JPH0340013B2 - - Google Patents

Description

[Detailed description of the invention]

The present invention relates to a process for producing olefins from methanol, and in particular to a process for the catalytic conversion of methanol to ethylene. The increasing demand for olefin feedstocks has periodically caused petrochemical feedstock shortages due to limited availability of petroleum of suitable quality traditionally used as feedstocks or limited naphtha cracking equipment. There is. Sources of ethylene from other non-petroleum sources are one means of keeping pace with ethylene and other olefin requirements. Methanol and/or dimethyl ether obtained from coal, natural gas or biological sources are ZSM-
It is known that more complex hydrocarbons such as olefins and aromatics can be converted through the use of catalytic zeolites such as 5 zeolite. Ethylene is one type of olefinic hydrocarbon obtained by this catalytic conversion method. This reaction is highly exothermic and the initially formed olefins are likely to undergo further reactions to form aromatic hydrocarbons useful in the production of motor vehicle gasoline. There is much patent literature on the conversion of methanol and/or dimethyl ether to light olefins. Thus, a process for producing olefins by catalytic conversion from aliphatic ethers using HZSM-5 zeolite is described in US Pat. No. 3,894,106. U.S. Pat. No. 3,979,472 describes the conversion reaction of lower alcohols and their ethers with a mixture of antimony oxide and ZM-5 zeolite to produce a mixture of ethylene, propylene and mononuclear aromatics. U.S. Pat. No. 4,025,572 discloses that ZSM-5 can be mixed with an inert diluent to improve ethylene selectivity, while U.S. Pat. I'm wearing it. Ethylene selectivity can be achieved by using only large crystalline ZSM-5 zeolites of at least about 1 micron (U.S. Pat. No. 4,025,571) or by mixing these zeolites with additive metals (U.S. Pat. No. 4,148,835).
will also be improved. Better selectivity can also be obtained by dispersing amorphous silica within the crystal structure of the zeolite catalyst (US Pat. Nos. 4,060,568 and 4,100,219). Although these methods have been particularly successful and very effective in converting lower aliphatic alcohols to olefinic hydrocarbons, the conversion reactions have been found to be exothermic to varying degrees depending on the particular reactants.
For example, the amount of heat generated in the conversion of a lower alcohol to a hydrocarbon product can be estimated to be in the following range.

[Table] Although it is desirable if the reaction is exothermic because it eliminates the need for an external heat source, the large amount of heat generated requires a substantial investment in significant internal cooling equipment after a complex reaction. From the table above it can be seen that the conversion of methanol is significantly exothermic at this point, with methanol being less exothermic. Additionally, due to the inherent properties and efficiency of crystalline aluminosilicate zeolite catalysts, the reaction of methanol is self-promoting and creates significant hot spots in the catalyst bed to complete the reaction, whereas ethanol is less so. This extremely exothermic reaction in an adiabatic catalyst bed reactor increases the rate of catalyst aging and causes thermal catalyst damage. Furthermore, this high temperature causes an undesirable product distribution. Therefore, it is important that the conversion of methanol to useful products be sufficiently thermally dispersed so that the temperature of any part of the catalyst bed is kept within predetermined limits. Furthermore, it is generally good industrial practice to use reactor capacity and ancillary equipment more effectively for reactant conversion at higher pressures. However, high pressures using charged methanol tend to produce large amounts of 1,2,4,5-tetramethylbenzene (durene), an undesirable by-product, while low pressures, for example below 450 kPa, are good for producing light olefins. Various methods have been used to control the heat generated in this process. For example, U.S. Pat.
