction

us
towards CO formation [171] and it is expected to improve DME production [172, 173]. In the

an
case of H2 -rich synthesis gas, in situ H2O removal would favor DME selectivity [44]. The

sorption-enhanced reaction process may offer an attractive possibility of in-situ H2O removal by
M
adsorption as shown by Carvill et al. [174]. Moreover, Iliuta et al. studied the sorption-enhanced

reaction process under in-situ H2O removal conditions for DME synthesis process in a fixed bed
d

reactor in order to analyze the effect of different parameters. By applying the adsorption-
te

enhanced concept, CO2 could be utilized as a constituent in the synthesis gas as in-situ H2O
p

removal that accelerates the reverse water gas shift reaction. The reason for in-situ H2O removal
ce

in this process displaced the water gas shift equilibrium to enhance the conversion of CO2 to

methanol and to improve the reactor productivity. The simulated results indicated that under H2O
Ac

removal conditions, DME yield and selectivity were favored and the fraction of unconverted

methanol was reduced. The role of H2O removal was prominent at higher CO2 feed

concentration, because a relatively large amount of water was produced. The preliminary

theoretical results indicated that the fixed bed reactor with in-situ H2O removal by adsorption

was more efficient in DME synthesis process than a fixed bed reactor without H2O removal [34].

37
Page 37 of 93
A possible disadvantage of H2O removal process is the deactivation of the catalyst metallic

function (CuO-ZnO-Al2O3) by coke deposition [175].

5.2. H2/CO ratio and CO2 content of the feed

t
ip
Syngas can be produced with different compositions through steam reforming (SR), carbon

dioxide reforming (CDR) and catalytic partial oxidation (CPO). The desired H2/CO ratio

cr
depends upon the intended use for the syngas. The SR process produces syngas too rich in

us
hydrogen while the syngas from CDR is too lean. The produced syngas from CPO is close to the

desired output ratio for DME production. In principle, the H2/CO blend can be adjusted by using

the water gas shift reaction [15].

an
M
Variation of H2/CO ratio can change the direction of water gas shift reaction. In low H2/CO

ratio, the reaction progresses to produce CO2 that results in enhancement of both methanol and
d

DME production. In high H2/CO ratio, CO2 production decreases which subsequently results in
te

DME reduction. Consequently, there is an optimum value for H2/CO ratio. An increase in

temperature leads to a decrease in optimum H2/CO ratio. This can be attributed to water gas shift
p
ce

reaction. With increasing temperature, water gas shift reaction reaches equilibrium conditions

rapidly. This, further results in the reduction of CO2 production [89]. In a recent work, it was
Ac

reported that CO2 removal prior to DME reactor greatly enhances the yield. Separation at this

stage would also provide high-purity CO2 and, therefore, would be beneficial for sequestration

[176]. But from another point of view, CO2 content of the feed must be under control. CO2 takes

part in methanol synthesis and is produced by WGS reaction. Therefore, CO2 is the bridge

relating methanol synthesis and WGS reaction [8]. CO2 molecules adsorb on the methanol

synthesis catalyst and occupy its active sites quicker than CO and H2, results in reduced

38
Page 38 of 93
methanol production [164]. Accordingly, increasing the concentration of CO2 in the feed stream

would be unfavorable to both reactions [24] and would lead to a decrease in CO and H2

conversion as well as in DME selectivity [2, 81, 117].

t
Since the H2/CO ratio can affect both synthesis gas conversion and product selectivity in

ip
direct synthesis of DME from syngas, it is worthy to find the optimum value H2/CO ratio. From

cr
the thermodynamic study, the optimum synthesis gas conversion can be obtained at the H2/CO

ratio of 1.0 [177]. As depicted in Fig.16 and Fig.17, selectivity towards DME decreases slightly

us
with an increase in H2/CO ratio while that of methanol has an increasing trend.

an
M Fig.16

Fig.17

5.3. Operational temperature

d
te

The temperature profile variations in reversible exothermic reactions like direct DME

synthesis have prominent effect on the reaction progress. At the beginning of the reaction, the
p

reaction is under kinetic control at low temperature and, consequently, DME production is
ce

enhanced by increasing temperature [8]. As reaction proceeds, the increased temperature causes
Ac

a reduction in equilibrium conversion of the reaction. Therefore, in reactors involving

exothermic reversible reactions, the temperature profile should decline as the reactions proceed.

Hence, applying a high temperature profile at the beginning of the one-step DME synthesis for a

higher reaction rate and then reducing the temperature gradually for increasing the equilibrium

conversion are appropriate methods for more DME production [89].

