Dimethyl Ether

Dimethyl Ether

10 DIMETHYL ETHER CHAPTER OUTLINE 10.1 Introduction 363 10.1.1 Dimethyl Ether as Sustainable Fuel 364 10.1.2 DME Production Routes 366 10.2 Physical P...

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10 DIMETHYL ETHER CHAPTER OUTLINE 10.1 Introduction 363 10.1.1 Dimethyl Ether as Sustainable Fuel 364 10.1.2 DME Production Routes 366 10.2 Physical Properties, Chemical Equilibrium, and Kinetics 10.2.1 Process Chemistry 368 10.2.2 Chemical Equilibrium 368 10.2.3 Kinetics of Methanol Dehydration 369 10.3 Conventional DME Process 371 10.3.1 Process Design and Optimization 371 10.3.2 Intensified Separations Using Dividing-Wall Column Technology 379 10.4 Novel Process Intensification Alternatives 384 10.5 Catalytic Distillation Process 385 10.6 Combined Gas-phase Reactor and Reactive Distillation Process 391 10.7 Conclusions 395 References 396

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10.1 Introduction Dimethyl ether (DME) has some remarkable properties (e.g., nontoxic colorless gas, noncorrosive, noncarcinogenic, environmentally friendly) that make it usable as green aerosol propellant, as precursor to organic compounds and especially as clean fuel for diesel engines or combustion cells (Muller and Hubsch, 2005; Arcoumanis et al., 2008). Other DME applications include refrigeration (blended with Freon), feedstock for dimethyl sulfate synthesis, production of acetic acid and olefins, liquid petroleum gas (LPG) substitute, household cooking fuel, and fuel for power generation. Similar to other second-generation biofuels, DME can be produced from a variety of feedstock, both fossil (natural gas and coal) and renewable (biomass): e.g., mainly

Applications in Design and Simulation of Sustainable Chemical Processes. https://doi.org/10.1016/B978-0-444-63876-2.00010-3 Copyright © 2019 Elsevier B.V. All rights reserved.

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coal in China; natural gas in Egypt, Middle East, and Indonesia; or lignocellulosic biomass (black liquor) in Sweden (Ptasinski, 2016). The global DME production was about 3 million ton/year, although the actual production capacity is much higher, about 10 million ton/year (Fleisch et al., 2012)dready to be unleashed for DME as fuel. The DME market is currently estimated to worth 9.7 billion USD by 2020 (www.marketsandmarkets.com).

10.1.1 Dimethyl Ether as Sustainable Fuel According to the International DME Association (www.aboutdme. org), DME is an ultra low-emission and sulfur-free fuel, with diesellike performance and propane-like handling. DME has a quiet and sootless (free of soot) combustion, high cetane rating (55e60 vs. 40e55 for diesel), low emissions, relatively high energy density (about 80% that of diesel), and low cost. However, other properties of DME are less favorable compared to conventional diesel fuels, e.g., liquid density of 735 kg/m3 (at 25 C and atmospheric pressure), lower heating value of 28.9 MJ/kg (equivalent to 66% of the energy content by mass or w50% by volume of diesel fuel), low viscosity (a factor 20 lower than diesel), and lack of lubricity. Nonetheless, because of its high cetane rating (over 50) and low boiling point (25.1 C, requiring pressurization as an LPG fuel), DME has several advantages as a fuel suitable for compressionignition engines: fast mixing of fuel and air, lower levels of ignition delay, good starting in cold weather (Fleisch et al., 2012; Ptasinski, 2016). A key point of attention is that using DME as a diesel fuel alternative means that DME must be liquefied by pressurization above its vapor pressure (at about 5.1 bar at 25 C). The air/fuel ratio of DME fuel is about 9 versus 14.6 for diesel, which means that the complete combustion of 1 kg DME requires less air than for diesel fuel. DME has a much higher and wider flammability range in air than gasoline, diesel, and propane, but it is very similar to natural gas. In liquid state, DME has low viscosity and low lubricityd crucial properties that strongly affect the achievable injection pressure in a fuel injection system. Viscosity allows it to pass through narrow passages, while the lack of lubricity can accelerate the wear of moving surfaces (e.g., feed pump, high-pressure injection pump, injector nozzles). Hence fuel additives are essential to improving DME viscosity and make it a viable fuel for diesel engines. In addition, DME adversely affects some plastics and rubbers and it dissolves most elastomers in a fuel system, but it is compatible with Teflon and Buna-N rubber. As the compressibility of DME is higher than that of diesel, the compression energy

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in the DME fuel pump is also greater. Overall, the differences regarding lubricity, viscosity, bulk modulus, and energy density mean that many components of fuel systems must be changed when converting from diesel to DME while the fuel injection timing and duration must be altered as well. Studies indicate that DME can be considered as one of the most efficient fuels (Ptasinski, 2016): • The well-to-tank (WTT) efficiency represents the energy efficiency of the fuel production from a specific feedstock (strong dependence on raw material). DME produced from natural gas has a WTT efficiency of about 70% (Semelsberger et al., 2006), which is somewhat lower than that of gasoline and diesel (80%e90%). Still, the DME production is the most efficient process as compared to all alternative fuels (e.g., methanol, hydrogen, or ethanol). • The tank-to-wheel (TTW) efficiency represents the efficiency of energy conversion in a vehicle, including transmission, engine, and vehicle operation. DME as a fuel in the compressionignition engine has the same energy efficiency as the traditional diesel fuel, about 28% (Semelsberger et al., 2006). However, higher TTW efficiencies are expected in new vehicles such as diesel hybrid electric vehicle (35%) and fuel celleequipped vehicles (41%e42%). • The well-to-wheel (WTW) efficiency represents the overall energy efficiency (WTT$TTW), including the production and vehicle efficiency. DME also takes a top position here, among the alternative fuels. The WTW efficiencies for DME produced from natural gas range from 18% for the compression-ignition engine up to 23% for fuel celleequipped vehicles. Depending on the production pathways, reductions in GHG emissions relative to diesel fuels are substantial, ranging 93%e101% in the United States and 87%e92% in the European Union (Lee at el., 2016). Compared to other transport fuels, DME is environmentally benign (Ptasinski, 2016) and has a very low global warming potential (GWP). The troposphere lifetime of DME is about 5.1 days that corresponds to a GWP of 1.2 (20-year time horizon), 0.3 (100-year horizon), and 0.1 (500-year horizon). For comparison only, the GWP values for methane are 56, 21, and 6.5, respectively (Semelsberger et al., 2006). In terms of emissions, the DME produced from natural gas is similar to diesel, whereas in case of DME produced from biomass, the overall emission level is only 17% as compared to diesel (Arcoumanis et al., 2008; Lee at el., 2016). Several companies (e.g., Unitel Technologies, www.uniteltech. com) consider the DME production from various biofeedstock

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such as high yield crops grown specifically for energy applications; food wastes from manufacturing, preparation, and postconsumer usage; agricultural residues from harvesting or processing; virgin wood from forestry and arboricultural sources. However, these routes also imply issues related to collection, transportation costs, and storage of raw materials.

10.1.2 DME Production Routes Currently, DME can be produced via two main routes (Azizi et al., 2014): • Indirect route by the dehydration of methanol (n.b. for methods for methanol manufacturing see Chapter 3). Overall, this is known as the indirect route or the two-step process: (1) methanol synthesis and (2) methanol dehydration. Toyo, MGC, Lurgi, and Uhde have their own indirect processes for DME production. • Direct route in a single stage using bifunctional catalysts. This one-step process (arguable more efficient) requires a dual-catalyst system that allows both the methanol synthesis and dehydration to take place in the same unit, with no methanol isolation and purification. Companies such as Haldor Topsoe, JFE Holdings, Korea Gas Corporation, Air products, and NKK produce DME via the direct route. Both types of DME production processes are commercially available, but the two-step process is more widespread industrially because it is relatively simple and the start-up costs are rather low. Quite a number of companies have been developed processes to convert coal- or natural gasebased syngas into DME including Haldor Topsoe, NKK Corporation, Air Products, Toyo Engineering Company. A detailed description of these DME processes is given in Encyclopedia of Chemical Processing (Sardesai, 2006).

