# Patent

## Title
Synthesis of DME using a fluid pluralized bed reactor

## Abstract

This invention discloses a method for making a dimethylether (DME) product from a synthesis gas (syngas) in the presence of a catalyst in a fluid pluralized bed reactor operating in the gas phase. The reactions generate a significant amount of heat and the heat management is balanced between supplying quench recycle syngas to the pluralized sections along the reactor and also by controlling the preheat temperature of the reactant streams. Gas phase fluidization of the catalyst is controlled so that the pluralized reactive zones are maintained in a backmix configuration.

## Background
REFERENCES CITED

















U.S. Pat. No. 4,417,000
November 1983
Slaugh et al.



U.S. Pat. No. 5,254,596
October 1993
Irick, Jr. et al.



U.S. Pat. No. 5,908,963
June 1999
Voss et al.



U.S. Pat. No. 6,191,175
February 2001
Haugaard et al.










    Lu, W. Z., Teng, L. H., Xiao, W. D., 2002. Simulation analysis of fixed-bed reactor for dimethyl ether synthesis. Natural Gas Chemical Industry 27 (4), 53-61
    U.S. Pat. No. 5,218,003 June 1993 Lewnard et al.
    US Patent Application Publication No. 2005/0038129 A1 February 2007 Lattner
    Brown, D. M., Bhatt, B. I., Hsiung, T. H., 1991. Novel technology for the synthesis of dimethyl ether from syngas. Catalysis Today 8, 279-304
    Guo, Q. G., Huang, Y. M., 1997. Studies on preparation of the catalyst for synthesizing of low carbon olefin from CO+H2. Natural Gas Chemical Industry (Cl Chemistry and Technology), 22(2)25-29
    Wang, Z. L., Wang, J. F., Diao, J., Jin, Y., 2001. The synergy effect of process coupling for dimethyl ether synthesis in slurry reactors. Chemical Engineering & Technologies 24 (5), 507-511.
    Liu, D. H., Xu, J. Zhang, H. T. Fang, D. Y. 2002 Direct synthesis of dimethyl ether from syngas in three phase agitated reactors. Chinese Journal of Chemical Engineering. 53(1): 103-106.
    Takashi Ogawa, Norio Inoue, Tutomu Shikada, Yotario Ohno, 2003. Direct Dimethyl Ether Synthesis. Journal of Natural Gas Chemistry 12, 219-227.
    Xiao, W. D., Lu, W. Z., 2002. A Novel Technology of DME Synthesis from Syngas. Patent No. CN1332141, China.
    Lu, W. Z., Teng, L. H., Xiao, W. D., 2004. Simulation and experiment study of dimethyl ether synthesis from syngas in a fluidized-bed reactor. Chemical Engineering Science 59, 5455-5464.


FIELD OF THE INVENTION
This invention relates to the production of dimethylether (DME). In particular, this invention relates to the production of DME using a fluid pluralized bed reactor.
BACKGROUND OF THE INVENTION
Current technology proposed three kinds of reactor for DME synthesis: the fixed bed (U.S. Pat. Nos. 5,254,596; 4,417,000; 5,908,963; 6,191,175; Xiao et al., 2002; Lu et al., 2002), the slurry reactor (U.S. Pat. No. 5,218,003; US Patent Application Publication No. 2005/0038129 A 1; Brown et al., 1991; Guo, et al., 1997; Wang, et al., 2001; Liu, et al., 2002; Takashi, et al., 2003) and the fluid bed reactor (Xiao, et al., 2002; Lu, et al., 2004). The process for making DME from a hydrogen and carbon monoxide syngas mixture is highly exothermic. The fixed bed reactor can only provide a limited heat removal capability, thus only a low conversion of the syngas can be achieved. At the conditions of P=4 MPa (565 psig), H2/CO molar ratio of 1.0, and gas space velocity of 15,000 ml/gcat/hr to keep the reactor temperature below 270° C. (518° F.), a fixed bed only yields 10.7% CO conversion and 91.9% DME selectivity (Lu, et al., 2002). On the contrary, the slurry reactor can provide an effective temperature control, but the two mass transfer resistances through the gas-liquid and the liquid-solid interfaces limit the DME productivity. In a slurry reactor under the conditions of T=260° C. (500° F.), P=4 MPa (565 psig), H2/CO molar ratio of 1.0, and gas space velocity of 3,000 ml/gcat/hr, the CO conversion and DME selectivity are 17.0% and 70.0% (Wang, et al., 2001).
In a fluid bed reactor, the gas-solid mass transfer resistance is so small that it can be neglected. Excellent temperature control is also achievable due to the vigorous mixing of catalyst particles in the fluid bed. Compared to a fixed bed or a slurry bed under the same conditions of the slurry bed, the CO conversion and DME selectivity in a fluid bed are 48.5% and 97.0% (Lu, et al., 2004). Therefore, the fluid bed reactor is the most promising for production of DME from syngas.
In this invention an improved embodiment to the fluid bed reactor called “fluid pluralized bed reactor” is described in the following sections, figures and examples to further promote effective removal of reaction heat in the reactor in order to minimize the recycle flow rate within the DME synthesis loop.
SUMMARY OF THE INVENTION
This invention provides a fluid pluralized bed reactor for producing DME from syngas. In particular, a relatively even temperature profile is achieved by the balancing of the heat generated within the reactor during the syngas to DME synthesis against the heat needed to bring the make-up syngas and recycle gas up to the desired DME reaction temperature.
The fluid pluralized bed is well mixed to provide proper heat transfer. In one embodiment, proper mixing is attained within the reactor by a combination of superficial velocities, aspect ratio of the catalyst bed and the catalyst particle size and density.
The fluid pluralized bed reactor temperature can be controlled either with or without the internal heat transmission tube in the reactor. A higher recycle rate is required when there is no internal heat transmission tube in the reactor.
The kinetics of the reaction are such that the bulk of the reaction is typically completed within a short distance downstream of the entry of the feed gases. In order to spread the reaction kinetics in a more isothermal fashion, it is convenient and appropriate to be able to introduce the feed gases in controlled quantities at several sections along the reactor. The feed gases in that case serve to also quench the reactor temperatures and bring them somewhat closer to an isothermal mode.
Each gas quench section also generates an environment of backmixing of catalyst solids that effectively distributes the heat generated uniformly throughout the bed. The circulating pattern of the catalyst lifts the catalyst upwards initially, picking up heat of reaction. This catalyst then circulates downwards to meet the fresh feed gases and preheat them quickly to reaction temperatures. The deployment of catalyst solids in this fashion greatly simplifies the kinetics and reduces or eliminates the amount of overall heat transfer surface area needed to control the process.
Appropriate injection of quench recycle syngas into different sections along the reactor then creates a series of independent, yet connected, backmix environments for optimizing the reactions and isothermality. In each section, temperature management can further be conducted by controlling the incoming temperature of the make-up syngas—whether they need to be cooled or heated is a function of the catalyst. In one embodiment of the invention, the fluid pluralized bed reactor is maintained at a temperature ranging from 150° C. to 350° C. (302° F. to 662° F.). Preferably, it is maintained at a temperature of 180° C. to 320° C. (356° F. to 608° F.), more preferably from 200° C. to 280° C. (392° F. to 536° F.).
In one embodiment, the temperature of the fluid pluralized bed is maintained by controlling the temperature of the feed gases entering the reactor through the reactor bottom gas distributor. In another embodiment, the feed gases may be entering the reactor at several different sections. The temperatures of the different incoming feed gas streams may be independently controlled to generate axial isothermality of the reactor. Preferably, the make-up syngas stream flowing into the reactor enters at ambient temperature, 37.8° C. (100.0° F.) while the recycle syngas enters at methanol absorber outlet temperatures ranging from −21° C. to −5° C. (−6° F. to 23° F.).