39
Page 39 of 93
 In the case of direct DME synthesis by an isothermal fixed bed reactor over a CZA-based

catalyst, at low temperatures the CO conversion is low owing to competitive adsorption between

CO and CO2 on the metallic function of the catalyst [41]. By increasing temperature, the CO

conversion decreases owing, firstly, to the thermodynamic restrictions of the exothermal reaction

t
ip
and, secondly, to Cu sintering which provokes partial loss of catalyst activity [2, 178]. Erena et

al. showed that the highest concentration of oxygenates (methanol and DME) could be obtained

cr
in the 250-300°C range. Above 300°C, hydrocracking reactions were dominant leading to a

us
sharp decrease in the selectivity towards DME [41].

an
For the synthesis of DME directly from syngas in slurry reactors, the CO conversion shows

a different trend towards temperature change. The CO conversion and DME productivity are
M
likely to increase with temperature. This can be attributed to the positive effect of temperature on

syngas solubility in liquid paraffin which improves the volumetric mass transfer coefficient
d

while accelerating methanol synthesis and dehydration rates. Tan et al. [179] demonstrated that
te

by increasing the reaction temperature more methane and other hydrocarbons were produced.
p

They also reported that at higher temperature, the methanol dehydration proceeded at a relatively
ce

high rate; and, consequently, stronger synergetic effect on syngas conversion resulted. However,

the temperature rise has certain limitations owing to a sintering phenomenon occurs at high
Ac

temperatures [180].

Raoof and coworkers studied the effect of temperature on catalytic dehydration of methanol

to dimethyl ether in an adiabatic fixed bed reactor. Since the reaction was exothermic and the

reactor was adiabatic, the temperature of catalyst bed increased from the feed inlet temperature

to a maximum value. Moreover, the reactor operating temperature increased relatively linear

with an increase in the feed temperature. This study showed that the methanol conversion to

40
Page 40 of 93
DME was not substantial at feed temperatures below 230°C and increased to the limit of about

85% at 250°C [4]. Also, Rownaghi et al. indicated that although higher reaction temperature

resulted in increased methanol conversion, selectivity towards DME decreased with increasing

the reaction temperature from 270 to 320°C [9].

t
ip
5.4. Operational pressure

cr
It can be concluded from literature that pressure is likely to increase the conversion of CO

us
whereby methanol synthesis is the limiting step of the overall reaction [2, 176]. This can be

explained by mole-number reducing stoichiometry of the methanol synthesis. Since water gas

an
shift and methanol dehydration reactions have the same number of moles on both sides of the

reaction, increasing pressure has no effect on these reactions and methanol synthesis by the
M
hydrogenation process of both CO and CO2 would be the only controlling steps [8, 164].

Although increased pressure is associated with increased CO conversion and DME productivity,
d

reactions at high pressures are limited by high operating costs [164]. Furthermore, it is observed
te

that in the transformation of H2+CO2 into dimethyl ether, CZA/γ-Al2O3 bi-functional catalyst
p

undergoes a slight deactivation owing to coke deposition. The increase in the coke content with
ce

pressure is explained by the enhancement of condensation reactions that leads to coke formation

[8]. In slurry reactors increasing pressure has an additional effect: it would decrease the bubble
Ac

size, hence increases the volumetric mass transfer coefficient that leads to a reduction in mass

transfer resistance in the slurry phase [50].

5.5. Space velocity

Space velocity is a crucial factor which influences catalyst performance. In the direct

synthesis of DME in fixed bed reactors, the conversion of CO dramatically decreases with

41
Page 41 of 93
increasing space velocity [24, 41, 117]. A similar behavior is expected for CO conversion when

superficial gas velocity is increased in slurry reactors. Increasing superficial gas velocity

decreases both mass transfer coefficient and mean residence time. Since the influence of mean

residence time on CO conversion is greater than that of the mass transfer coefficient, at higher

t
ip
gas velocities, CO conversion is reduced owing to inadequate time for syngas to diffuse into the

slurry phase and reach the catalyst surface. In such systems, DME productivity is under the

cr
influence of two different behaviors. By increasing superficial velocity at higher catalyst

us
concentration, DME production increases according to the enhanced flow rate of the entering

syngas to the column. In contrast, at low catalyst concentration, the DME productivity decreases

an
at higher velocity because of the decrease in CO conversion. Taking both of these effects into

account, an optimum superficial gas velocity should be found [164].

M
Considering the effect of space velocity on DME selectivity, there is still no general relation
d

between these two factors. Wang et al. observed that by increasing the space velocity, the
te

DME/CO2 ratio decreased considerably. This means that selectivity towards CO2 increases
p

significantly at the expense of DME selectivity [24]. It is observed by Erena and coworkers that
ce

both selectivity and yield of DME increased sharply at low values of space time and then they

increased monotonically to constant values. These results are in accordance with the fact that low
Ac

space time values favor the water shift reaction while high values of space time favor the

methanol dehydration reaction [41]. In another work by Stiefel et al., it was shown that while

DME selectivity remained constant in the case of γ-Al2O3 and H-MFI 400 catalysts, it decreased

by employing H-MFI 90 and H-MOR 90. This can be explained by strongly acidic character of

H-MFI 90 and H-MOR 90 at higher residence times which favors the DME conversion to higher

hydrocarbons [117].