10.2 Physical Properties, Chemical Equilibrium, and Kinetics Process simulation implies mainly the vapor-liquid equilibrium (VLE) for the mixture methanol/DME/water. The nonideality of the liquid phase can be modeled using the UNIQUAC or NRTL liquid activity models. Methanolewater and methanoleDME binary interaction parameters are available in the Aspen Plus database, while the watereDME binary interaction parameters

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Figure 10.1 Residue curve map of the mixture dimethyl ether (DME)emethanolewater (left). Txy diagrams of the (A) methanol/DME and (B) water/MeOH mixtures (right).

were taken from the NIST database. The nonideality of the vapor phase was modeled by the RedlicheKwong equation of state. RedlicheKwong EOS requires critical properties of the pure components, which are also available in the Aspen Plus database. The property model was also validated against experimental data reported in literature (Teodorescu and Rasmussen, 2001; Ihmels and Lemmon, 2007; Wu et al., 2011). Fig. 10.1 shows the residue curve map of the DMEemethanolewater mixture, calculated at 10 bar. There are no azeotropes present in this system, but there is a small region where phase splitting is possible. Fig. 10.1 also shows the Txy diagrams for MeOH/DME and Water/MeOH, calculated at 10 and 1 bar, respectively. If performed by distillation at 10 bar, the MeOH/DME separation appears easy, being possible to obtain DME of very high purity. The main required utilities are low-pressure steam (160 C) for reboiler and cooling water (30 C) for condenser. The separation of water/MeOH can be achieved also by distillation at atmospheric pressure, using the same utilities.

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10.2.1 Process Chemistry The traditional two-step DME production method comprises conversion of syngas to methanol as the first step, followed by methanol dehydration. The most important reactions taking place in the first step are the syngas conversion to methanol that is coupled with the water gas shift (WGS) reaction, typically using Cu-based catalysts (Ptasinski, 2016): CO þ 2H2 %CH3 OH

DH298K ¼ 90:6 kJ=mol

CO2 þ 3H2 %CH3 OH þ H2 O CO þ H2 O%CO2 þ H2

DH298K ¼ 49:4 kJ=mol

DH298K ¼ 41:2 kJ=mol

(10.1) (10.2) (10.3)

In the second step, methanol is dehydrated to DME according to the following chemical reaction: 2CH3 OH%CH3 OCH3 þ H2 O

DH298K ¼ 23:4 kJ=mol

(10.4)

All these reactions form a synergetic system and can be preferentially carried out in the same reactor, as applied in the one-step DME synthesis. The net reaction of the DME synthesis (note that only the reactions (10.1), (10.2), and (10.4) are stoichiometric-independent) is a combination of these three reactions (Ptasinski, 2016): 3CO þ 3H2 %CH3 OCH3 þ CO2

DH298K ¼ 246 kJ=mol

(10.5)

10.2.2 Chemical Equilibrium When the etherification reaction (10.4) is considered, the methanol equilibrium conversion is rather high, exceeding 80% in the temperature range 200e400 C. The reaction is slightly exothermic and proceeds at constant number of moles; therefore the equilibrium conversion slightly decreases with temperature and is independent of pressure. The main synergetic advantage of the reaction network (10.1)e(10.4) is that the equilibrium of methanol synthesis reaction is shifted toward methanol formation, as the methanol produced here is simultaneously consumed in the methanol dehydration reaction. Furthermore, the water formed in the dehydration reaction is consumed in the WGS reaction. The hydrogen generated from the WGS reaction also shifts the methanol synthesis reaction toward the methanol formation. Consequently, the CO equilibrium conversion for the DME synthesis reaction is much higher as compared to that of the methanol synthesis reaction.

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Figure 10.2 Comparison of the CO equilibrium conversion for dimethyl ether (DME) and methanol synthesis (left). Yield of DME production from CO (right).

Fig. 10.2 shows the equilibrium conversion of CO versus temperature, for different values of the pressure and starting from an equimolar H2/CO mixture. The calculations have been performed in Aspen Plus using the REQUIL model and Soavee RedlicheKwong equation of state. When only the MeOH synthesis reactions are considered, the CO equilibrium conversion does not exceed 50%. The equilibrium conversion decreases with temperature (because the reactions are exothermic) and increases with pressure (because the number of moles decreases). When the DME synthesis is added to the reaction network, the CO conversion dramatically increases. However, as important amounts of CO2 are formed, the equilibrium yield is far from the maximum value of 0.5 kmol DME/kmol CO (Fig. 10.2, right). In practice, various catalysts are applied for the one-step DME synthesis (Sun et al., 2014). They include either two separate beds for methanol and DME synthesis or a single dual-catalyst bed, which has a methanol synthesis component (usually Cu-based) and a methanol dehydration component, such as alumina, silica-alumina, and zeolites (Ptasinski, 2016). In this respect, the catalyst becomes important as it should increase the selectivity by favoring the hydrogenation reaction (10.1) over the WGS reaction (10.3).

10.2.3 Kinetics of Methanol Dehydration DME is formed in the reversible chemical reaction of methanol dehydration, which takes place either in gas phase (catalyzed by g-alumina) or in liquid phase (catalyzed by ion-exchange resins). 2CH3 OH%CH3 OCH3 þ H2 O

(10.6)

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Gas-phase reaction: The kinetics of the gas-phase methanol dehydration, on a commercial g-Al2O3 catalyst, was studied experimentally in the temperature range 250e400 C and pressures up to 10 bar (Bercic and Levec, 1993). The authors provide two rate equations. The intrinsic rate was developed using the Langmuir-Hinshelwood-Hougen-Watson (LHHW) formalism, assuming that the dissociative adsorption of methanol on the catalyst surface takes place, and the surface reaction is the controlling step. The global reaction rate was measured under several experimental conditions. Then, a reactionediffusion model was developed, for which the transport parameters (conductivity and effective diffusivity) were predicted based on literature data and correlations. The intrinsic kinetic parameters were determined by fitting the model prediction to the experimental data. However, using the intrinsic kinetics for designing a reactor in commercial process simulation software (such as Aspen Plus) is cumbersome because one would have to write a user kinetic subroutine to calculate the global reaction rate by solving the reactionediffusion equations. Bercic and Levec (1993) also provide an apparent reaction rate (10.7) and kinetic parameters which predict well the temperature and conversion profiles in a catalytic reactor. This can be easily used in a commercial process simulation software. However, the original apparent rate constant ks had to be increased by a factor 100 to match the effectiveness factor presented in the same paper. Other similar kinetics are reported in literature for temperatures of 250e400 C and pressures up to 10 bar (Mollavali et al., 2008).   yW yD 2 ks KM2 yM  Keq (10.7) r ¼ pffiffiffiffiffiffiffiffiffiffiffiffiffi 4 ð1 þ 2 KM yM þ KW yW Þ In this chapter, the following constants were used: ks ¼ 3.30  109 exp(10,800/T ), kmol/(kg h) KM ¼ 0.72  102 exp(830/T ), dimensionless KW ¼ 0.45  102 exp(1130/T ), dimensionless The equilibrium constant Keq was calculated (by Aspen Plus) from Gibbs free energies. Liquid-phase reaction: Methanol dehydration can also take place in liquid phase, in the presence of different solid acidic catalysts. Note that performing the reaction in liquid phase makes possible the application of reactive distillation (RD) technology. For example, Hosseninejad et al. (2012) determined the initial reaction rate of the liquid-phase methanol dehydration on Amberlyst-35, which could be described by the following rate equation.