BRIEF DESCRIPTION OF THE DRAWINGS
FIG. 1 is the voidage variation along the height of a fluid pluralized bed.
FIG. 2 is a proposed fluid pluralized bed reactor equipped with three distributors.
FIG. 3 is a simplified process flow block diagram for the production of 150 tons/day of fuel grade DME from biomass.
FIG. 4 is a detailed process flow block diagram for the DME synthesis, gas cooling, acid gas removal and DME purification sections.



DETAILED DESCRIPTION OF THE INVENTION
Exothermicity & Equilibrium Considerations
The process of making dimethylether (DME) from a hydrogen and carbon monoxide syngas mixture is a strongly exothermic and equilibrium dictative reaction. Under LeChatelier principles, this process requires relatively high operating pressures and low temperatures to attain reasonable rates of reaction.















Exothermic



ΔH
Heat



















Methanol Synthesis from CO





2CO + 4H2   2CH3OH
{circle around (1)}
−197.48 kJ/mole*
74%


Methanol Dehydration to DME


2CH3OH   CH3OCH3 + H2O
{circle around (2)}
−21.26 kJ/mole
10%


Water Gas Shift


CO + H2O   CO2 + H2
{circle around (3)}
−39.57 kJ/mole
16%









Overall DME Synthesis Reaction
−258.31 kJ/mole 
100% 


Direct DME from Syngas


Producing CO2


3CO + 3H2   CH3OCH3 + CO2





*1 KJ/mole = 430.21 Btu/lb mole






All the three reactions are reversible and release a significant amount of heat for all these forward reactions. Consequently, a critical factor for DME reactor design is the management of the heat released by the reactions. The heat released by DME production can generate 2.4 tons of steam per ton of DME, equivalent to an adiabatic temperature rise of about 1,000° C. (1,832° F.) at a complete conversion of syngas with a 1:1 molar ratio of H2:CO in the feed gas.
However, the catalyst for Reactions {circle around (1)} and {circle around (3)} is subject to severe deactivation when overheated to above 280° C. (536° F.). To avoid thermodynamic limitations and excessive catalyst deactivation, conventional gas-phase reactors must be operated at a low per-pass conversion to maintain reactor temperature below 280° C. (536° F.), implementing a high syngas recycle rate, and resulting in large capital investments and operating costs.
Under such reaction conditions, the attainable conversion is strongly limited by the thermodynamic equilibrium. Finding a satisfactory compromise as to the reaction conditions between reaction rate and conversion percentage is therefore difficult. Effective control of the reaction temperature across the catalyst bed proved to be a technically problematic consideration.
In industrially applied processes, in which the catalyst is present in the form of a fixed bed of particles, high gas velocities are applied to promote effective removal of reaction heat and to allow good control of the reaction temperature. Due to these high velocities and the thermodynamic limitations, low CO conversions per pass are obtained. To achieve acceptable yields of DME from synthesis gas it is customary to recompress unconverted synthesis gas and recycle it to the reactor inlet. This requires recycle compressors of large capacities, which are costly and have high power consumptions. It is because of these constraints that using a fluid pluralized bed becomes an optimal solution.
Fluidization
The description of fluidization, its characteristics and attributes are based upon the seminal work of D. Kunii and O. Levenspiel, Fluidization Engineering, 1977.
Voidage variation along the height of a bubbling-fluid pluralized bed is carefully understood and implemented in the design and layout of the gas distributors. Broadly speaking, the bed comprises of two zones—a dense bubbling zone, and a lean fluidizing zone (FIG. 1). The bubbling zone has a lower entrance zone that has a slightly decreased solids density due to the adjacent gas distributor, and a main zone of constant density. An increase in gas velocity results in a linear decrease in bed density, and a commensurate increase of the height of constant density zone.
A dense bubbling bed has regions of low solid density, sometimes called gas pockets or voids. These regions are called bubble phase. The region of higher solid density is called the emulsion region or the dense phase. Within the dense phase, there is also an on-going solids circulation phenomena that is conventionally referred to as backmixing. Additionally, based upon the fluidization gas velocities, the bed may be operating in a bubbling, slugging or turbulent fluidization regime. Maximization of the backmixing phenomena is preferably attained under turbulent fluidization conditions.
At a certain height, the bubbling bed transitions into a lean fluidized zone of decreasing densities. There is minimal entrainment at the top of the lean fluidized zone and at a certain height the particulate entrainment becomes approximately constant. This is referred to as the transport disengaging height (TDH) and this is where the vessel exit is positioned. When an internal cyclone is used, the unit is placed below the TDH position and this results in an economy regarding the overall height and cost.
In practice, the gaseous reactant products leaving the top of the fluid pluralized bed are directed into a disengaging zone to separate the catalyst fines from the gases. In one embodiment, one or more cyclones are located below the TDH in the upper portion of the reactor. The cyclones are equipped with diplegs with the leg of the primary cyclone dropping into the bottom section of the fluid pluralized bed and the leg of the secondary or fine cyclone dropping into an area of the fluid pluralized bed above the prior dipleg.
Other embodiments for particulate separation may also be satisfactorily utilized. These embodiments include the use of blow back filters, either internal to the reactor or external. Other embodiments include increasing the height of the reactor and various other methods including meshes, plates, etc. In all cases, the fines separated by an appropriate embodiment should be directed back into the fluidized catalyst section.
Fluidization Parameters
Particle size is an important property that contributes to appropriate fluidization and appropriate backmixing of the catalyst within the designated reactor bed/section. In one embodiment of the invention, the catalyst bed includes catalyst particles having a particle size (i.e., average diameter) of from 20 to 300 microns. Preferably, the catalyst particles have a particle size of from 50 to 200 microns.
Superficial gas velocity (SGV) is a measurement of the gas flowing through the catalyst bed. It is defined as:




  
    Superficial
    ⁢
    
      
    
    ⁢
    gas
    ⁢
    
      
    
    ⁢
    velocity
    ⁢
    
      
    
    ⁢
    
      (
      
        m
        /
        sec
      
      )
    
  
  =
  
    
      
        
          
            
              Volumetric
              ⁢
              
                
              
              ⁢
              flow
              ⁢
              
                
              
              ⁢
              rate
              ⁢
              
                
              
              ⁢
              of
            
            ⁢
            
              
            
          
        
      
      
        
          
            gases
            ⁢
            
              
            
            ⁢
            leaving
            ⁢
            
              
            
            ⁢
            the
          
        
      
      
        
          
            reactor
            ⁢
            
              
            
            ⁢
            
              (
              
                
                  m
                  3
                
                /
                sec
              
              )
            
          
        
      
    
    
      
        
          
            Average
            ⁢
            
              
            
            ⁢
            cross
            ⁢
            
              
            
            ⁢
            section
          
        
      
      
        
          
            area
            ⁢
            
              
            
            ⁢
            of
            ⁢
            
              
            
            ⁢
            the
          
        
      
      
        
          
            reactor
            ⁢
            
              
            
            ⁢
            
              (
              
                m
                2
              
              )
            
          
        
      
    
  




Typically, the fluid pluralized bed reactor SGV ranges from 0.1 m/sec to 2.0 m/sec. Low SGV may result in lack of fluidization while a high SGV may convert a dense phase fluid pluralized bed into a lean phase fluid pluralized bed where the voidage becomes so excessive that it inhibits contact between the catalyst and the reactive gases resulting in significant reaction yield loss.
In one embodiment, the SGV is not greater than 1.5 m/sec and in another not greater than 1.25 m/sec. Preferably, the reactor is maintained at a SGV of 0.3 m/sec to 1 m/sec and more preferably from 0.3 m/sec to 0.5 m/sec.
Particle density is also a significant contributor in the maintenance of fluid pluralized beds. It is calculated by using the following equation.