(Operating conditions selected to suppress it to 25%) and
See No. 4238631 (rising reactor and dense fluidized catalyst bed). US Pat. No. 4,035,430 describes a method for controlling the heat generated when converting methanol to gasoline boiling range products. After converting the feedstock to an equilibrium mixture in a dehydrating catalyst bed, it is passed through a series of zeolite catalyst beds of increasing size with interstage cooling with methanol, dimethyl ether, and/or light hydrocarbons to remove the heat of reaction. . The temperature rise in any bed does not exceed approximately 28°C and the total temperature rise does not exceed approximately 110°C. There is no detailed description of how the relative size of each catalyst bed affects the thermal stability of the reaction. Furthermore, there is no mention of the success of converting the alcohol charge primarily into olefinic hydrocarbons rather than aromatics. In US Pat. No. 2,319,620, multiple fixed beds of catalysts are used for endothermic catalytic cracking of hydrocarbons and for periodic exothermic regeneration of coking catalysts. The hydrocarbon reactant is divided into as many streams as there are fixed catalyst beds, with each stream passing through only one bed. Reactant flow through each bed may be axial or radially inward. US Patent No.
No. 2475855 also uses a series of fixed catalyst beds for endothermic cracking of hydrocarbons. This hydrocarbon stream flows continuously in a radial fashion through a series of beds. The radial flow may be inward or outward. The coked catalyst is regenerated in situ and the heat generated during regeneration is absorbed within each bed and released during endothermic decomposition of the hydrocarbons. Since heat storage results in an increase in the temperature of the reactants passing through the bed, an additional amount of hydrocarbon liquid is introduced between each bed through the injection nozzle to lower the temperature of the mixture sent to the next bed. The size of the catalyst bed is increased in the flow direction to account for the introduction of additional reactants between the beds. This patent does not suggest that the method described is applicable to exothermic reactions or that the radial flow configuration used has a lower pressure drop than other flow configurations. Although the use of high pressures in the conversion of alcohols to olefins usually provides efficient use of processing equipment, low pressures produce significantly higher amounts of light olefins in the product. Therefore, it is desirable to limit the pressure loss of the device so that low charging pressures can be used. Furthermore, although it has been observed that olefin selectivity can only be improved by partial conversion of the feed alcohol ether, this approach also incurs significant economic losses since, for example, methanol costs require recovery and recycling of unreacted reactants. It is also often recognized that The present invention combines methanol alone or its corresponding ether in the presence of a zeolite catalyst by placing the zeolite catalyst in a series of beds and cooling the effluent from each bed between stages to limit the temperature rise of each bed. This is based on the observation that the heat generated in the conversion reaction of the mixture can be effectively removed and stable operation can be achieved.
It has also been discovered that the pressure drop across each catalyst can be minimized by passing the reaction mixture radially through each bed. It has also been discovered that significant economic improvements in operation can be made by recycling unconverted alcohol to the dehydration step and unconverted ether to the conversion step. In accordance with the present invention, a methanol-containing feed is contacted with a dehydration catalyst at an elevated temperature to produce an ether-rich product and the product is characterized in that: (i) a pore size greater than 5 angstroms, a silica to alumina molar ratio of at least 12; with a constraint index of 12 or