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r ¼ 

ks K W cW 1þ KM cM

 2

1

1 cD cW 2 Keq cM

 (10.8)

The authors (Hosseninejad et al., 2012) performed experiments at 9 bar and 110e135 C using catalysts of 0.2e0.6 mm diameter. They provide the following kinetic parameters: ks ¼ 6.12  109exp(11,793/T ), kmol/(kg s) KW/KM ¼ exp(6.46 þ 2964.0/T ) In this work, we correct the kinetic expression provided by Hosseninejad et al. (2012) by including the equilibrium departure term. The equilibrium constant (Keq) was regressed from equilibrium values calculated by Aspen Plus from Gibbs free energies. In this way, the kinetic equation can also be applied to conditions in which methanol conversion approaches chemical equilibrium (e.g., in RD processes). Keq ¼ exp(2.6305 þ 2787/T ).

10.3 Conventional DME Process In the conventional process, the raw material is methanol, synthesized from syngas over a copper-based catalyst (Cu/Zn, Cu/Zn/Al, Cu/Zn/Co). The methanol dehydration reaction is carried out in a gas-phase reactor. Different types of solid acid catalysts can be used, such as g-alumina (g-Al2O3), HZSM-5, silica-alumina, phosphorous-alumina, and fluorinated-alumina. Among them, g-alumina is the preferred one because of its thermal stability, mechanical resistance, high surface area, and catalytic properties. The reaction takes place in a fixed-bed gas-phase reactor which is operated at 250e400 C and pressure up to 20 bars. Conversion of methanol lies within the 70%e80% range, depending on the catalyst and the operating conditions. Because of the incomplete conversion, the outlet of the reactor consists of a ternary mixture: DME, water, and methanol. Thus, the mixture is cooled and subsequently separated by distillation in a direct sequence: first column is required to deliver highpurity (minimum 99.99%wt) DME (Muller and Hubsch, 2005; Kiss and Ignat, 2013), while the unreacted methanol is separated from water in a second distillation column, and then recycled back to the reactor (Kiss and Ignat, 2013).

10.3.1 Process Design and Optimization Fig. 10.3 shows the process flowsheet, mass balance, and the key design parameters of the reactoreseparationerecycle (RSR)

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Figure 10.3 Conventional reactioneseparationerecycle process for dimethyl ether (DME) production.

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process for DME production (Bildea et al., 2017). The fresh methanol feed is mixed with the recycled methanol stream and fed via a feed-effluent heat exchanger (FEHE) to a furnace, which vaporizes the whole stream. Methanol vapor is converted in the reactor to DME and water, but the conversion is incomplete. The hot reactor outlet stream is used in FEHE to preheat the feed to the reactor and then the remaining heat of the hot stream is used for steam generation (which can be used for the reboiler of methanol recovery column). The outlet of the steam generator is fed directly, as vapor, to the DME distillation column (COL-1). Note that further cooling to liquid at boiling point (about 64 C) increases the total annual cost (TAC), as an additional heat exchanger is needed and transferring the heat to cooling water increases the duty required for DME separation. In COL-1, the DME is separated as top distillate, while the bottom product (mixture of methanol and water) is fed to the methanol recovery column (COL-2), which separates methanol as distillate (a stream that is recycled in the process) and water by-product as bottom stream. For the conventional process, the choice of the reactor is a major design decision. As the reaction is slightly exothermic (for pure methanol feed, the adiabatic temperature rise is about 125 C), the reaction can be performed in an adiabatic tubular reactor. This option was chosen in this study. The feed temperature should be as high as possible such that reaction proceeds at high rate, but the constraint on maximum temperature allowed by the catalyst (400 C) must be fulfilled. Another option is to use a multitubular cooled reactor. This has the advantage of being operated at higher temperature (using thus less catalyst for the same conversion) but is more complex to build and operate. The second decision concerns the conversion for which the reactor is designed. Higher conversion requires more catalyst. However, approaching the chemical equilibrium has a beneficial effect on the TAC, as the cost reduction of the separation section (CapEx and OpEx) overcomes the expense of using more catalyst. The process design is optimized using the minimization of the TAC as objective function (Bildea et al., 2017): TAC ¼ OpEx þ

CapEx payback period

(10.9)

A payback period of 3 years was used, and a running time of 8000 h/year was considered. The following heating and cooling costs were taken into account: high-pressure (HP) steam (42 bar, 254 C, $9.88/GJ), medium-pressure (MP) steam (11 bar, 184 C, $8.22/GJ), low-pressure (LP) steam (6 bar, 160 C, $7.78/GJ), and cooling water ($0.72/GJ). These costs of utilities are typical

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for a US plant (Luyben, 2011), but they might differ for other locations. The total investment costs (TIC) (CapEx) include the reactor, all heat exchangers, and distillation columns. The cost of the heat exchangers (reboiler, condenser, FEHE, steam generator, cooler) is given by (Dimian et al., 2014)   CHEX ðUS$Þ ¼ ðM&S=280Þ$ 474:7$A0:65 ð2:29 þ Fm ðFd þ Fp ÞÞ 

Cfurnace ðUS$Þ ¼ ðM&S=280Þ$ 15668$Q

 0:85

(10.10) ð1:27 þ Fc Þ

(10.11)

where M&S is the Marshall and Swift equipment cost index (M&S ¼ 1536.5 in 2012), A is the area (m2), Fm ¼ 1 (carbon steel), Fd ¼ 0.8 (fixed tube), Fp ¼ 0 (less than 20 bar). A heat transfer coefficient U ¼ 0.5 kW/m2/K was assumed to calculate the heat transfer area. For the reboilers, the design factor was taken as Fd ¼ 1.35. For the furnace, Q is the duty (in MW), while the correction factor was taken as Fc ¼ 1 (process heater, carbon steel, design pressure less than 40 bar). The distillation columns diameter (D) was obtained by the tray sizing utility from Aspen Plus, while the height was evaluated from the number of trays (NT), as H ¼ 0.6∙(NT  1) þ 2 (m). Afterward, the cost of the columns shell was calculated as   Cshell ðUS$Þ ¼ ðM&S=280Þ$ 957:9$D1:066 $H 0:82 $ð2:18 þ Fc Þ (10.12) The cost of the trays was given by (Dimian et al., 2014) Ctrays ðUS$Þ ¼ NT $ðM&S=280Þ$97:2$D1:55 $ðFt þ Fm Þ

(10.13)

with Ft ¼ 0 (sieve trays) and Fm ¼ 1 (carbon steel). and where Fc ¼ Fm Fp, Fm ¼ 1 (carbon steel) and Fp ¼ 1 þ 0:0074$ ðP  3:48Þ þ 0:00023$ðP  3:48Þ2 . The relationship for Cshell was also used to estimate the cost of the tubular reactor. For the catalyst (g-alumina, bulk density 882 kg/m3, particle size 3 mm), a purchased cost of 10 $/kg was used. When a single-bed adiabatic reactor is used, the design decisions concern reactor size (diameter and length) and reactor inlet temperature. The diameter was calculated by imposing a fluid velocity u ¼ 0.25 m/s, which is in line with the reported value of an industrial DME reactor (Bai et al., 2013). Note that a velocity larger than 2 m/s results in excessive pressure drop while a lower velocity decreases the mass transfer rate from the fluid to the solid