  
    ρ
    p
  
  =
  
    
      1
      
        
          1
          /
          
            ρ
            s
          
        
        +
        
          V
          p
        
      
    
    ⁡
    
      [
      
        g
        /
        
          cm
          3
        
      
      ]
    
  




    ρs: the true density of the substance constituting the particles (g/cm3)
    VP: the pore volume (cm3/g)
    ρp: the particle density (g/cm3)


A good range for particle density in this reactor is 1.5 g/cm3 to 3 g/cm3.
If the particle density is more than 3 g/cm3, the bed will require extremely high gas flow rates to attain a fluidized condition. Such high gas flow rates will result in low per pass conversions because the volumetric hourly space velocity will become excessive.
If the particle density is less than 1.5 g/cm3, there will be a significant excess of catalyst going off into the lean phase and a small amount in the dense reactive phase. A lot of the catalyst will flow into the upper section and put a heavy burden on the cyclones and diplegs. It is possible that the catalyst layer will swell up at relatively low superficial linear velocities and flood the reactor.
Bed density is another criteria that has to be carefully considered. The bed density of the reactor is defined as the volume fraction of catalyst solids in the reactor. Generally, the catalyst in the dense phase is maintained at a solids volume fraction ranging from 0.25 to 0.6. This fraction translates to 25% to 60% of the volume in the bed is occupied by the solid catalyst particles. In the lean phase, which is above the dense phase zone, the solids void fraction typically ranges from 0.15 to 0.3, which translates into 15 to 30% volume in the bed occupied by the solid catalyst particles. Preferably, the catalyst in the dense phase is maintained at a solids volume fraction ranging from 0.4 to 0.5 and in the lean phase, the catalyst is maintained at a solids volume fraction ranging from 0.2 to 0.25.
Backmixing
Backmixing is a state in a reactor where the contents are well stirred and uniform in composition throughout. Consequently, the exit stream leaving this reactor should have the same composition as the fluid within the reactor.
Backmixing of the catalyst in this application is typically attained by utilizing proper reactor design and a combination of superficial gas velocities, aspect ratio of the catalyst bed and the catalyst particle size and density. The flow gas through the dense phase zone is adequate to keep the catalyst in the dense phase zone with sufficient backmixing. Good backmixing results in a remarkably temperature stable system without hot spots, making for a good control of the reaction. Further, the large gas solid contact area as well as good contacting of the solid-gas phases makes this an efficient system for effecting the catalytic reactions and heat transfer which achieve a low ΔT across the reactor, both radially and axially.
In addition to the proper superficial gas velocity, bed density and particle size and density as mentioned above, in order to achieve the proper level of backmixing in the dense phase bed, the aspect ratio of the catalyst bed should be kept relatively low. According to this invention, the aspect ratio is the ratio of the height of the catalyst bed to the diameter of the catalyst bed. Preferably, the dense phase bed is maintained at a catalyst bed height to diameter ratio of not greater than 10:1, more preferably not greater than 5:1, and most preferably not greater than about 2:1.
DME Fluid Pluralized Bed Reactor
The fluid pluralized bed reactor is configured to ensure that the heat generated within the reactor during the syngas to DME conversion reaction is balanced by the heat needed to bring the feed gases up to the desired temperature either with or without the internal heat transmission tube in the reactor.
The kinetics of the reaction are such that the bulk of the reaction is typically completed within a short distance downstream of the entry of the feed gases. In order to spread the reaction kinetics in a more isothermal fashion, it is convenient and appropriate to be able to introduce the feed gases in controlled quantities at several sections along the reactor. The feed gases in that case serve to also quench the reactor temperatures and bring them somewhat closer to an isothermal mode.
Each gas quench section also generates an environment of backmixing of catalyst solids that effectively distributes the heat generated uniformly throughout the bed. The circulating pattern of the catalyst lifts the catalyst upwards initially, picking up heat of reaction. This catalyst then circulates downwards to meet the fresh feed gases and preheat them quickly to reaction temperatures. The deployment of catalyst solids in this fashion greatly simplifies the kinetics and reduces the amount of overall heat transfer surface area needed to control the process.
Appropriate injection of quenched recycle syngas into different sections along the reactor then creates a series of independent, yet connected, backmix environments for optimizing the reactions and isothermality. In each section, temperature management can further be conducted by controlling the incoming temperature of the make-up syngas—whether they need to be cooled or warmed is a function of the catalyst. In one embodiment of the invention, the fluid pluralized bed reactor is maintained at a temperature ranging from 150° C. to 350° C. (302° F. to 662° F.). Preferably, it is maintained at a temperature of 180° C. to 320° C. (356° F. to 608° F.), more preferably from 200° C. to 280° C. (392° F. to 536° F.).
In one embodiment, the temperature of the fluid pluralized bed is maintained by controlling the temperature of the feed gases entering the reactor through the reactor bottom gas distributor. In another embodiment, the feed gases may be entering the reactor at several different sections. The temperatures of the different incoming feed gas streams may be independently controlled to generate axial isothermality of the reactor. Preferably, the make-up syngas stream flowing into the reactor enters at ambient temperature, 37.8° C. (100.0° F.) while the recycle syngas enters at methanol absorber outlet temperatures ranging from −21° C. to −5° C. (−6° F. to 23° F.).
The treated gas from the methanol absorber contains unconverted syngas. In order to maximize the production of DME, this gas needs to be recycled around the DME synthesis loop. A small amount of the recycle gas is purged in order to remove the accumulating inert gases in the loop.
Distributors
Kinetics in the reactor is conducted by injecting the reactant gases through appropriate distributors. The function of the distributors is to evenly distribute the reactant gases so as to fluidize the catalyst in the reactor in such a way as to maintain sufficient backmix capabilities.
There may be several distributors axially positioned approaching a logarithmic distribution along the fluid pluralized bed reactor. A reactor, for example, may typically be equipped with three distributors, one at the bottom, one at position x and one at position 10×, where x is an arbitrary axial dimension. Each distributor then generates its own fluidization characteristic along with its own backmix envelope.
Each one of the backmix envelopes that has been developed will have its independent kinetics, the commensurate heat release (these reactions are highly exothermic) and volume contraction (the overall DME synthesis reaction converts three moles of CO plus three moles of H2 into one mole of DME and one mole of CO2). The temperature increase within the backmix envelope along with a volumetric decrease of gas creates issues with regards to fluidization and backmix.
It has been discovered that introducing the gas through the distributors and into the reactor must be done under a temperature control algorithm where temperature sensors located in the bed directly above the distributors control the amount and the temperature of the reactant gas. The backmix characteristics of the fluid pluralized bed are normally analyzed by embedding several temperature sensors within the bed and looking at the temperature differences between the sensors. Ideally a good backmix system will exhibit the temperature differences to be approaching zero.
The proposed reactor is equipped with three distributors as shown in FIG. 2. In one embodiment of the reactor, the reactant gases enter at three sections: one at the bottom of the reactor through a distributor and two entries are along the side of the reactor at different heights. These three gas entries through their distributors results in three independent yet linked backmix reactors where the temperature is controlled not only by the temperatures of the feed gases but also by limiting the conversion of the reactants. It is obvious that at gas flow rates will also be a determinating factor of the reaction completion. Because of the highly exothermic nature of the overall reaction, it is sometimes important to control the overall mole conversion of the CO in a range from 25 to 60%. Preferably, conversions are maintained in the range from 30-55% and more preferably from 40-50%.
The fluid pluralized bed reactor is proposed as the ideal device for DME synthesis. Compared with the slurry reactor, the gas-solid mass transfer resistance in a fluid pluralized bed reactor is so small that it can be neglected, and excellent temperature control is also achievable due to the vigorous mixing of catalyst particles in the bed. Almost all of the reactions occur in the dense phase, which contains the catalyst particles, whereas the bubble phase does not contribute significantly to the reaction due to low solid concentration. Concentration gradients are established between the two phases, due to the depletion of reactants and the synthesis of products, inducing the diffusion of products from the dense phase to the bubble phase, and that of the reactants in the opposite direction.
The syngas-to-DME process is highly exothermic. A critical factor for DME reactor design is the management of the heat released by the reactions. The heat released by DME production can generate 2.4 tons of steam per ton of DME, equivalent to an adiabatic temperature rise of about 1,000° C. (1,832° F.) at a complete conversion of syngas with a 1:1 molar ratio of H2:CO in the feed gas. This problem is especially significant as the catalyst of DME synthesis may be deactivated rapidly when the temperature is over 280° C. (536° F.).
Our preliminary theoretical investigation has shown that fluid pluralized bed technology is more efficient for DME synthesis than slurry reactor and fixed bed technologies. When H2/CO molar ratio equals 1.0 in the feed gas, the CO conversion and DME selectivity are 48.5% and 97% in a fluid pluralized bed reactor, compared to the values of 17% and 70% in a slurry reactor under the same conditions, and to 10.7% and 91.9% in a fixed bed reactor under its normal conditions. The superior efficiency of the fluid pluralized bed results from the elimination of diffusional limitations, giving rise to an effectiveness factor very close to one, and also because of the shift of equilibrium to more favorable conditions, such as the product diffusion from the dense phase to the bubble phase.
The sensitivity simulation shows that, the effect of the mass transfer coefficient can be ignored, the optimum H2/CO molar ratio in the make-up syngas is 0.9 to 1.1 while in the feed gas (make-up syngas plus recycle syngas) to the DME reactor is between 0.9 to 1.5. The enhancement of pressure improves DME productivity substantially. High temperature is also favorable for DME synthesis up to a maximum temperature of 285° C. (545° F.), past which it starts dropping gradually.
The direct conversion of the equimolar hydrogen and CO gas mixture in the make-up syngas into DME is an extremely exothermic reaction. In view of this, our reactor has been specially designed to maintain a highly isothermal profile. The DME catalyst is susceptible to rapid coking in case the operating conditions are upset for some unforeseen reason(s), therefore it's essential that the catalyst can be added or replaced with a minimum of difficulty and effort.
Catalyst
This invention uses an attrition resistant bifunctional catalyst where methanol synthesis—water gas shift reaction is one function and the methanol dehydration to DME reaction is the other function.
The procedure to manufacture of this catalyst comprises the following steps:

    
    
        (a) co-precipitation preparation of CuCO3, ZnCO3, CrO2 and La2(CO3)3 using reagents of Cu(NO3)2, Zn(NO3)2, Cr(NO3)3, La(NO3)3 and Na2CO3 at 80° C. to a final pH value of 7, then Al(OH)3 power was added to the co-precipitate,
        (b) preparation of a slurry with the dehydrogenation component HZSM-5,
        (c) addition of silica sol as binder with a particle size less than 15 nm and a pH value of 4-5 into the slurry,
        (d) spray drying the slurry to form microspheres and
        (e) calcination.
    
    


The catalyst comprises of particles having a size ranging from 20 to 300 μm and is optimized for fluid pluralized bed reactor operation.
The following examples provide data for the balancing of the heat generates within the fluid pluralized bed reactor during the syngas to DME synthesis against the heat needed to bring the make-up syngas and recycle gas up to the desired DME reaction temperature, i.e. 260° C. to 280° C. (500° F. to 536° F.) either with or without the internal heat transmission tube in the reactor. Two different absorption solvents for the removal of the acid gases generated in the DME synthesis loop were evaluated. Other options include the use of a water gas shift reactor after the autothermal reformer and the recycle of the whole DME synthesis loop purge gas to the autothermal reformer were also investigated. Finally, the pressure of the two light end columns is changed to three different pressure levels: 224.6 psig (16 bar), 347.9 psig (25 bar) and 463.9 psig (33 bar) in order to increase the condenser temperature. All the examples are given by way of illustration only and not by way of limitation to the present invention.
EXAMPLE 1
36,358 lb/hr of biomass with the composition shown below is fed to a gasifier operated at 1750° F. and 150 psig.










Biomass Proximate Analyses










Comp.
Wt %














Moisture
10.000



Fixed Carbon
11.954



Volatiles
83.400



Ash
4.646



Total
110.000




















Biomass Ultimate Analyses










Comp.
Wt %














Carbon
49.454



Hydrogen
5.755



Nitrogen
0.061



Sulfur
0.024



Oxygen
40.060



Ash
4.646



Total
100.000










An oxidative autothermal reformer operated at 1557° F. and 144 psig is provided for the simultaneous removal of tars, benzene/toluene/xylene components, and for decreasing methane concentration by reforming while optimizing energy efficiency. A syngas with the composition below is obtained:










Syngas from Autothermal Reformer


















Phase
Vapor







Temp., ° F.
1556.8



Pressure, psig
144.0



Flowrate, lbmol/hr
2637.62



H2/CO, Molar Ratio
0.9560







Composition
Mol %







CH4
2.00



CO2
14.80



N2
0.35



H2O
15.29



CO
34.54



H2
33.02










This gas is then cooled down to 108° F. to knock out most of its moisture content before it is compressed by a two-stage compressor with intercooler, aftercooler and water knockout to 710 psig.
The compressed syngas is then passed through an absorber using CH3—O—(CH2CH2O)5—CH3 (C12H26O6) as the absorption solvent to remove 88 mol % of the CO2 in the stream. The rich solvent is regenerated by a simple flash, and no thermal energy is required. The lean gas from the absorber is combined with H2 recovered from the DME synthesis loop purge gas before it is fed to a DME fluid pluralized bed reactor. Part of the heat released by the exothermic reactions (Q=−10.48×106 Btu/hr) is removed by passing a heat transfer medium through a heat transmission tube in the fluid pluralized bed. This make-up syngas having the following composition is introduced through the distributor located at the bottom of the fluid pluralized bed reactor.