42
Page 42 of 93
6. Process intensification (PI)

Traditionally, high purity DME is synthesized by dehydration of methanol produced from

syngas in as conventional gas phase process that involves a catalytic fixed-bed reactor followed

t
by a direct sequence of two distillation columns. The main problem of this process is the high

ip
investments costs for several units (e.g. reactor, columns, heat exchangers) that require a large

cr
overall plant footprint, as well as the associated energy requirements [181].

us
The main objective of PI is to improve processes and products to obtain technologies more

safe and economic. Due to the intrinsic characteristics of distillation separations, efforts to

an
improve distillation technology are still in twofold: one is to reduce the energy consumption and

the other is to reduce the capital investment. This calls for process intensification principles to
M
achieve intensified distillation systems to save both energy and capital costs. Catalytic

distillation has become in few decades very popular as demonstrated by the increasing
d

application of this technology to new and old production processes. The attractiveness of this
te

intensified process is based on the demonstrated potential for capital productivity improvements,
p

selectivity enhancement, reduced energy and polluting solvent consumption. These advantages
ce

are greatest when the combination of reaction and separation implies a reciprocal synergetic

effect. Besides the industrial production advantages, the catalytic distillation synthesis of DME
Ac

has some features, e.g. at the selected operative conditions, no side reactions is expected and the

only side product is water [182].

A very innovative solution to overcome the drawback of energy intensive distillation is

using dividing-wall column (DWC) technology [183] that can save up to 30% in CapEx and up

to 40% in OpEx [184]. DWC technology is very versatile and it can be used also in extractive

43
Page 43 of 93
distillation [185], azeotropic separations or reactive distillation [186]. Reactive distillation and

dividing-wall column technology can be effectively used for improving existing and new DME

processes. For example, the conventional DME purification and methanol recovery distillation

sequence can be successfully converted into a single-step separation based on DWC. Compared

t
ip
to the conventional direct sequence of two distillation columns, the novel proposed DWC

alternative reduces the energy requirements by 28% and the equipment costs by 20% [181].

cr
Moreover, reactive distillation is a feasible process intensification alternative to produce DME

us
by methanol dehydration, using solid acid catalysts. The innovative reactive DWC process has

excellent performance with the elimination of one process step which implies the reduction in

an
capital costs and the intensification of the process contributes to a quantum leap towards the

theoretical maximum in terms of mass, energy, space and time efficiency [187]. Consequently,
M
the R-DWC process can be considered as a serious candidate for the DME production in new as
d

well as revamped industrial plants. Other challenges that are in accordance with DME process
te

intensification are:
p

 To develop micro-channel catalytic reactors for process intensification and

ce

downsizing.

 To improve catalyst stability, in particular in the presence of sulfur and other

Ac

impurities.

 Seeking different techniques for mesoporosity generation in zeolites, efficient

control of acidity, metal dispersion and characterization of the catalyst using a

modern set of techniques.

44
Page 44 of 93
7. Conclusions and future perspectives

Significant research has been conducted to explore a variety of methods for DME to be

produced efficiently. Although the commercially proven technology for DME production is the

t
dehydration of pure methanol, many researchers are working on the direct conversion of syngas

ip
to DME (STD) with dual heterogeneous catalysts over which both catalysts are used in one

cr
reactor (bi-functional catalyst) to perform methanol synthesis and dehydration reactions

us
simultaneously [188]. A summary of the studies on DME synthesis methods, applied catalysts

and operating conditions are presented in Table 3. From the table it can be concluded that most

an
of the studies are conducted in the temperature range of 200-300°C and pressures up to 70bar.

According to the table, it is obvious that researchers have shown strong interest in applying CZA
M
and it can be understood that γ-Al2O3 and ZSM-5 zeolite are the most commonly used methanol

dehydration catalysts.
d
te

Table 3

Another conclusion that can be derived from the table is the great interest of researchers in
p
ce

applying fixed bed reactors due to their simplicity. However, they require high investment costs

for several units as well as the associated energy requirements. According to the aim of process
Ac

intensification, development of dramatically improved process alternatives, as compared to

the present state-of-art is necessary to bring a substantially smaller, cleaner, and more

energy-efficient technology such as catalytic distillation (CD), dividing-wall column (DWC) and

reactive dividing-wall column (R-DWC). However, the industrial companies urgently require a

comprehensive methodology enabling to proceed from the design phase to the working process.

45
Page 45 of 93
 Proper formulation of bi-functional catalysts related to their strength of acidic sites is also

important to have better process performance. However, the deactivation of catalysts as an area

of discussion has not been neglected; acid site coverage and pore blockage mechanisms are

believed to deactivate the zeolites, while water has influence on deactivation of γ-Al2O3. The

t
ip
comparison of catalyst deactivation in fixed-bed and slurry-phase reactors also exhibited stronger

deactivation of Cu-based catalysts in slurry reactors.

cr
Essential factors affecting DME production performance were also investigated. It has been

us
proved that the presence of water has an inhibiting effect on the reaction rate by competing with

an
methanol molecules over acid sites. Under H2O removal conditions with hydrophilic membranes,

the water-gas shift equilibrium displaced so that the conversion of CO2 into methanol increases
M
thereby enhancing DME yield and selectivity.