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catalyst particle. Concerning the reactor length, a longer reactor (more catalyst) leads to higher conversion and thus reduces the separation costs, but pressure drop increases and the reactor is more difficult to build and operate. As the selectivity is very high in DME synthesis by methanol dehydration, the main issue here is the increase of pressure drop: if more catalyst is needed, this cannot be achieved by larger cross-sectional area of the reactor because this decreases the velocity. A higher feed temperature increases the reaction rate and therefore reduces the amount of catalyst required to achieve a certain conversion. On the other hand, more heat is required and the achievable conversion is reduced because of the exothermal nature of the chemical reaction. Moreover, the temperature along the reactor bed is increased, with negative effect on catalyst stability. Fig. 10.4 presents the TAC versus inlet reactor temperature for different lengths (Bildea et al., 2017). For each reactor, the optimal operating point TAC* is marked, together with the optimal values of TIC* (same as CapEx), total operating cost TOC* (same as OpEx), pressure drop along the reactor DP, and maximum temperature along the reactor bed Tout. It can be observed that the reactor leading to minimum TAC while fulfilling the constraint of outlet temperature below 400 C (when the catalyst starts to deactivate) has 12 m length, being fed at 275 C. This will be considered for further comparison with other design alternatives. Fig. 10.5 presents lines of constant reaction rate in the temperatureeconversion plane (Bildea et al., 2017). The line of

Figure 10.4 Total annual cost (TAC) versus reactor inlet temperature for reactors with various lengths. TAC, pressure drop, and reactor outlet temperature are indicated for each optimal point.

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Figure 10.5 Methanol conversion versus temperature at fixed reaction rate.

maximum reaction rate (at a given conversion) is represented by the dashed line. Also, the figure shows the reaction trajectories corresponding to reactors of length 8 m, 12 m, and 16 m, fed at the optimum temperature. It can be noticed that all the optimal reactors approach the equilibrium; therefore the reduction of separation costs exceeds the cost of using more catalysts. While evaluating the TAC, the performance of the distillation columns was specified in terms of distillate purity (99.99%wt DME, COL-1) and bottoms purity (>99.95%wt water, COL-2). The usual design condition of R ¼ 1.2 Rmin (Luyben, 2011) was used, while the location of the feed tray was chosen such that the minimum reboiler duty is achieved. The specification for the methanol recycle purity (in the range 0.95%wt to 0.999%wt) had little influence on the optimal reactor design and minimum TAC value. The DME column (COL-1) was operated at 10 bar, which allows use of cooling water in the condenser. The FEHE increases the temperature of the reactor feed stream up to 140 C by using the reactor effluent as heat source. Note that the amount of heat that can be recovered in the FEHE is limited by the temperature crossover which occurs when higher outlet temperature of the cold stream is specified. Tables 10.1 and 10.2 list the simulation results of the RSR process for DME production, while Table 10.3 provides a summary of the economic results (Bildea et al., 2017). Note that the furnace

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Table 10.1 Simulation Results for the ReactoreSeparationeRecycle Process for Dimethyl Ether Production (100 Ktpy): Reaction System (Without Downstream Processing). Operating Unit

Parameter/Unit

Value

Reactor

Length/[m] Diameter/[m] CapEx [k$] Area/[m] Duty/[MW]Tab CapEx [k$] Duty [MW] CapEx/[k$] OpEx/[k$/year] Duty/[MW] Area/[m2] CapEx/[k$] OpEx/[k$/year]

12 1.89 754.48 12.32 2.29 41.18 6.88 1006 1983.7 0.576 33.3 104.1 129.21

FEHE

Furnace

Steam generator

contributes by 29% to CapEx and 61% to OpEx, so heat recovery (using FEHE, steam generation, and vapor feed to the DME column) plays an important role in the process. The specific energy requirements are 2.58 MJ/kg (714 kWh/ton) DME. Alternatively, the energy efficiency of the DME separation can be enhanced by recycling methanol containing water (e.g., partially by-passing Col-2 and recycle or by using a low-purity distillate of Col-2 as methanol recycle), but this requires a water-tolerant catalyst (e.g., K-modified H-ZSM-5). The conventional RSR process could be further improved by making the following key changes, as described in a more recent short communication (Luyben, 2017): the liquid methanol feed is preheated close to the boiling point, vaporized in an evaporator, and superheated in a furnace. This solution (which is a standard engineering practice for evaporation) allows for better heat integration and reduces energy requirements. A cooled reactor is used instead of an adiabatic reactor, allowing the same conversion with less catalyst. Moreover, operating the methanol column under vacuum (at the lowest temperature which allows use

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Table 10.2 Simulation Results for the ReactoreSeparationeRecycle Process for Dimethyl Ether (DME) Production (100 ktpy): Downstream Processing (Direct Distillation Sequence). DME Column

Methanol Column

667.25 170.0

394.25 86.5

18 10 1.68 11.6 44.5 681.1

30 21 1.16 18.8 77.0 144.7

273 394.3 3.94 0.625 6.45 10 99.99 e

121.27 273 1.37 2.044 2.82 1 99.95 99.98

911.4 558.9 111.7 221.6 19.2 282.2 133.9 148.2 586.0

601.0 204.2 159.2 218.2 18.9 516.6 58.5 458.1 716.9

Feed Condition

Feed rate/[kmol/h] Feed temperature/[ C] Design Data

Number of stages, NT Feed stage, NF Diameter/[m] Height/[m] Reboiler area/[m2] Condenser area/[m2] Operating Data

Distillate rate/[kmol/h] Bottoms rate/[kmol/h] Reflux ratio Reboiler duty/[MW] Condenser duty/[MW] Condenser pressure/[bar] Distillate purity/[%wt] Bottoms purity/[%wt] Economic Data

CapEx/[k$] Condenser/[k$] Reboiler/[k$] Column shell/[k$] Trays/[k$] OpEx/[k$/year] Cooling/[k$/year] Heating [HP steam]/[k$/year] Total annual cost/[k$/year]

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Table 10.3 Summary of Economic Results for the ReactoreSeparationeRecycle Process for Dimethyl Ether (DME) Production (100 ktpy). Reactor FEHE Furnace Steam generator DME column Methanol column Total

CapEx/[k$]

OpEx/[k$/year]

TAC/[k$/year]

754.5 41.2 1006.4 104.1 910.8 600.6 3417.6

e e 1983.7 129.2 281.9 516 2652.4

251.5 13.7 2319.2 94.5 585.5 716.2 3791.6

of cooling water in the condenser) increases the relative volatility; therefore, the methanol/water separation is easier. This alternative, which decreases the energy costs by 32% but requires investing 53% more capital, is left as an exercise for the reader.

10.3.2 Intensified Separations Using Dividing-Wall Column Technology This section shows that the conventional direct distillation sequence for the DME separation and methanol recovery can be successfully replaced with a single-step separation process in a dividing-wall column (DWC), with significant economic benefits. Moreover, the possibility of reusing the existing equipment is especially beneficial for revamping existing DME production plants. A single-step DME separation process is described, taking place in a DWC that effectively integrates into one shell the separation tasks of DME purification and methanol recovery (Kiss and Ignat, 2013). The DWC technology is one of the best examples of process intensification applied at industrial scale, as it offers large savings in the capital and operating costs (20%e35%), as well as in the equipment footprint. Several comprehensive reviews and research papers were published on the DWC topic, covering the design, simulation, control, optimization, and applications of DWC (Dejanovic et al., 2010; Yildirim et al., 2011; Kiss and Bildea, 2011).

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Figure 10.6 Dimethyl ether (DME) separation alternatives using Dividing-wall column (DWC) technology.