Make-up Syngas to DME Reactor


















Phase
Vapor







Temp., ° F.
100.0



Pressure, psig
710.5



Flowrate, lbmol/hr
1766.36



H2/CO, Molar Ratio
1.0095







Composition
Mol %







CH4
2.25



CO2
2.67



N2
0.48



H2O
0.02



CO
47.07



H2
47.51










Due to the efficient backmixing is maintained in the fluid pluralized bed, the fresh make-up syngas is preheated quickly to reaction temperatures. The cooled recycle gas after the purge is served as the two other entries along the side of the reactor at different heights in order to have additional control of the reaction temperatures.
The effluent from the reactor is cooled to 14° F. to condense out most of the water and methanol and about 40% of the DME product. The remaining of the DME product and most of the acid gas (CO2) in the vapor phase is removed by an absorber using methanol as the absorption solvent. The stream descriptions of the condensate from the condenser and rich solvent from methanol absorber are shown below.










Crude DME Product










Condensate from
Rich Solvent from


Stream
the Condenser
Methanol Absorber





Phase
Liquid
Liquid


Temp., ° F.
14.0
35.1


Pressure, psig
710.5
710.5


Flowrate, lbmol/hr
212.44
3426.74


Component Flow, lbmol/hr


CH4
5.94
24.24


CO2
39.35
276.73


N2
0.72
2.17


H2O
12.63
4.82


CO
1.58
4.78


H2
1.26
4.83


CH4O
39.73
2944.28


DME
111.22
164.89









The light end in the stream of condensate from the condenser is removed by light end distillation column 1, and the crude DME is purified by DME distillation column 1:










Fuel Grade DME Product and Recycled Crude Methanol









Stream
DME Product
Methanol & Water





Phase
Liquid
Liquid


Temp., ° F.
96.4
264.4


Pressure, psig
101.3
101.3


Flowrate, lbmol/hr
110.67
52.37


Component Flow, lbmol/hr


CH4
0.00
0.00


CO2
0.01
0.00


N2
0.00
0.00


H2O
0.00
12.63


CO
0.00
0.00


H2
0.00
0.00


CH4O
0.00
39.74


DME
110.67
0.00









The methanol & water stream obtained from the bottom of the distillation column is recycled to the DME fluid pluralized bed reactor.
Similarly, the light end in the stream of rich solvent from the methanol absorber is removed by light end distillation column 2, and the crude DME is purified by DME distillation column 2:










Fuel Grade DME Product and Regenerated Absorption Solvent









Stream
DME Product
Absorption Solvent





Phase
Liquid
Liquid


Temp., ° F.
96.4
260.1


Pressure, psig
101.3
101.3


Flowrate, lbmol/hr
160.70
2953.00


Component Flow, lbmol/hr


CH4
0.00
0.00


CO2
0.00
0.00


N2
0.00
0.00


H2O
0.00
4.82


CO
0.00
0.00


H2
0.00
0.00


CH4O
0.00
2944.28


DME
160.70
3.90









The absorption solvent thus regenerated is recycled to the methanol absorber for reuse. The DME produced (150 tons/day) from these two DME distillation columns has a purity greater than 99.98 mol % which is suitable for fuel grade DME applications.
A simplified process flow block diagram for the production of 150 tons/day of fuel grade DME from biomass is shown in FIG. 3. This process was simulated using Aspen Plus Version 7.1. The corresponding stream description is shown in Table 1.
A detailed process flow block diagram for the DME synthesis, gas cooling, acid gas removal and DME purification sections is shown in FIG. 4 and its corresponding stream description in Table 2.







TABLE 1





150 TPD DME PLANT

















Stream No.
















S1
S2
S3
S4
S5
S6
S7
S8





Phase
V
V
V
V
V
L
V
L


Temperature ° F.
800
350
1750
350
1557
108
108
108


Pressure, psig
150
150
150
150
144
135
135
284


Flowrate, #mol/hr
238.5294
394.4257
2423.1599
29.0630
2637.6238
386.3010
2251.3236
12.4345


Composition, mol %


CH4
0.0000
0.0000
6.5902
0.0000
1.9992
0.0000
2.3423
0.0000


CO2
0.0000
0.0000
16.3313
0.0000
14.7950
0.0000
17.3311
0.0000


N2
0.0000
2.0000
0.3540
2.0000
0.3466
0.0000
0.4060
0.0000


O2
0.0000
98.0000
0.0000
98.0000
0.0000
0.0000
0.0000
0.0000


H2O
100.0000
0.0000
18.4786
0.0000
15.2890
100.0000
0.7563
100.0000


CO
0.0000
0.0000
32.9540
0.0000
34.5381
0.0000
40.4644
0.0000


H2
0.0000
0.0000
25.2792
0.0000
33.0191
0.0000
38.6847
0.0000


H2S
0.0000
0.0000
99 PPM
0.0000
91 PPM
0.0000
106 PPM
0.0000


COS
0.0000
0.0000
 4 PPM
0.0000
 4 PPM
0.0000
 5 PPM
0.0000


NH3
0.0000
0.0000
23 PPM
0.0000
35 PPM
0.0000
 41 PPM
0.0000


CH4O
0.0000
0.0000
0.0000
0.0000
0.0000
0.0000
0.0000
0.0000


DME
0.0000
0.0000
0.0000
0.0000
0.0000
0.0000
0.0000
0.0000












Stream No.
















S9
S10
S11
S12
S13
S14
S15
S16





Phase
V
V
V
V
V
V
V
V


Temperature ° F.
108
114
108
100
48
23
23
23


Pressure, psig
711
0
710
710
710
710
710
710


Flowrate, #mol/hr
2238.8948
487.2304
1751.4422
1766.3551
5302.6448
3585.0645
48.7748
3536.2897


Composition, mol %


CH4
2.3553
2.6506
2.2734
2.2543
13.9408
19.7783
19.7783
19.7783


CO2
17.4271
70.3783
2.6969
2.6741
4.8734
5.9716
5.9716
5.9716


N2
0.4083
0.1511
0.4799
0.4758
7.6955
11.3021
11.3021
11.3021


O2
0.0000
0.0000
0.0000
0.0000
0.0000
0.0000
0.0000
0.0000


H2O
0.2054
0.8406
197 PPM
196 PPM
 66 PPM
 1 PPM
 1 PPM
1 PPM


CO
40.6891
16.3485
47.4655
47.0649
30.3942
22.0670
22.0670
22.0670


H2
38.8996
9.5661
47.0646
47.5114
43.0062
40.7562
40.7562
40.7562


H2S
107 PPM
452 PPM
0.0000
0.0000
0.0000
0.0000
0.0000
0.0000


COS
 5 PPM
 22 PPM
0.0000
0.0000
0.0000
0.0000
0.0000
0.0000


NH3
 41 PPM
173 PPM
0.0000
0.0000
0.0000
0.0000
0.0000
0.0000


CH4O
0.0000
0.0000
0.0000
0.0000
598 PPM
897 PPM
897 PPM
0.0000


DME
0.0000
0.0000
0.0000
0.0000
234 PPM
351 PPM
351 PPM
0.0000












Stream No.

