It is shown that there is an optimum value of H2/CO ratio in DME synthesis affected by
d

temperature. It is also proved that the CO2 content of the feed might decrease CO conversion as
te

well as DME selectivity. Effect of temperature on CO conversion is case dependent. It would

p

decrease CO conversion in direct DME synthesis through isothermal fixed bed reactors while it
ce

is likely to increase CO conversion and DME productivity in slurry reactors. In addition, higher

methanol conversion is achieved by the catalytic dehydration of methanol to dimethyl ether in an

Ac

adiabatic fixed bed reactor, at the expense of DME selectivity. From the stoichiometry of the

reaction it is concluded that pressure increases the conversion of CO and it is specified that the

methanol synthesis is the limiting step of the overall reaction. The last factor studied was space

velocity which had a decreasing effect on CO conversion in both fixed-bed and slurry reactors.

46
Page 46 of 93
 Despite numerous studies on DME production in literature, there is still lack of an entire

research area encompassing economic aspects of the process. Moreover, improving optimization

and scaling up of the one-step synthesis of DME should be considered. Relatively few researches

have been performed on the leading role of CO2 in DME synthesis, especially at the conditions

t
ip
of high space velocity. Moreover, the most favorable CO/CO2 ratio in the feed stream has not

been widely defined yet. More research is needed for the development of novel catalysts which

cr
show better performance in terms of activity, selectivity and most importantly stability towards

us
water. Investigating the optimum ratio of catalyst components to provide DME/MeOH mixtures

as required for different product requirements is also noteworthy. Verifying whether Lewis sites

an
can be converted to Bronsted sites in the presence of water is also suggested. It is also offered to

determine the optimum value of water removal in the case of membrane reactors. Developing a
M
long life-time catalyst for CH3Br hydrolysis reaction in order to make this approach practical and
d

investigating the efficiency of immobilizing catalysts in micro reactors are further proposed.
te

Abbreviations:
p

AF-SPMR Axial-flow Spherical Packed-bed Membrane Reactor

ATR Autothermal Reforming
ce

CapEx Capital Expenditure

CD Catalytic Distillation
CDR Carbon Dioxide Reforming
CFC Chlorofluorocarbon
Ac

CPO Catalytic Partial Oxidation

CZA CuO-ZnO-Al2O3
DME Dimethyl Ether
DWC Dividing-wall Column
GHSV Gas Hourly Space Velocity
HC Hydrocarbon
H-SOD Hydroxy-sodalite
LHSV Liquid Hourly Space Velocity
LPG Liquefied Petroleum Gas
MTG Methanol to Gasoline
MTH Methanol to Hydrocarbons
MTO Methanol to Olefins

47
Page 47 of 93
MWCNT Multi-walled Carbon Nanotube
OpEx Operating expense
PEFC Polymer Electrolyte Fuel Cell
POX Partial Oxidation
RD Reactive Distillation
R-DWC Reactive Dividing-wall Column
SR Steam Reforming

t
STD Syngas to DME

ip
WGS Water Gas Shift

cr
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us
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an
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p te
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67
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Figure Captions:

Fig.1. The trend of related publications on DME

t
Fig.2. Percentage of publications on (a) the subject of DME and (b) different synthesis methods

ip
Fig.3. Dimethyl ether production diagram

cr
Fig.4. A Scheme of Indirect Synthesis Process

Fig.5. A Scheme of Direct Synthesis Process

us
Fig.6. A schematic diagram of indirect DME production from natural gas in adiabatic fixed-bed reactor

(modified from [62]).

an
Fig.7. A schematic diagram of dual-type DME reactor configuration (modified from [62]).

Fig.8. The thermally coupled reactor configuration (modified from [65]).

M
Fig.9. A schematic diagram of the co-current mode for a recuperative reactor configuration (modified

from [65]).
d

Fig.10. Simplified DME production process via CD process (modified from [76]).
te

Fig.11. Path from conventional setup to reactive dividing-wall column (R-DWC) (modified from [76]).
p

Fig.12. A schematic diagram of co-current mode for fluidized bed membrane reactor (FBMR)
ce

configuration modified from [89]).

Fig.13. Catalytic F-4SF/ceramic composite tubular membrane (modified from [82]).

Ac

Fig.14. A scheme illustrating a catalytic membrane (modified from [82]).

Fig.15. Schematic diagram of two-bed reactor system (modified from [95]).

Fig.16. Equilibrium conversion of synthesis gas at 280oC and 50 atm (data from [177]).

Fig.17. Conversion and selectivity as a function of H2/CO ratio at 260oC and 50 atm (data from [177]).

85
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 Table 1

Properties

Molecular Formula C2H6O

Molar Mass 46.07 g mol−1

Appearance Colorless Gas

t
Odor Typical

ip
Density 1.97 g cm-3

Melting Point -141 °C, 132 K, -222 °F

cr
Boiling Point -24 °C, 249 K, -11 °F

Solubility in Water 71 g dm-3 (at 20 °C)

us
log P 0.022

Vapor Pressure >100 kPa

an
M
d
p te
ce
Ac

86
Page 69 of 93
 Table 2

Reactor Type Characteristics/Usages Benefits in a DME plant Cautions

Fixed-beds Simplicity and lower cost - Catalyst deactivation
Catalytic heterogeneous gas High recycle of

t
phase reactions syngas

ip
For catalytic reactions with low High operational
or intermediate heat of reaction investment
High conversion achieved by High pressure drop
decreasing the temperature