For the separation of the ternary mixture DMEemethanole water (outlet from the reactor), there are three potential configurations using a DWC, with the partition wall located in the middle, in the top, or in the bottom of the column, as shown in Fig. 10.6 (Kiss and Ignat, 2013). In all these configurations, direct steam injection can be used instead or in addition to the reboiler producing water vapors. The top wall system is more suitable for higher water content in the feed, while the setup with a bottom wall system is preferred in case of high DME content in the feed. However, they both suffer from larger temperature difference across the wall, as compared to a classic midwall DWC. Moreover, because of the reaction stoichiometry, the molar amount of DME and water in the feed is equaldhence the midwall DWC configuration is the preferred one. An industrial process for DME production by methanol dehydration is considered in this study, with a production rate of 100 ktpy DME. Process simulations were carried out in Aspen Plus for the base case scenario and the DWC alternative proposed, using the rigorous RADFRAC unit with the equilibrium model enabled and 80% stage efficiency (Kiss and Ignat, 2013). Because of the incomplete conversion of methanol (76%) in the gas-phase reactor, a ternary mixture of DMEemethanole water is produced: 22,880 kg/hr consisting of 24% methanol, 38% DME, and 38% water (mol base). The feed stream of the distillation sequence is the reactor outlet stream that is cooled to 120 C at 10 bar before being distilled. Note that the methanol conversion considered in this section (76%) is the typical for industrial operation but lower compared to the optimized RSR plant previously described (82%). Therefore, the economics of the separation section is worse. This proves that it is beneficial to use more catalyst, as the additional reactor-related costs are overcome by the reduction of separation costs.

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381

Table 10.4 Design Parameters of the Conventional Direct Sequence of Two Distillation Columns (for dimethyl ether (DME) Synthesis by Methanol Dehydration). Design Parameters

Value

Unit

Flow rate of feed stream Feed composition (molar fractions) DME:methanol:water Temperature of feed stream Pressure of feed stream Operating pressure (DC1) Operating pressure (DC2) Column diameter (DC1) Column diameter (DC2) Total number of stages (DC1) Total number of stages (DC2) Feed stage (DC1) Feed stage (D-C2) Reflux ratio (DC1) Reflux ratio (DC2) Distillate to feed ratio (DC1) Distillate to feed ratio (DC2) DME product purity Methanol recycle purity Water product purity Reboiler duty DC1 Condenser duty DC1 Reboiler duty DC2 Condenser duty DC2

22,880 0.38:0.24:0.38

kg/hr e

120 10 10.0 1.0 1.3 1.7 32 42 16 23 1.87 2.04 0.546 0.529 99.99/99.99 99.99/99.99 99.99/99.99 1400 3730 4200 5123

C

Table 10.4 provides the design and operating parameters of both distillation columns of the direct sequence (Fig. 10.3), based on an industrial DME process (Kiss and Ignat, 2013). The first column separates the pure DME product while the second one removes water by-product and recovers the unreacted methanol that is recycled back. The DME purification column has 32 stages and it is operated at 10 bar, with a reflux

bar bar bar m m e e e e kg/kg kg/kg kg/kg kg/kg %wt/%mol %wt/%mol %wt/%mol kW kW kW kW

382

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ratio of 1.87 and a reboiler duty of 1400 kW. The methanol recovery column has 42 stages, being operated at atmospheric pressure, using a reflux ratio of 2.04 and a reboiler duty of 4200 kW. It should be noted that the two distillation columns of the direct sequence are operated at different pressures, while a DWC can be operated at only one pressure. This aspect is important for the feasibility of using a DWC, whenever there are limitations regarding the maximum operating temperature and pressure. But fortunately this is not the case here because both distillations can be carried out at the same pressure, as in the DME purification column. For the optimal design of the DWC, the sequential quadratic programming optimization method and the sensitivity analysis tool from AspenTech Aspen Plus were employed. The main objective of the optimization procedure is to minimize the total reboiler duty required, as follows: Min (Q) ¼ f (NF, NSS, NDWS, NDWC, V, RR, rV, rL) (10.14) Subject to ! ym! x m where the optimization parameters used here are feed location (NF), side-draw stage (NSS), wall size (NDWS) and location (NDWC), boil-up rate (V ), reflux ratio (RR), liquid and vapor split (rL and rV), while ym and xm are the vectors of the obtained and required purities for the m products. In view of the revamping approach, the total number of stages (NT) was considered equal to the number of stages of the longest column in the conventional sequence. Fig. 10.7 presents the Vmin diagram of the ternary system DMEemethanolewater (considering the feed flow and feed

Figure 10.7 Vmin diagram for the ternary mixture dimethyl ether (DME)emethanolewater.

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Figure 10.8 Temperature (left) and composition (right) profiles in the Dividing-wall column unit. DME, dimethyl ether.

composition from Table 10.4)dnoted here as A, B, C for convenience (Kiss and Ignat, 2013). The Vmin diagram is a straightforward and common measure used to compare the energy requirements in different distillation arrangements. The minimum amount of energy required for the separation in a DWC is given by the highest peak in the diagramdmeaning that the rest of the separations are obtained “for free” (Kiss and Ignat, 2013). Fig. 10.8 illustrates the temperature and composition profiles along the DWC, while Table 10.5 provides the design and operating parameters of the optimal DWC unit (Kiss and Ignat, 2013). DME and water are the top and bottom end high-purity products (>99.99%wt), while methanol accumulates toward the middle of the column, being withdrawn as a side stream (>99%wt) and then recycled. The temperature difference between the two sides of the wall is low, such conditions being feasible for practical implementation, with little heat transfer expected and negligible effect on the column performance. For the cases when the heat transfer cannot be neglected at all, an insulated partition wall can be considered for the practical implementation (Kiss, 2013). The two process alternatives were compared in terms of TIC, TOC, and TAC. The equipment costs were estimated using correlations from the Douglas textbook to the price level of 2010, as described by Dejanovic et al. (2010). Table 10.6 provides a head-to-head comparison of the key performance economic indicators (Kiss and Ignat, 2013). Remarkable is the fact that the DWC alternative requires less equipment and 20% lower capital costs, while being the most energy efficient allowing significant energy savings of over 28%, as compared to the conventional distillation sequence considered here. The new separation

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Table 10.5 Design and Operating Parameters of a Dividing-wall column (DWC) for Single-Step Dimethyl Ether (DME) Separation (for DME Synthesis by Methanol Dehydration). Design Parameters

Value

Unit

Flow rate of feed stream Feed composition (molar fractions) DME:methanol:water Temperature of feed stream Pressure of feed stream Operating pressure Column diameter Number of stages prefractionator side Total number of stages DWC Feed stage prefractionator Side stream withdrawal stage Wall position (from/to stage) Distillate to feed ratio Reflux ratio Liquid split ratio (rL) Vapor split ratio (rV) DME product purity Methanol recycle purity Water product purity Reboiler duty Condenser duty

22,880 0.38:0.24:0.38

kg/hr e

120 10 10 1.7 17 42 12 20 16e32 0.546 3.87 0.27 0.42 99.99/99.99 99.10/98.40 99.99/99.99 4028 6322

C

Bar Bar M e e e e e kg/kg kg/kg kg/kg kg/kg %wt/%mol %wt/%mol %wt/%mol kW kW

scheme also requires less equipment units and reduced plant footprintdthus sparing existing equipment (column shell, heat exchangers), useable elsewhere in the chemical plant. Moreover, the revamping of the existing equipment is actually possible, within a very reasonable payback time of half a year.