S17
S18
S19
S20
S21
S22
S23
S24
S25





Phase
V
V
V
L
V + L
V
L
L
L


Temperature ° F.
23
23
−56
100
33
−51
260
14
96


Pressure, psig
710
710
101
710
101
101
101
710
101


Flowrate, #mol/hr
33.8657
14.9091
313.0383
52.7272
3639.1756
49.3894
2953.0180
1.9730
271.3786


Composition, mol %


CH4
28.4855
0.0000
7.7423
0.0000
0.8293
12.0315
0.0000
0.0000
0.0000


CO2
8.6005
0.0000
88.4006
0.0000
8.6854
79.6639
0.0000
0.0000
2 PPM


N2
16.2778
0.0000
0.6918
0.0000
792 PPM
1.4569
0.0000
0.0000
0.0000


O2
0.0000
0.0000
0.0000
0.0000
0.0000
0.0000
0.0000
0.0000
0.0000


H2O
 1 PPM
0.0000
0.0000
23.9567
0.4796
0.0000
0.1633
0.0000
12 PPM 


CO
31.7818
0.0000
1.5272
0.0000
0.1749
3.2056
0.0000
0.0000
0.0000


H2
14.6747
100.0000
1.5429
0.0000
0.1673
2.5451
0.0000
0.0000
0.0000


H2S
0.0000
0.0000
0.0000
0.0000
0.0000
0.0000
0.0000
0.0000
0.0000


COS
0.0000
0.0000
0.0000
0.0000
0.0000
0.0000
0.0000
0.0000
0.0000


NH3
0.0000
0.0000
0.0000
0.0000
0.0000
0.0000
0.0000
0.0000
0.0000


CH4O
0.1292
0.0000
0.0000
75.3623
81.9971
0.0000
99.7043
100.0000
4 PPM


DME
505 PPM
0.0000
953 PPM
0.6810
7.5873
1.0972
0.1324
0.0000
99.9982
















TABLE 2





150 TPD DME PLANT

















Stream No.


















S1
S2
S3
S4
S5
S6
S7
S8
S9
S10





Phase
V
V
V
L
V
V
V
V
L
V


Temperature ° F.
100
48
500
14
14
23
23
23
116
−51


Pressure, psig
710
710
710
710
710
710
710
710
101
101


Flowrate,
1766.3551
5302.6448
4270.1830
212.4433
4057.6448
3585.0645
48.7748
3536.2897
163.0440
49.3894


lbmol/hr


Composition,


mol %


CH4
2.2554
13.9408
17.3117
2.7974
18.0716
19.7783
19.7783
19.7783
0.0000
12.0315


CO2
2.6741
4.8734
12.4154
18.5221
12.0957
5.9716
5.9716
5.9716
3 PPM
79.6639


N2
0.4758
7.6955
9.5563
0.3387
10.0389
11.3021
11.3021
11.3021
0.0000
1.4569


H2O
196 PPM
66 PPM
0.3177
5.9462
230 PPM
 1 PPM
 1 PPM
 1 PPM
7.7478
0.0000


CO
47.0649
30.3942
18.6755
0.7453
19.6142
22.0670
22.0670
22.0670
0.0000
3.2056


H2
47.5114
43.0062
34.3597
0.5917
36.1277
40.7562
40.7562
40.7562
0.0000
2.5451


CH4O
0.0000
0.0000
0.9597
18.7045
306 PPM
897 PPM
897 PPM
897 PPM
24.3712
0.0000


DME
0.0000
0.0000
6.4040
52.3540
3.9983
351 PPM
351 PPM
351 PPM
67.8806
1.0972












Stream No.


















S11
S12
S13
S14
S15
S16
S17
S18
S19
S20





Phase
L
L
L
L
V
L
L
L
L
L


Temperature ° F.
259
96
35
234
−56
96
260
260
14
96


Pressure, psig
101
101
710
101
101
101
101
101
710
101


Flowrate,
52.7272
110.6786
3426.7323
3113.6980
313.0383
160.7000
2953.0180
0.9288
1.9730
271.3786


lbmol/hr


Composition,


mol %


CH4
0.0000
0.0000
0.7073
0.0000
7.7423
0.0000
0.0000
0.0000
0.0000
0.0000


CO2
0.0000
 5 PPM
8.0756
0.0000
88.4006
0.0000
0.0000
0.0000
0.0000
2 PPM


N2
0.0000
0.0000
632 PPM
0.0000
0.6918
0.0000
0.0000
0.0000
0.0000
0.0000


H2O
23.9567
29 PPM
0.1407
0.1549
0.0000
0.0000
0.1633
100.0000
0.0000
12 PPM 


CO
0.0000
0.0000
0.1395
0.0000
1.5272
0.0000
0.0000
0.0000
0.0000
0.0000


H2
0.0000
0.0000
0.1409
0.0000
1.5429
0.0000
0.0000
0.0000
0.0000
0.0000


CH4O
75.3623
0.0000
85.9209
94.5590
0.0000
7 PPM
99.7043
0.0000
100.0000
4 PPM


DME
0.6810
99.9966
4.8119
5.2861
953 PPM
99.9992
0.1324
0.0000
0.0000
99.9982









EXAMPLE 2
Same as Example 1 except that the internal heat transmission tube in the fluid pluralized bed reactor is removed. A higher recycle rate (from 3536 μmol/hr to 6363 μmol/hr) in the DME synthesis loop is required to maintain the same effluent temperature of the DME reactor. This higher recycle rate is obtained by injecting more steam to the gasifier (from 239 μmol/hr to 378 μmol/hr). The H2/CO molar ratio in the feed syngas to the DME reactor also increases from 1.4149 to 2.8088 due to the water gas shift reaction:










Feed Gases to DME Reactor











Feed Gas to DME Reactor
Example 1
Example 2







Phase
Vapor
Vapor



Temp., ° F.
47.5
39.5



Pressure, psig
710.5
710.5



Flowrate, lbmol/hr
5302.72
8183.06



H2/CO Molar Ratio
1.4149
2.8088



Composition, Mol %



CH4
13.94
13.46



CO2
4.87
4.91



N2
7.70
5.89



H2O
0.00
0.00



CO
30.40
19.86



H2
43.01
55.79



CH4O
0.06
0.07



DME
0.02
0.02










The higher recycle rate in the DME synthesis loop also reduces the partial pressure of the DME product in the phase separator, only 0.715 mol % of the DME produced is condensed out in the DME reactor effluent condenser compared to 40.672 mol % in Example 1:










Condensate from the Phase Separator











Feed Gas to DME Reactor
Example 1
Example 2







Phase
Liquid
Liquid



Temp., ° F.
14.0
14.0



Pressure, psig
710.5
710.5



Flowrate, lbmol/hr
212.44
79.86



Component Flow, lbmol/hr



CH4
5.94
0.07



CO2
39.35
2.02



N2
0.72
0.00



H2O
12.63
37.99



CO
1.58
0.01



H2
1.26
0.02



CH4O
39.74
37.78



DME
111.22
1.97










Therefore, a higher methanol solvent flowrate is needed in Example 2 for the methanol absorber to absorb the additional DME content in the recycle syngas, 3935 μmol/hr as compared to 2953 μmol/hr in Example 1.
The DME and most of the CH4O in the condensate from the condenser can be recovered by a distillation column and are recycled back to the DME reactor:










Recovered DME and CH4O Recycled to DME Reactor










Phase
Vapor














Temp., ° F.
100.0



Pressure, psig
710.5



Flowrate, lbmol/hr
38.00



Component Flow, lbmol/hr



CH4
0.07



CO2
2.02



N2
0.00



H2O
2.05



CO
0.01



H2
0.02



CH4O
31.86



DME
1.97










The light end in the stream of rich solvent from the methanol absorber bottom is separated by a light end distillation column, and the crude DME is purified by the DME distillation column:










Fuel Grade DME Product & Regenerated Absorption Solvent









Stream
DME Product
Absorption Solvent





Phase
Liquid
Liquid


Temp., ° F.
96.5
260.3


Pressure, psig
101.3
101.3


Flowrate, lbmol/hr
271.40
3935.44


Component Flow, lbmol/hr


CH4
0.00
0.00


CO2
0.01
0.00


N2
0.00
0.00


H2O
0.00
3.46


CO
0.00
0.00


H2
0.00
0.00


CH4O
0.00
3928.38


DME
271.39
3.60









In this example, all the DME is produced from a single DME distillation column. The absorption solvent thus regenerated at the column bottom is recycled to the methanol absorber for reuse.
EXAMPLE 3
Same as Example 2 except that the absorption solvent in the methanol absorber is replaced by CH3—O(—CH2—CH2—O)5—CH3 (C12H26O6). Due to the higher solubility of the syngas in C12H26O6, more biomass, oxygen and steam are required to produce the same amount of DME product:










Rich Solvent from Methanol/C12H26O6 Absorber










Rich Solvent from Methanol/





C12H26O6 Absorber
Example 1
Example 2
Example 3





Phase
Liquid
Liquid
Liquid


Temp., ° F.
35.1
34.3
32.0


Pressure, psig
710.5
710.5
710.5


Flowrate, lbmol/hr
3426.74
4533.33
2358.64


Flowrate, #/hr
114796
151721
479062


Wt Ratio, wt %
23.96
31.67
100.00


Component Flow, lbmol/hr


CH4
24.24
27.92
38.16


CO2
276.73
283.49
280.64


N2
2.17
1.93
5.03


H2O
4.82
3.46
0.71


CO
4.78
3.83
17.11


H2
4.83
9.08
42.24


CH4O
2944.28
3928.38
3.17


DME
164.89
275.24
271.58


C6H26O6
0.00
0.00
1700.00



















Feedstocks










Feedstocks
Example 1
Example 2
Example 3













Biomass (10 wt % moisture), TPD
438
445
464


Oxygen (98 mol % pure), TPD
162
164
170


Steam, TPD
52
82
100









The absorption solvent C12H26O6 also has much higher molecular weight and boiling point than CH4O which means higher energy is required to heat, to cool and to pump the absorption solvent C12H26O6:










Pure Absorption Solvent Physical Properties











Absorption Solvent
MW
TF, ° F.
TB, ° F.
ΔHv, Btu/#mol














CH4O
32.04
−143.82
148.46
15108


C12H26O6
266.33
11.21
647.51
29585



















Energy Requirement










Energy Requirement
Example 1
Example 2
Example 3





Reboiler Heat Duty, Btu/hr
31.88 × 106
48.72 × 106
230.07 × 106


To Pump from 101 psig to
157.16
209.57
855.86


710 psig, hp


To Cool from Reboiler Temp.
20.79 × 106
27.74 × 106
212.60 × 106


to 14° F., Btu/hr









EXAMPLE 4
Same as Example 2 except that 18.75 mol % of the cooled syngas from the autothermal reformer is passed through a water gas shift (WGS) reactor. The effluent from the WGS reactor is then combined with the remaining 81.25 mol % of the syngas in order to have a similar compressed syngas composition as in Example 2:










Compressed Syngas Composition











Compressed Syngas Composition
Example 2
Example 4







Phase
Vapor
Vapor



Temp., ° F.
108.0
108.0



Pressure, psig
710.5
710.5



Component Flow, lbmol/hr



CH4
56.26
52.76



CO2
445.19
444.35



N2
9.23
9.24



H2O
4.79
4.78



CO
873.48
871.69



H2
930.05
930.94










Meanwhile the steam feed to the gasifier is greatly reduced:










Steam Feed to Gasifier











Feedstocks
Example 2
Example 4







Steam, TPD
82
46










EXAMPLE 5
Same as Example 1 except that the PSA unit for the H2 recovery from the DME synthesis loop purge gas is eliminated and the whole purge gas stream is recycled to the autothermal reformer. The methane content in the purge gas is reformed to produce more H2 and CO in the autothermal reformer, and less biomass, oxygen and steam are required to produce the same amount of DME product.










Feedstocks











Feedstocks
Example 1
Example 5















Biomass (10 wt % moisture), TPD
438
429



Oxygen (98 mol % pure), TPD
162
161



Steam, TPD
52
50










The H2/CO molar ratio in the feed syngas to the DME reactor also reduces from 1.4149 to 1.0392 due to the elimination of the PSA unit for the H2 recovery from the purge gas. Meanwhile the N2 concentration increases from 7.70 to 19.12 mol %.










Feed Gases to DME Reactor











Feed Gas to DME Reactor
Example 1
Example 5







Phase
Vapor
Vapor



Temp., ° F.
47.5
46.6



Pressure, psig
710.5
710.5



Flowrate, lbmol/hr
5302.72
5308.02



H2/CO Molar Ratio
1.4149
1.0392



Composition, Mol %



CH4
13.94
14.00



CO2
4.87
4.95



N2
7.70
19.12



H2O
0.00
0.00



CO
30.40
30.33



H2
43.01
31.52



CH4O
0.06
0.06



DME
0.02
0.02










EXAMPLE 6
Same as Example 1 except that the pressure of the two light end columns is changed to 224.6 psig, 347.9 psig and 463.9 psig in order to increase the condenser temperature. The pressures and the resulting molar reflux ratios, temperatures and heat duties of the condensers and reboilers for these two light end columns are shown below:










LIGHT END DISTILLATION COLUMN #1










Condenser
Reboiler












Pressure,
Molar
Tempera-
Heat
Tempera-
Heat


Bar
Reflux
ture,
Duty,
ture,
Duty,


(psig)
Ratio
° F.
106 Btu/hr
° F.
106 Btu/hr















 8.0 (101.3)
1.0
−50.92
−0.38
116.14
0.99


16.5 (224.6)
2.0
−18.15
−0.65
175.05
1.55


25.0 (347.9)
3.0
−0.66
−1.08
214.27
1.57


33.0 (463.9)
4.0
18.13
−1.05
243.40
2.30



















LIGHT END DISTILLATION COLUMN #2










Condenser
Reboiler












Pressure,
Molar
Tempera-
Heat
Tempera-
Heat


Bar
Reflux
ture,
Duty,
ture,
Duty,


(psig)
Ratio
° F.
106 Btu/hr
° F.
106 Btu/hr















 8.0 (101.3)
1.0
−55.54
−2.05
233.77
20.32


16.5 (224.6)
2.0
−20.46
−3.65
293.31
28.04


25.0 (347.9)
2.5
1.58
−4.11
330.61
31.06


33.0 (463.9)
3.0
19.01
−4.43
357.64
36.31









EXAMPLE 7
Same as Example 5 except that the Australian brown coal (sometimes the brown coal is called lignite) is used to replace the biomass feed. Brown coal is a relatively soft material that has a heating value only about one-quarter of that of black coal. The typical composition of an Australian brown coal is shown below:










Brown Coal Proximate Analyses










Comp.
Wt %














Moisture
0.00



Fixed Carbon
25.87



Volatiles
70.13



Ash
4.00



Total
100.00




















Brown Coal Ultimate Analyses










Comp.
Wt %














Carbon
66.0



Hydrogen
4.5



Nitrogen
0.5



Sulfur
1.0



Oxygen
24.0



Ash
4.0



Total
100.0










Due to the higher carbon and lower oxygen contents in the brown coal ultimate analyses, less brown coal and more steam are required to produce the same amount of DME product:










Feedstocks











Feedstocks
Example 5
Example 7















Dry Solid Feed, TPD
386
267



Oxygen (98 mol % pure), TPD
161
136



Steam, TPD
50
167










EXAMPLE 8
Same as Example 5 except that the natural gas with the following typical composition is used to replace the biomass feed. Since natural gas is rich in hydrogen, the syngas thus produced has a hydrogen to carbon monoxide molar ratio of 3 to 4. In order to have a molar ratio of about 1, a lot of carbon dioxide is needed to shift the following reaction to the left:

CO+H2OCO2+H2 











Typical Natural Gas Composition










Component
Mole %














Methane
95.2



Ethane
2.5



Nitrogen
1.3



Carbon Dioxide
0.7



Propane Plus
0.3



Total
100.0



Specific Gravity
0.58



Gross Heating Value
37.8 MJ/m3 (dry)










Due to the higher heating value of natural gas (23276 Btu/lb) as compared to brown coal (11130 Btu/lb) or biomass (8225 Btu/lb), much less of natural gas is required to produce the same amount of DME product.










Feedstocks










Feedstocks
Example 5
Example 7
Example 8













Fuel, TPD
386
267
133


Oxygen (98 mol % pure), TPD
161
136
179


Steam, TPD
50
167
—


Carbon Dioxide, TPD
—
—
272









EXAMPLE 9
Same as Example 8 except that the high carbon dioxide content streams of the acid gas removal unit off gas and the light end distillation column overhead are used to replace the pure carbon dioxide feed. Meanwhile additional hydrogen and carbon monoxide in these two streams are also recovered, thus less natural gas feed is required. In addition, much less carbon dioxide (the greenhouse gas) is rejected to the environment.










Feedstocks & Carbon Dioxide Released to the Environment












Exam-
Exam-
Exam-
Exam-


Feedstocks
ple 5
ple 7
ple 8
ple 9














Fuel, TPD
386
267
133
126


Oxygen (98 mol % pure),
161
136
179
171


TPD


Steam, TPD
50
167
—
—


Carbon Dioxide, TPD
—
—
272
276


Addition Hydrogen & Carbon
—
—
—
23


Monoxide Recovered, TPD


Carbon Dioxide (Greenhouse
413
361
339
51


Gas) Released to the


Environment for the


Production of DME, TPD


Net Carbon Dioxide Released
0
648
626
338


to the Environment after


Combustion of DME, TPD









Thus using the recycle of the carbon dioxide rich streams to replace pure carbon dioxide feed stream in the production of DME from natural gas releases the minimum amount of greenhouse gas to the environment. It should be noted that the combustion of DME also releases carbon dioxide (287 TPD). However, the amount of carbon dioxide released during the production and the combustion of DME is balanced exactly the amount of carbon dioxide consumed during the lifecycle of the biomass used to produce DME. Therefore, use of DME made from biomass as an energy source causes no net increase in the carbon dioxide levels of the atmosphere.
EXAMPLE 10
Same as Example 1 except that the two DME distillation columns are replaced by a single light and DME distillation column. The regenerated solvents in both cases are first cooled by the two cold streams fed to the light end columns before the final cooling to the absorber inlet temperature. Since the regenerated solvent in Example 10 contains about 1.25 wt % of DME and gives a higher partial pressure of DME in the treated recycle gas, thus a lower temperature such as −20° F. is required for the absorber inlet temperature.










Energy Requirement









Heat Duties
Example 1
Example 10












Regenerated Solvent Flowrate, lbmol/hr
2988
1715


Absorber Inlet Temp., ° F.
14
−20


Total Condenser Heat Duties, 106 Btu/hr
−23.24
−8.97


Total Reboiler Heat Duties, 106 Btu/hr
31.40
13.32


Cool Reactor Effluent from 100° F. to
−4.83
−7.27


Absorber Inlet Temperature, 106 Btu/hr


Cool Regenerated Solvent from the Light End
−8.39
−2.46


Heat Exchanger Outlet Temperature to the


Absorber Inlet Temperature, 106 Btu/hr









As can be seen above, there are clear advantages to use the methanol/DME mixture which contains about 1.25 wt % of DME instead of pure methanol as the scrubbing solvent in terms of saving both refrigerating and heating energy.

## Claims

1. A process for making DME product in a fluid pluralized bed reactor, comprising: flowing a syngas containing hydrogen and carbon monoxide through a fluid pluralized bed reactor at a superficial gas velocity ranges from 0.1 msec to 2.0 msec; contacting the syngas with a bifunctional catalyst in a dense phase fluid pluralized bed reactor having a catalyst bed height to diameter ratio of not greater than 10:1 so as to form DME product; and separating the catalyst from the product gas in the disengaging zone via one or more cyclones equipped with diplegs.

2. The process of claim 1, wherein about 75% of the hydrogen content in the purge stream from the DME synthesis loop is recovered via a PSA unit, which is then combined with the fresh make-up syngas.

3. The process of claim 1, wherein the vent from the PSA unit may be fed to the autothermal reformer, thus the methane content in the vent stream is reformed to produce additional hydrogen and carbon monoxide, and the hydrogen and carbon monoxide in the vent stream are conserved.

4. The process of claim 1, wherein the PSA unit for the hydrogen recovery from the DME synthesis loop purge gas is eliminated and the whole purge gas stream is recycled to the autothermal reformer to reform the methane content and conserve the hydrogen and carbon monoxide content in the purge gas stream.

5. The process of claim 1, wherein the feed gases are entering the reactor at several different sections and the flowrates of the different incoming feed gas streams are independently controlled to generate axial isothermality of the reactor.

6. The process of claim 1, wherein the make-up syngas stream flowing into the reactor enters at ambient temperature, 37.8° C., while the recycle syngas enters at DME synthesis loop absorber outlet temperatures ranging from −21° C. to −5° C.

7. The process of claim 1, wherein the positions of the distributors in the fluid pluralized bed reactor are optimally approaching a logarithmic distribution.

8. The process of claim 1, wherein the feed gases flowing through the fluid pluralized bed reactor further contains about 2 mol % to 5 mol % of carbon dioxide to maintain the high activity of the catalyst.

9. The process of claim 1, wherein the bifunctional catalyst has a particle size ranging from 20 to 300 μm.

10. The process of claim 1, wherein the fluid pluralized bed reactor temperature is controlled either with or without the internal heat transmission tube in the reactor.

11. The process of claim 1, wherein the catalyst in the dense phase is maintained at a solid volume fraction ranging from 0.4 to 0.5 and in the lean phase, the catalyst is maintained at a solid volume fraction ranging from 0.2 to 0.25.

12. The process of claim 1, wherein the catalyst that is separated via the cyclones is returned to the dense phase fluid pluralized bed through the diplegs for reuse.

13. The process of claim 1, wherein the fluid pluralized bed reactor is maintained at a temperature ranging from 150° C. to 350° C.

14. The process of claim 1, wherein the fluid pluralized bed reactor is maintained at a pressure ranging from 3 MPa to 8 MPa.

15. The process of claim 1, wherein a water gas shift reactor is added after the autothermal reformer to shift part of the syngas produced in order to reduce the flowrate of steam fed to the gasifier and keep the same hydrogen to carbon monoxide molar ratio.