cr
along the reactor
Slurry Phase Catalytic heterogeneous gas phase Manageable temperature
Complicated
reactions better heat transfer equipment

us
Loss of catalyst
particles
Fluidized-bed Catalytic heterogeneous gas phase Lower gas-solid mass transfer Collision between
reactions resistance catalyst particles and
Excellent temperature control the reactor wall

an
High conversion and no need for Loss of catalyst
recirculation
Moderate operating pressure
Coupled and Dual For both highly exothermic and Lowering both capital and operating
M
Type Reactors endothermic reactions costs
Highly energy-efficient
Hot spots can be controlled
Coupling Reactor For methanol dehydration Higher selectivity/conversion CD: requires moderate
and Separation CD (or RD): distillation column Reducing operational cost temperature, while the
d

Units and the reactor are combined. R-DWC: lowers footprint with milder employed catalyst is
DWC: split the middle section operating condition, better active at higher
of a single tower into two performance (energy saving, reduced temperature
te

sections. CO2 emission, reduced total annual

R-DWC: reactive dividing-wall cost)
column (based on DWC design)
p

Micro Reactors For both highly exothermic and High controllability of the reaction laminar flow behavior
endothermic reactions conditions
Small holdup value
ce

Avoiding thermal runaway

Compactness and parallel
processibility
Membrane Has been used in indirect and also Good reaction yield May produce
Ac

Reactors direct methods. No additional steps of separation and undesired HC

purification Pore blockage
Prevent the catalyst deactivation Thermal/mechanical
Dual bed membrane reactor: stability issues
higher thermal efficiency
 reduces the cost of syngas
production
Spherical membrane reactor:
Decreases the pressure drop
Increases the DME production

87
Page 70 of 93
 Table 3
Synthesis Type of Temperature Pressure Ref.
Catalyst Other Essential Factors Authors
Method Reactor (oC) (bar) No.
Space velocity = 15000

t
mL(gcath)-1

ip
Fixed-bed
Direct Cu–ZnO–Al2O3/ZSM5 200-280 40 H2/CO/CO2/N2 = Chen et al. 2
Reactor
a. 61/30/5/4
b. 48/32/16/4

cr
H2/CO/CO2/N2/CH4 =
Micro Packed- Mixture of CuO–ZnO– 56/28/5/5/6 (mol%) Hayer et
Direct 220- 320 50–70 5
bed Reactor Al2O3 and γ-Al2O3 GHSV = 7500 al.

us
(Nml/gcat/min)
GHSV = 6000 L/(kgcat h)
Slurry Phase Mixture of the methanol
Direct 25-240 76.5 H2/CO = 0.45 Chen et al. 10
Reactor catalyst and γ - Al2O3
CO2 content = 4.8 (mol%)

an
Fixed-bed H2/CO = 1.82
Direct Cr/ZnO–S–Z 300, 325, 350 50 Yang et al. 17
Reactor CO2 content = 5.16 mol%
GHSV = 4500–60,000
Micro-channel Nml/gcat/h Hayer et
Direct CuO–ZnO–Al2O3/γ-Al2O3 210-300 10-50 23
Reactor al.
M
H2/CO = 1, 2, 4
Space velocity = 3000, 8000
Tubular Fixed- h-1 Wang et
Direct CuO-ZnO 100-250 20 24
bed al.
H2/CO = 1.8, 2.0, 2.1
d

Slurry-bed H2/CO = 0.6-1.5 Yuanyuan

Direct Bifunctional Catalyst 40-105 1.2-51.7 29
Reactor CO2 content = 0.0009-0.079 et al.
te

Space velocity = 1500 h–1

Fixed-bed CZA/HBFZ Wang et
Direct 250 50 H2/CO = 2.24 30
Reactor CZA/HY al.
CO2 content = 4.8%
p

H2/(CO+CO2) = 1, 1.5, 2
Fixed-bed
Direct Cu-ZnO- Al2O3/H-ZSM-5 250 50 Water removal (volume Iliuta et al. 34
Reactor
ce

fraction) = 0.5
H2/CO/ CO2 = 3/1.5/1
Isothermal
CuO-ZnO-Al2O3/NaH- Space time = 8.33, 16.66, Ereña et
Direct Fixed-bed 275 40 41
ZSM-5 33.33 and 66.66 (g of al.
Reactor
Ac

catalyst) h/mol of (H2 + CO)

Fixed-bed GHSV = 1000 h-1
Direct CuO–ZnO–Al2O3/γ-Al2O3 280 50 Lee et al. 43
Reactor H2/CO = 2
Fluidized Bed Cu–ZnO– Space velocity = 3000 ml/g/h
Direct 260 30 Lu et al 44
Reactor Al2O3/HZSM-5 H2/CO =1.0
Slurry Bubble H2/CO = 2.03 Papari et
Direct Cu-based C301/γ-Al2O3 230-270 50 52
Column CO2 content = 3 %mol al.
Pipe-Shell
Fixed-bed 62
(Thermally H2/CO = 2.52 Vakili et
Direct CuO-ZnO-Al2O3 /γ-Al2O3 220 50 and
Coupled Heat CO2 content = 4.09 mol% al.
Exchanger 64
Reactor)
Fixed-bed
Direct CuO-ZnO-Al2O3/H-ZSM-5 250 50 H2/(CO2+CO) = 1.0, 1.5 Iliuta et al. 84
Membrane