10.4 Novel Process Intensification Alternatives To increase the capacity of DME production, the chemical industry needs novel eco-efficient processes that can meet the growing

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Table 10.6 Comparison of Key Indicators for the Conventional Sequence versus Dividing-wall column (DWC) Alternative. Key Performance Indicators

Conventional Process

DWC Alternative

Total investment cost (TIC) Total operating costs (TOC) Total annual costs (TAC) Specific energy requirements (kW$h/ton DME) CO2 emissions (kg CO2/h$ton DME)

$1,760,752 $1,388,550 $1,564,625 448.0 569.7

$1,412,490 $997,735 $1,138,984 322.2 409.8

market demand (Azizi et al., 2014). The suggested improvements of the methanol dehydration process followed a number of technologies reported in literature: fully thermally coupled distillation columns (Petlyuk) or DWC replacing the direct distillation sequence (Kiss and Ignat, 2013); self-heat recuperation (Kansha et al., 2015); simultaneous synthesis and separation in a catalytic distillation process (An et al., 2004; Lei et al., 2011); catalytic cyclic distillation (Patrut et al., 2014); or reactive DWC (Kiss and Suszwalak, 2012). Fig. 10.9 illustrates some of these DME process intensification alternatives. Additional DME technologies (e.g., coupled and dual-type reactors, microreactors, membrane reactors, and spherical reactors), including the direct route from syngas, are also well described in the review paper of Azizi et al. (2014). Among these technologies, RD seems to be the most promising (Azizi et al., 2014), being a proven process intensification method that effectively combines the reaction and separation into a single unit. RD can considerably improve the performances of an equilibrium-limited process, by pulling the conversion to completion (thus avoiding recycles), increasing selectivity and productivity, reducing the energy use, and overall reducing the CapEx and OpEx.

10.5 Catalytic Distillation Process RD conveniently combines reaction and separation in a single unit that can drive conversion to completion by continuously removing the products from the system. In such processes,

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386

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M

Figure 10.9 Novel dimethyl ether (DME) process intensification alternatives.

designing an RD column at the maximum driving force results in an optimal design in terms of controllability and operability, which is less sensitive to disturbances in the feed and has the inherent ability to reject disturbances (Mansouri et al., 2015). As shown in Fig. 10.10 (Bildea et al., 2017), methanol is fed at the top of the reactive zone in the RD column (on stage 9). The reaction takes place in the reactive zone where the solid acid catalyst is placed (stages 9e42). DME is the lightest component in the system and hence removed as top distillate, while water is the heaviest component and hence removed as bottoms (this stream could be used for alternative energy recovery options). Unlike the gas-phase reaction, the dehydration of methanol in liquid phase is catalyzed by thermally stable resins, such as Amberlyst-35, which has high activity and selectivity at temperatures up to 150 C. The optimization was performed using the TAC as objective function. The price of the structured packing (KATAPAK-SP 11) was taken as 10,000 $/m3. A purchased cost of 10 $/kg was considered for the solid acid catalyst (Amberlyst-35 ion-exchange resin, 560 kg/m3 bulk density). The optimal design of the RD process (and the next one combining gas-phase reaction and RD) is a mixed-integer

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387

Figure 10.10 Single-step process using a reactive distillation column for dimethyl ether (DME) production.

nonlinear problem (MINLP). Such problems are intrinsically very difficult to solve, and the solution time increases rapidly with the number of variables and constraints. The decision variables used in the optimization are the following: • Discrete (integers): number of stages (NT), feed stage (NF), reactive stages (NR1 e NR2) • Continuous (real numbers): reflux ratio (RR), pressure (P), amount of catalyst/tray (mcat) The optimization considers the following constraints: • Distillate purity: 99.99%wt DME • Bottoms purity: exceeding 99.95%wt • Temperature on the reactive stages: below 150 C (to avoid catalyst deactivation) • Catalyst must fit into the available space (no more than 20% of the packing volume) Note that the optimal steady-state operating point is often defined by the intersection of active constraints (Kookos and Perkins, 2016). This well-established feature of chemical processes is neatly exploited to simplify the solution of the optimization problem. Thus, the RD column for DME synthesis has the following features, confirmed by several sensitivity analysis runs: • When the distillate rate is set to half of the feed (based on reaction stoichiometry), because of the relative volatilities of

388

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the species involved in the process, methanol is equally distributed between the distillate and the product streams. Therefore, achieving 99.99%wt distillate purity (DME) by manipulating the reflux ratio also ensures that the purity of the bottoms stream (water) exceeds the 99.95%wt requirement. • The operating costs (OpEx) represent more than 75% of the TAC. The main contributions to the capital costs (CapEx) are the condenser, reboiler, column shell, and structured packing, with the catalyst representing less than 3%. For this reason, increasing the amount of catalyst per theoretical stage invariably results in a lower value of the TAC, as the reduction of the other costs clearly outweighs the cost of catalyst. However, the amount of catalyst is limited to 20% of the available volume (Götze et al., 2001). • Increasing the operating pressure leads to higher temperature along the RD column. This results in faster reaction rates and, in turn, to lower TAC because of reduced effort necessary to achieve the separation. The optimal pressure leads to highest allowable temperature in the reactive section. Based on these observations, the following optimization strategy is employed. In Aspen Plus, set the discrete decision variables: number of stages, feed stage, and reactive stages. The distillate molar rate is set to half of the feed, according to reaction stoichiometry. By means of Design Specification blocks, the continuous decision variables (reflux ratio, pressure, and amount of catalyst on each tray) are adjusted such that the constraints are fulfilled (distillate purity 99.99%wt DME; temperature on reactive stages < 150 C; catalyst fits into the available space of 20% of the packing volume). Because of mass balance and high purity of the distillate, the purity of the bottoms stream always exceeds the 99.95%wt requirement. The simulation is run and the value of the TAC is obtained. Afterward the discrete variables are adjusted, and new runs are performed, until the minimum TAC is obtained. In this work, the optimal values of the discrete variables are found using the genetic algorithm implemented in Matlab by means of the ga function. The user-provided Matlab objective function uses the COM interface to communicate with Aspen Plus: sends to Aspen Plus the values of the discrete decision variables, requests the simulation run, and checks the convergence. After a successful run, the objective function (TAC, calculated in Aspen Plus by a FORTRAN Calculator block) and the continuous decision variables (obtained by the Design Specification blocks) are retrieved and saved in an Excel file

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389

Figure 10.11 Temperature and reaction rate profiles (left) and L-V composition profiles (liquiddcontinuous line; vaporddashed line) along the catalytic distillation column (right). DME, dimethyl ether.

together with the discrete decision variables. Then, the value of the objective function is passed to the ga function, which performs a new iteration. If Aspen Plus simulation does not converge, the simulation is reinitialized and a new run is attempted. If this also fails, a large value is returned to ga as the value of the objective function. During an initial run, rather large ranges of the discrete variables were assumed and many Aspen Plus simulations were unsuccessful. After a better estimation of the optimum, the ranges were restricted as follows: feed tray: 7 . 12; first reactive stage: 7 . 12; last reactive stage 30 . 50; total number of stages: 45 . 60. In this case, more than 95% of the simulations were successful. Fig. 10.10 also shows the process flowsheet, mass balance, and the key operating parameters, while Fig. 10.11 plots the temperature and reaction rate profiles, as well as the liquid and vapor composition profiles along the RD column (Bildea et al., 2017). In addition, Table 10.7 provides the optimal design parameters of the RD column for DME production (Bildea et al., 2017). The operating pressure (11.36 bar) leads to reactive stages temperature in the range 130e150 C. The amount of catalyst (193.2 kg/stage) occupies 20% of the stage volume, this being in line with the recommendations for KATAPAK packing (Götze et al., 2001). The reflux ratio (6.174 kmol/kmol) ensures the required purity. Along the reactive stages, the liquid-phase methanol mole fraction exceeds 0.80, which leads to a large reaction rate. For the RD process, the CapEx is only 2395 k$, the OpEx is 2604 k$/year, and the energy needs are 2.43 MJ/kg (672.5 kWh/t) DME.