or (ii) alternating cooling zones with catalyst beds consisting of any crystalline aluminosilicate zeolite with a predominant pore size of 6 angstroms or less and a pore size defined by eight-membered rings with oxygen atoms. The reaction mixture is brought into contact with a catalyst under heating several times to achieve a predetermined conversion rate to olefinic hydrocarbons, and the temperature reduction of the reaction mixture by passing through each cooling zone is equal to the temperature in the catalyst bed just before that. and the temperature rise of the reaction mixture through each catalyst bed shall be no greater than the sensitivity parameter for the conversion of methanol to ethylene, and the number of catalyst beds shall be equal to the total adiabatic temperature rise for a given conversion reaction. There is provided a method for converting methanol to olefinic hydrocarbon products such that the ratio of methanol to the sensitivity parameter is at least equal to the ratio of . The ether-rich product may be passed axially or radially through the zeolite catalyst bed, and the zeolite catalyst bed may be a fixed bed or a moving bed. In the case of a moving catalyst bed, the reaction mixture stream can be cocurrent or countercurrent with the catalyst stream. The olefinic hydrocarbon product of this process always contains some unreacted methanol and/or ether. This unreacted material can be suitably separated from the product mixture and recycled to the operation. The methanol is returned to the dehydration step and the ether to the conversion step. According to the invention, methanol alone or together with its corresponding ether reacts with olefinic hydrocarbons, in particular in a reactor with specifications that allow effective regulation of the heat released during the conversion of the ether-rich intermediate to olefinic hydrocarbons. Specifically, it is converted to ethylene. The reactants may consist of methanol and dimethyl ether, and the olefinic product may be primarily ethylene. Although the process of the present invention involves the conversion of methanol to olefinic products, the feed stream to the operation may also contain small amounts of other alcohols, such as ethanol, propanol, and isopropanol. In the first step of the process, methanol is contacted with a dehydration catalyst to produce a water- and ether-rich intermediate product. The dehydration catalyst can be any catalyst that provides intermolecular dehydration of the alcohol reactant to an ether-rich product having a higher carbon to oxygen ratio than the feed. Dehydration reactions that can occur include those that produce simple ethers and mixed ethers such as dimethyl ether and diethyl ether. These intermediates are produced by intermolecular dehydration of the corresponding alcohol reactants, and all of these condensations are exothermic reactions. It is generally known that these dehydration reactions use alumina compositions such as gamma alumina, but
Other acidic catalysts known in the art are also very effective for dehydration. It will be appreciated that intramolecular dehydration was not possible using the methanol charge and the dehydration reaction proceeded exothermically only to form dimethyl ether. The method of the invention consists of two successive catalyst contacting steps,
Both processes generate heat. The exotherm in the first step is limited by the inhibition of conversion of methanol to a near-equilibrium mixture of dimethyl ether, methanol, and water. The conversion product or first stage effluent is at a temperature of about 310 to about 400°C due to the exotherm, which is heated to a temperature of about 270 to about 400°C by indirect heat exchange with the circulating heat exchange fluid according to the second stage zeolite catalyst. Well regulated to a temperature of 430℃. For example, the heat exchange fluid may be water or methanol passed through the first step. A second step catalytic conversion reaction converts the first step effluent consisting of methanol, dimethyl ether and water to an olefin-rich product. The operation is a highly exothermic reaction and some crystalline zeolites, especially ZSM
- It occurs rapidly in the presence of type 5 crystalline zeolite and small-pore crystalline zeolite. Generally, the ZSM-5 type zeolite used in the present invention is a crystalline zeolite with a silica to alumina ratio of 12 or greater and a constraint index (CI) of about 1 to about 12. These zeolites and their use as catalysts for the conversion of lower aliphatic alcohols are described in the above-mentioned U.S. patents, particularly Nos. 3,894,106 and 4,025,571;
4058576 and 4148835. Preferred zeolites are ZSM-5 type zeolites,
For example, ZSM-5, ZSM-11, ZSM-12, ZSM-
23, ZSM-35 and ZSM-38, and ZSM-5
is particularly preferred. ZSM-5 is described in U.S. Pat. No. 3,702,886, and ZSM-11 is described in U.S. Pat. No. 3,709,979.
ZSM-12 is covered by U.S. Patent No. 3832449, ZSM-23 is covered by U.S. Patent No. 4076842, and ZSM-35 is covered by U.S. Patent No.
4016245, and ZSM-38 is U.S. Patent No.
Each is described in No. 4046859. Particularly preferred catalysts are large crystals, i.e., U.S. Pat.
At least one as described in Nos. 4025571 and 4148835
It is made of ZSM-5 type zeolite with a crystal diameter of microns. Other particularly preferred catalysts include additional components for improving ethylene selectivity, such as amorphous silica dispersed within the zeolite crystal structure.