88
Page 71 of 93
 Reactor Water removal = 0.0-5.0e-10
Tube
Fluidized Bed Space velocity = 3000
side: 40 Mardanpo
Direct Membrane - 220-300 ml/gcat/hr 89
Shell ur et al.
Reactor H2/CO = 1
side: 50
Water removal = 1e−10-
Autothermal 5e−10 kmol/(sm2 Pa)
Direct Dual-bed Cu-ZnO-Al2O3/H-ZSM-5 250 32 Iliuta et al. 95
H2/CO = 100
Membrane

t
CO2 content = 21.5 %mol

ip
Fixed-bed Hadipour
Direct CuO–ZnO–Al2O3/γ-Al2O3 230-300 9 CO/CO2/H2 = 64/32/4 vol.% 96
Micro-reactor et al.
Cu/Y

cr
Fixed-bed Cu-Mn/Y Space velocity = 1500 h-1
Direct 245 20 Fei et al. 111
Reactor Cu-Zn/Y H2/CO = 1.5
Cu-Mn-Zn/Y

us
GHSV = 5500 L/(kgcat h)
Fixed-bed Cu–ZnO–Al2O3/Zr-
Direct 250 40 CO/CO2/H2 = 41/21/38 Bae et al. 112
Reactor ferrierite
(%mol)
CuO/ZnO system

an
Fixed-bed Residence time = 10-90 s Stiefel et
Direct γ-Al2O3 up to 450 up to 100 117
Reactor H2/CO = 0.67 , 1.0 al.
zeolites and γ-Al2O3
GHSV = 25,000 mL STP/(h
Fixed-bed g-hydr.catal.) Zhang et
Direct CuZr–PdCNTs/HZSM-5 190-270 50 130
Reactor al.
M
H2/CO2/N2 = 69/23/8
H2/CO = 2.0 Khoshbin
Fixed Micro-
Direct CuO-ZnO-Al2O3/H-ZSM-5 200-300 10-40 and 153
reactor GHSV = 600 cm3/grcat.hr Haghighi
Cu-ZnO-Al2O3/H-ZSM-5
d

Fixed-bed GHSV = 6000 ml/gcat h

Direct Cu-ZnO-Al2O3/ Na- 250-280 42 Kim et al. 155
Reactor H2/CO = 1.5
ZSM-5
te

HZSM-5 zeolites modified GHSV = 1500 ml/(h gcat)

Tubular
Direct with various contents of 260 40 Mao et al. 165
Reactor H2/CO/CO2 = 0.66/0.30/0.04
magnesium oxide
p

Slurry Phase H2:CO = 1:1.5 Wang et

Direct CuO–ZnO–Al2O3/γ-Al2O3 260 50 166
Reactor Space velocity = 4000 h-1 al.
ce

CZZr
Fixed-bed CZAZr Flores et
Direct 250 50 H2/CO = 2.0 167
Reactor Katalco mixed with H- al.
ferrierite zeolite
Ac

F51-8PPT GHSV = 5000-20000 h-1

Micro-channel
Direct ZSM-5 220-320 10-40 H2/CO = 2.0, 3.0 Hu et al. 173
Reactor
Acidic Al2O3 CO2 content = 4 mol%
Space time = 12.8 (g of
catalyst) h (mol of
Isothermal reactants)−1
H2/CO = 3/1 Sierra et
Direct Fixed-bed CuO-ZnO-Al2O3/γ-Al2O3 275 30 175
al.
Reactor Time on stream = 30 h
Water/syngas molar ratio in
the feed = 0-0.6
Bifunctional Catalyst
Isothermal GHSV = 800 ml/gcat h
(blend of methanol Kabir et
Direct Plug-flow 900 30-60 176
dehydration catalyst and γ- H2/CO = 0.81 al.
Reactor
Al2O3)

89
Page 72 of 93
 Slurry Phase GHSV = 2.0 L/g cat h,
Direct Mn/CuZnAl 240, 260, 280 50 Tan et al. 179
Reactor H2/CO = 2/1
Fixed-bed Ereña et
Direct CuO–ZnO–Al2O3/γ-Al2O3 225-325 20 H2/CO = 2 189
Reactor al.
Fixed-bed GHSV = 5500 L/(kgcat h)
Direct Mesoporous Cu–γ-Al2O3 285-325 50 Jiang et al. 190
Reactor CO/CO2/H2 = 41/21/38
Cu/Zn
Cu/Zn/Al

t
H2/CO = 0.5-2.0
Fixed-bed CZ-A (Copper Acetate,

ip
Direct 240–290 30-70 Space velocity = 3000-6000 Kim et al. 192
Reactor Zinc Acetate)
h-1
CZ-N (Copper Nitrate,
Zinc Nitrate)

cr
Fixed-bed GHSV = 2000 mLg−1cat h−1
Direct Al-MCM-41 260 50.7 Naik et al. 193
Reactor H2/CO2 = 3
Slurry Bubble