Table 10.7 Optimal Design Parameters of a Reactive Distillation Column for Dimethyl Ether Production (100 ktpy). Parameter/Unit

Value

Design Data

Number of stages, NT Feed stage, NF Reactive stages, NR1 e NR2 HETP/[m] Amount of catalyst/[kg/stage] Catalyst volumetric fraction Diameter/[m] Height/[m] Reboiler area/[m2] Condenser area/[m2] Feed preheater/[m2]

54 9 9e42 0.5 193.2 0.2 2.1 28.0 248.8 755.0 132.3

Operating Data

Feed rate/[kmol/h] Feed temperature/[ C] Distillate rate/[kmol/h] Bottoms rate/[kmol/h] Reflux ratio Reboiler duty/[MW] Condenser duty/[MW] Feed preheater/[MW] Condenser pressure/[bar] Distillate purity/[%wt DME] Bottoms purity/[%wt water]

546 85.4 273 273 6.177 8.485 9.148 0.943 11.36 99.99 99.97

Economic Data

CapEx/[k$] Condenser/[k$] Reboiler/[k$] Column shell/[k$] Packing/[k$] Catalyst/[k$] Trays/[k$] Feed preheater OpEx/[k$/year] Cooling/[k$/year] Heating (HP steam]/[k$/year] Total annual cost/[k$/year]

2394.9 597.6 342.2 578.4 586.5 65.7 31.9 192.6 2604.2 189.7 2414.5 3402.5

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Notably, Oberon Fuels (www.oberonfuels.com) launched the first commercial catalytic distillation column producing 9700 ton/year of fuel-grade DME (98.5%wt), in Brawley (California, US).

10.6 Combined Gas-phase Reactor and Reactive Distillation Process The main idea of the combined process is to use only the reaction section of existing RSR processes for DME production and include an RD column in the downstream processing, as shown in Fig. 10.12 (Bildea et al., 2017). This allows complete reactant conversion and avoids the need of recycling. The reactor diameter decreases, but the effect on the TAC is small. However, it should be remarked that the gas-phase reactor of existing DME plants can process more methanol (for example, the RSR process illustrated in the previous section could process 748 kmol/h instead of 546 kmol/h). The methanol conversion is eventually driven to completion by the RD column. As the feed to the downstream processing section is not pure methanol (as in the single-step RD process), but a ternary mixture (DME, methanol, and water from the gas-phase reactor), an additional unit is necessary. Because water has the highest boiling point (and therefore in the RDC column will remain in the liquid phase) and has a detrimental effect on the reaction rate, this component is removed in the first distillation column. The water column was designed for 99% water recovery, at 99.97%wt purity. The methanoleDME mixture is fed above the reactive section of the RDC. In the RD column, the room for the amounts of catalyst (67 kg/stage) is the maximum, occupying 20% of the available space. Despite the higher methanol liquid-phase concentration, the reaction rate is lower compared to the single RD column process because of less catalyst. Along the reactive stages, the temperature is in the range 134e147 C. Tables 10.8 and 10.9 list the simulation results of the combined process for DME production, while Table 10.10 (Bildea et al., 2017) provides a summary of the economic results. For the same production capacity (100 ktpy), the cost of the reaction section is lower as compared to the classic RSR process because of the absence of the methanol recycle. However, the downstream processing section is somewhat more expensive than the classic

391

392 Chapter 10 DIMETHYL ETHER

Figure 10.12 Combined gas-phase reactor and reactive distillation process for dimethyl ether (DME) production.

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393

Table 10.8 Simulation Results for the Combined Gas-phase Reactor and Reactive Distillation Process for Dimethyl Ether Production (100 ktpy): Gas-phase Reactor Section. Parameter/Unit

Value

Reactor

Length/[m] Diameter/[m] CapEx [k$]

12 1.71 231.8

FEHE

Area/[m] Duty/[MW] CapEx [k$]

11.8 1.981 7.65

Furnace

Duty [MW] CapEx/[k$] OpEx/[k$/year]

5.63 848.66 1671.55

Steam Generator

Duty/[MW] Area/[m2] CapEx [k$] OpEx [k$/year]

0.366 25.1 86.54 82.08

direct distillation sequence, as the reaction must also be carried out together with the separation. Moreover, the reactor outlet is not fed as vapor to the RD column (because of the reaction taking place in liquid phase) but as saturated liquid (condensed in the prefractionator). As a consequence, the key economic indicators are slightly increased: CapEx is 3437 k$, OpEx is 3152 k$/year, and the specific energy requirements are 2.905 MJ/kg (807 kWh/ton) DME. Nonetheless, the key advantage of this process alternative is that because of the elimination of the methanol recycle, the production capacity could be increased by a factor equal to the reverse of the methanol conversion (1/XMeOH).

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Table 10.9 Simulation Results for the Combined Gas-phase Reactor and Reactive Distillation Process for Dimethyl ether (DME) Production (100 ktpy): Downstream Processing Section. Parameter/Unit

Water Column

RD Column

546 164.2

324.8 66.6

18 9 e 1.38 112 1066 e e

35 8 8e25 1.25 21.8 90.8 386.4 67 0.2

324.8 221.2 1.867 1.727 7.118 5 0.003 (water) 0.9997 (water)

273 51.8 1.937 3.159 3.822 10.2 0.9999 (DME) 0.9997 (water)

1143.8 747.9 203.6 177.5 14.8 e 556.6 147.6 409.0 936.8

956.7 386.8 177.8 270.2 13.3 108.6 1006.7 79.3 927.4 1325.6

Feed Condition

Feed rate/[kmol/h] Feed temperature/[ C] Design Data

Number of stages, NT Feed stage, NF Reactive stages Diameter/[m] Height/[m] Reboiler area/[m2] Condenser area/[m2] Catalyst/[kg/stage] Catalyst volumetric fraction [%] Operating Data

Distillate rate/[kmol/h] Bottoms rate/[kmol/h] Reflux ratio Reboiler duty/[MW] Condenser duty/[MW] Condenser pressure/[bar] Distillate purity/[%wt] Bottoms purity/[%wt] Economic Data

CapEx/[k$] Condenser/[k$] Reboiler/[k$] Column shell/[k$] Trays/[k$] Packing/[k$] OpEx/[k$/year] Cooling/[k$/year] Heating (HP steam]/[k$/year] Total annual cost/[k$/year]

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395

Table 10.10 Summary of Economic Results for the Combined Gas-phase Reactor and Reactive Distillation Process for Dimethyl Ether Production (100 ktpy). Reactor FEHE Furnace Steam generator Water column RD column Total

CAPEX/[k$]

OPEX/[k$/year]

TAC/[k$/year]

231.8 7.65 848.67 86.54 936.8 1325.6 3437.06

e e 1671.55 82.08 556.6 1006.7 3152.77

77.3 2.55 1954.44 53.23 868.86 1448.56 4298.4

10.7 Conclusions DME manufacturing is boosted today by the perspective of a wide use as sustainable fuel that behaves similarly with diesel or GPL. The classic process for DME production by methanol dehydration can be further improved by optimization, but above all by employing process intensification techniques. The DME purification and methanol recovery distillation sequence can be effectively converted into a single-step separation based on the DWC technology. As compared to the conventional direct sequence of two distillation columns, the DWC alternative reduces the energy requirements by 28% and the equipment costs by 20%. The new separation scheme also requires less equipment units and reduced plant footprintdthus sparing existing equipment (column shell, heat exchangers), useable elsewhere in the chemical plant. The revamping of the existing equipment is actually possible, within a very reasonable payback time. Also, the DME process simulations showed that the novel process alternatives based on RD can significantly improve the classic RSR process for DME production. Based on the results presented here, the following conclusions can be drawn: • The classic RSR DME process should be designed for high (close to equilibrium) methanol conversion. Further improvements are possible by advanced heat integration, such as feeding the reactor outlet as vapor to the direct distillation sequence (about 8% savings in TAC).