us
Direct - 250 52.7 H2/CO = 2.2 Chen et al. 194
Column
−1
Space velocity = 1000 h
Direct Fixed-bed Cu-Zn-Al/HZSM-5 240 50 H2/CO = 2/1 Jia et al. 196
CO2 content = 7.06 (%mol)

an
Tubular Fixed- GHSV = 8400 h-1
Mixtures of Cu/ZnO/Al2O3 Montesano
Direct Bed Micro- 250 50 197
and γ-Al2O3 H2/CO = 2/1 et al.
reactor
CuOZnO-Al2O3/ γ -Al2O3 Space time = 8.33 g catalyst
M
and CuO-ZnO- h/(mol reactants) Aguayo et
Direct Fixed-bed 275 40 198
Al2O3/NaHZSM-5 hybrid al.
catalysts H2/CO = 4/1

GHSV = 1700 mL García-

syngas/(gcat h) Trenco
Direct Fixed-bed CZA 260 40 199
d

and
H2/CO/CO2 = 0.66/0.30/0.04 Martínez
te

Mixture of CuO---ZnO--- GHSV = 1500 ml/(h gcat)

Direct Fixed-bed Al2O3 and sulfate-modified 260 40 Mao et al. 200
γ -Al2O3 H2/CO/CO2 = 0.66/0.30/0.04
Direct Isothermal
p

H2/CO = 3/2
and Fixed-bed Cu–Mn–Zn/Ce–HY 245 20 Jin et al. 100
Reactor Space velocity = 1500 h–1
Indirect
ce

Alumina Impregnated
Tubular SBA-15 (Al@SBA-15)
High Bronsted acidity of Tokay et
Indirect Packed Mesoporous 120-450 1 1
Al@SBA-15 al.
Reactor Aluminosilicate
Ac

γ-Al2O3
Fixed-bed Zhang et
Indirect Al2O3-HZSM-5 190-300 5 - 3
Reactor al.
Adiabatic Raoof et
Indirect γ-Al2O3 233-303 1 LHSV (h-1) = 8.5 4
Fixed-bed al.
Fixed-bed Rownaghi
Indirect ZSM-5 180-320 1.1 - 9
Reactor et al.
γ-Al2O3
Zeolites (HY, HZSM-5
Catalytic
and HM) Hosseinine
Indirect Distillation 110-135 9 - 20
Ion exchange resins jad et al.
Column
(Amberlyst 15, 35, 36
and 70)
Fixed-bed
Indirect γ-Al2O3 260 18.2 - Farsi et al. 60
Reactor

90
Page 73 of 93
 Heat
exchanger 65
Reactor Pt/Al2O3 Khademi
Indirect 150-250 1 - and
Adiabatic γ-Al2O3 et al.
Fixed-bed 90
Reactor
Platelet Milli-
Indirect H-ZSM5/SiC 250 1 - Liu et al. 77
reactor
Spherical

t
Packed-bed Samimi et

ip
Indirect γ-Al2O3 260 18.8 Water removal = 5 (%mol) 80
Membrane al.
Reactor
Catalytic
Volkov et

cr
Indirect Membrane F-4SF Resin 180 1 - 82
al.
Reactor
Fixed-bed Hosseini
Indirect Nanocrystalline γ-Al2O3 300 1 LHSV = 2.8, 11.7, 26.1 h−1 101
Reactor et al.

us
Fixed-bed Lertjiamra
Indirect AlPO4 150-300 1 - 102
Reactor tn et al.
γ- Al2O3
Fixed-bed Yaripour
Indirect Modified γ-Al2O3 with 300 1 GHSV = 15,600 h-1 161

an
Reactor et al.
silica
Fixed-bed
Indirect meso- γ-Al2O3 300 - Absence of an acid catalyst Khaleel 191
Reactor
Top stage =
M
Fixed-bed CuO–ZnO–Al2O3 268 Space velocity = 1800 h-1
Indirect 80 Zhu et al. 201
Reactor HZSM-5 Bottom H2/CO = 2.0
stage = 236
CO2 Isothermal
Ereña et
Hydrogena Fixed-bed CuO–ZnO–Al2O3/γ-Al2O3 225-325 20-40 - 8
d

al.
tion Reactor
CO2
te

Fixed-bed Blend of CuO–TiO2–ZrO2 Wang et

Hydrogena 200 30 - 33
Reactor and H-ZSM-5 al.
tion
CO2
Fixed-bed
p

Hydrogena CuO–ZnO–Al2O3/H-ZSM-5 262 30 - Zha et al. 99

Reactor
tion
CH3Br Glass Tube ZnCl2 loading from 1.0 mol%
ce

ZnCl2/SiO2 180 - You et al. 31

Hydrolysis Reactor to 12 mol%
CH3Br Prakash et
Batch Reactor PVP 50-125 - - 36
Hydrolysis al.
(*) f= (H2-CO2)/(CO+CO2)
Ac