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• The overall energy efficiency could be further enhanced by recycling methanol containing water, but this requires a watertolerant catalyst (e.g., K-modified H-ZSM-5) in the reactor. • The liquid methanol feed could be preheated close to the boiling point, vaporized in an evaporator, and superheated in a furnace. This solution allows for better heat integration and reduces energy requirements. Moreover, a cooled reactor could be used instead of an adiabatic reactor, allowing the same conversion with less catalyst. • The single-step process using an RD column is the most promising for new DME plants because of the lowest CapEx (2395 k$) and OpEx (2604 k$/year)dfor 100 ktpy plant capacitydas well as specific energy requirements (2.43 MJ/kg DME). • The combined process (gas-phase reactor þ RD) is not recommended for new plants (because of higher costs than the single-step RD process), but it is suitable for revamping existing DME plants. With only a minor additional investment, the plant capacity can be significantly increased by having complete methanol conversion and no recycles, thus allowing a higher processing capacity of the fresh methanol feed.

References An, W., Chuang, K., Sanger, A., 2004. Dehydration of methanol to dimethyl ether by catalytic distillation. Canadian Journal of Chemical Engineering 82, 948e955. Arcoumanis, C., Bae, C., Crookes, R., Kinoshita, E., 2008. The potential of di-methyl ether (DME) as an alternative fuel for compression-ignition engines: a review. Fuel 87, 1014e1030. Azizi, Z., Rezaeimanesh, M., Tohidian, T., Rahimpour, M.R., 2014. Dimethyl ether: a review of technologies and production challenges. Chemical Engineering and Processing: Process Intensification 82, 150e172. Bai, Z., Ma, H., Zhang, H., Ying, W., Fang, D., 2013. Process simulation of dimethyl ether synthesis via vapor phase dehydration. Polish Journal of Chemical Technology 15, 122e127. Bildea, C.S., Gyorgy, R., Brunchi, C.C., Kiss, A.A., 2017. Optimal design of intensified processes for DME synthesis. Computers and Chemical Engineering 105, 142e151. Bercic, G., Levec, J., 1993. Catalytic dehydration of methanol to dimethyl ether. Kinetic investigation and reactor simulation. Industrial and Engineering Chemistry Research 32, 2478e2484.  2010. Dividing wall column - a Dejanovic, I., Matijasevic, L., Olujic, Z., breakthrough towards sustainable distilling. Chemical Engineering and Processing: Process Intensification 49, 559e580. Dimian, A.C., Bildea, C.S., Kiss, A.A., 2014. Integrated Design and Simulation of Chemical Processes, second ed. Elsevier, Amsterdam.

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Fleisch, T.H., Basu, A., Sills, R.A., 2012. Introduction and advancement of a new clean global fuel: the status of DME developments in China and beyond. Journal of Natural Gas Science and Engineering 9, 94e107. Götze, L., Bailer, O., Moritz, P., von Scala, C., 2001. Reactive distillation with KATAPAK®. Catalysis Today 69, 201e208. Hosseininejad, S., Afacan, A., Hayes, R.E., 2012. Catalytic and kinetic study of methanol dehydration to dimethyl ether. Chemical Engineering Research and Design 90, 825e833. Ihmels, E.C., Lemmon, E.W., 2007. Experimental densities, vapor pressures, and critical point, and a fundamental equation of state for dimethyl ether. Fluid Phase Equilibria 260, 36e48. Kansha, Y., Ishizuka, M., Song, C., Tsutsumi, A., 2015. Process intensification for dimethyl ether production by self-heat recuperation. Energy 90, 122e127. Kiss, A.A., 2013. Advanced Distillation Technologies - Design, Control and Applications. Wiley, Chichester, UK. Kiss, A.A., Bildea, C.S., 2011. A control perspective on process intensification in dividing-wall columns. Chemical Engineering and Processing: Process Intensification 50, 281e292. Kiss, A.A., Ignat, R.M., 2013. Revamping dimethyl ether separation to a singlestep process. Chemical Engineering and Technology 36, 1261e1267. Kiss, A.A., Suszwalak, D.J.-P.C., 2012. Innovative dimethyl ether synthesis in a reactive dividing-wall column. Computers and Chemical Engineering 38, 74e81. Kookos, I.K., Perkins, J.D., 2016. Control structure selection based on economics: generalization of the back-off methodology. AIChE Journal 62, 3056e3064. Lee, U., Han, J., Wang, M., Ward, J., et al., 2016. Well-to-wheels emissions of greenhouse gases and air pollutants of dimethyl ether from natural gas and renewable feedstocks in comparison with petroleum gasoline and diesel in the United States and Europe. SAE International Journal of Fuels and Lubricants 9. Paper 2016-01-2209. Lei, Z., Zou, Z., Dai, C., Li, Q., Chen, B., 2011. Synthesis of dimethyl ether (DME) by catalytic distillation. Chemical Engineering Science 66, 3195e3203. Luyben, W.L., 2011. Principles and Case Studies of Simultaneous Design. AIChE Wiley, Hoboken. Luyben, W.L., 2017. Improving the conventional reactor/separation/recycle DME process. Computers and Chemical Engineering 106, 17e22. Mansouri, S.S., Sales-Cruz, M., Huusom, J.K., Woodley, J.M., Gani, R., 2015. Integrated process design and control of reactive distillation processes. IFAC-PapersOnLine 48, 1120e1125. Mollavali, M., Yaripour, F., Atashi, H., Sahebdelfar, S., 2008. Intrinsic kinetics study of dimethyl ether synthesis from methanol on g-Al2O3 catalysts. Industrial and Engineering Chemistry Research 47, 3265e3273. Muller, M., Hubsch, U., 2005. Dimethyl ether. In: Ullmann’s Encyclopedia of Industrial Chemistry, seventh ed. Wiley-VCH, Weinheim. Patrut, C., Bildea, C.S., Kiss, A.A., 2014. Catalytic cyclic distillation - a novel process intensification approach in reactive separations. Chemical Engineering and Processing: Process Intensification 81, 1e12. Ptasinski, K.J., 2016. Efficiency of biomass energy: an exergy approach to biofuels, power, and biorefineries. Wiley, pp. 443e474 (Chapter 12): Dimethyl Ether (DME). Sardesai, A., 2006. Dimethyl ether. In: Lee, S. (Ed.), Encyclopedia of Chemical Processing. CRC-Press, Taylor & Francis.

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Chapter 10 DIMETHYL ETHER

Semelsberger, T.A., Borup, R.L., Greene, H.L., 2006. Dimethyl ether (DME) as an alternative fuel. Journal of Power Sources 156, 497e511. Sun, J., Yang, G., Yoneyama, Y., Tsubaki, N., 2014. Catalysis chemistry of dimethyl ether synthesis. ACS Catalysis 4, 3346e3356. Teodorescu, M., Rasmussen, P., 2001. High-pressure vapor-liquid equilibria in the systems nitrogen plus dimethyl ether, methanol plus dimethyl ether, carbon dioxide plus dimethyl ether plus methanol, and nitrogen plus dimethyl ether plus methanol. Journal of Chemical and Engineering Data 46, 640e646. Wu, J.T., Zhou, Y., Lemmon, E.W., 2011. An equation of state for the thermodynamic properties of dimethyl ether. Journal of Physical and Chemical Reference Data 40. Article Number: 023104. Yildirim, O., Kiss, A.A., Kenig, E.Y., 2011. Dividing-wall columns in chemical process industry: a review on current activities. Separation and Purification Technology 80, 403e417.