91
Page 74 of 93
Table Captions:

Table 1. Properties of DME

Table 2. Different types of DME reactors in comparison

Table 3. Production conditions for various works on DME synthesis

t
ip
cr
us
an
M
d
p te
ce
Ac

92
Page 75 of 93
 Highlights:
 Considering the papers accomplished over the years 1965-2013.
 Focusing on production methods with discussion on their wide variety of reactors.
 Investigating on catalyst configurations, operational temperature and H2/CO ratio.

t
ip
cr
us
an
M
d
p te
ce
Ac

93
Page 76 of 93
Figure(s)

30
Number of Publications

25
25
21 22 21
20
16 16 17
15 14
13 13
11 11

t
10 8 9
7 6

ip
5 3
2 2
0

cr
us
Time (Year)

Fig.1
an
M
ed
pt
ce
Ac

Page 77 of 93
 58%
Application & Properties of
DME
Catalyst Technologies
30%
12%

t
Reactor Technologies

ip
cr
us
(a)

an
M
58%

Direct Synthesis
ed

Indirect Synthesis
33%
9% Other Routes
pt
ce

(b)
Ac

Fig.2

Page 78 of 93
 Coal Synthesis Methanol
Sources: Natural Gas Gas DME
Oil Indirect Conversion
Biomass (CO+H2)
Direct Conversion
Fig.3

t
ip
cr
us
an
M
ed
pt
ce
Ac

Page 79 of 93
Methanol

v v
v To DME
Tank
v v
v v
v v

v v

t
Reactor To Waste

ip
DME Tower Water
Treatment

cr
v
Methanol/Water Tower

us
Fig.4

an
M
ed
pt
ce
Ac

Page 80 of 93
 To Burner
To Gas Pipe
v
Syngas v
Water
To DME Tank
v
v v
v
v v To
v Methanol
v
v v Tank
v
v
v v
DME Tower
Reactor v To Heat

t
v Exchanger

ip
Methanol/Water Tower

cr
Fig.5

us
an
M
ed
pt
ce
Ac

Page 81 of 93
 Methanol
Natural Gas Methanol
Production
Unit

t
Product

ip
DME
Distillation Unit

cr
DME Production Unit

us
Fig.6

an
M
ed
pt
ce
Ac

Page 82 of 93
 Second Reactor First Reactor
Preheated Syngas
Water Vapor
Steam Drum

Bi-functional
Catalysts

t
ip
Syngas from Reacting Gas

cr
Reforming
Distillation Unit Pure DME

us
Fig.7
an
M
ed
pt
ce
Ac

Page 83 of 93
 Hydrogen

Separator
Methanol from Storage

Benzene

Cyclohexane

t
ip
Distillation
Pure DME

cr
Unit

us
Fig.8
an
M
ed
pt
ce
Ac

Page 84 of 93
 C6H12
CH3OH Ar

Exothermic Section

Endothermic Section Heat

Transfer

t
ip
cr
us
CH3OH C6H12
DME C6H6
H2 O H2
Ar
an
M
Fig.9
ed
pt
ce
Ac

Page 85 of 93
 DME Methanol

Distillation
Methanol CD Column Column

t
ip
Water

cr
Fig.10

us
an
M
ed
pt
ce
Ac

Page 86 of 93
 Reactive Distillation

t
ip
cr
Distillation Reactive DWC

us
an
Dividing Wall Column
M
Fig.11
ed
pt
ce
Ac

Page 87 of 93
 DME

Pd-Ag Membrane

H2

t
H2

ip
Tube Side

cr
(Reaction Side)

us
Shell Side
(Permeation Side)
an Synthesis Gas
(P=40bar)
M
Synthesis Gas Synthesis Gas
ed

(P=50bar) (P=50bar)

Fig.12
pt
ce
Ac

Page 88 of 93
 α -Al2O3
TiO2

F-4SF
Resin Layer

t
ip
Fig.13

cr
us
an
M
ed
pt
ce
Ac

Page 89 of 93
 Product (DME, water, and
unreacted methanol) and N2

O-ring
Shell Catalytic membrane
Thermocouple

t
ip
cr
us
Sweep Gas N2 Product (DME, water, and
Feed unreacted methanol) and N2
Methanol and N2

an
Fig.14
M
ed
pt
ce
Ac

Page 90 of 93
 Inert Gas

Syngas
Glycerol + H2O

Heat

t
Heat Glycerol Reforming

ip
Membrane DME

cr
Synthesis
H 2O

us
DME + by-products
an
M
ed

Fig.15
pt
ce
Ac

Page 91 of 93
 100

Equilibrium conversion of syngas (%) 80

60

40

t
ip
20

cr
0

us
0 0.5 1 1.5 2 2.5 3

H2/CO (molar ratio)

3CO+3H2=CH3OCH3+CO2 an
2CO+4H2=CH3OCH3+H2O CO+2H2=CH3OH

Fig.16
M
ed
pt
ce
Ac

Page 92 of 93
 100

Conversion of syngas (%), Selectivity (%) 80

60

40

t
ip
20

cr
0

us
0.25 0.5 0.75 1 1.25 1.5 1.75 2 2.25
H2/CO (molar ratio)

Syngas Conv. DME an Methanol CH4

Fig.17
M
ed
pt
ce
Ac

Page 93 of 93