It consists of ZSM-5. This type of catalyst is described in US Pat. Nos. 4,060,568 and 4,100,219. In addition to ZSM-5 zeolite, other zeolites known in the art as small-pore crystalline aluminosilicate zeolites can also be used in the present invention. This small-pore zeolite can be either naturally occurring or artificial, such as erionite, chiabazite, zeolite T, zeolite ZK-5 and
There is ZSM-34. Zeolite T has U.S. Patent No.
No. 2950952, zeolite ZK-5 is U.S. Patent No.
3427195, and ZSM-34 is U.S. Patent No.
4079095 and 4079096, respectively. The crystal structure of this type of zeolite is such that the pore size of most of the zeolites is greater than 3 angstroms and smaller than 6 angstroms, allowing access to and exit from the intracrystalline free space. This zeolite also has almost 8
It is characterized by a hole window that can form a member ring. Of course, this eight-membered ring is created by the regular arrangement of tetrahedra that make up the anionic skeleton of crystalline aluminosilicate, and the oxygen atom itself is bonded to a silicon or aluminum atom at the center of the tetrahedron. . The pores that characterize this zeolite are approximately
A substantially circular or approximately 3.6 x 5.2 zeolite, such as ZSM-5, with uniform diameter pores of 3.9 angstroms.
It may also be somewhat oval, such as erionite, which has angstrom pores. In both cases, small pore zeolites primarily have pore sizes smaller than 6 angstroms. The pore sizes of these zeolites as well as other suitable zeolites are determined in Advances in
Chemistry Series. 101, 155-170 (1971)
Described in "Zeolite Organization" by WM Mayer and DH Olson. According to the invention, the conversion reaction of the ether-rich effluent from the first step to highly olefinic hydrocarbon products is carried out in a continuous multistage adiabatic reactor with interstage cooling. Cooling can be direct or indirect.
The zeolite catalyst is provided in each stage as a fixed bed or a moving bed. The amount of catalyst in each bed is such that a uniform temperature increase occurs throughout each bed as the heat of reaction is released during the conversion of the charge to high olefinic product.
The temperature rise through each bed and the amount of catalyst in each bed are based on the desired total conversion of the charge, the total adiabatic temperature rise of the conversion of methanol to olefins, and the sensitivity parameters of the conversion. By including the catalyst in each successive bed and providing cooling between stages of reaction heat removal, stable operation is achieved and undesirable conversion to gasoline boiling range hydrocarbons is avoided. The sensitivity parameters for a chemical reaction and a given catalyst are constant and can be approximated by assuming the Arrhenius dependence of the reaction rate on temperature and that the reaction rate is approximately linear over the relevant temperature range. Therefore, the sensitivity parameter can be calculated from the formula: θ=RTo ² /E (where θ represents the sensitivity parameter, R represents the gas constant, To represents the initial temperature, and E represents the activation energy). θ in the conversion of methanol to ethylene is approximately equal to 18°C. When _an exothermic chemical reaction is carried out in _an adiabatic reactor, the formula: The degree of conversion is obtained from the specific heat of the feed stream expressed in units of specific heat of the feed stream per mole of reactant entering. Solving for ΔT: ΔT=X _A (−ΔH)/Cp Therefore, the total adiabatic temperature rise for an exothermic reaction for a given feed composition and specific conversion can be calculated from the feed specific heat and the heat of reaction. In exothermic reactions which are less sensitive to variations in flow rate, catalyst stage, feed composition or charge temperature when carried out in an adiabatic catalyst bed, the temperature rise within the bed must be limited to the sensitivity parameter θ. If the total adiabatic temperature rise of the reaction is greater than the sensitivity parameter, multiple catalyst beds must be used to obtain the desired conversion under controllable conditions. The minimum number of beds required is equal to ΔT sum/θ. If the ratio of the total adiabatic temperature rise to the sensitivity parameter is not an integer, the next larger integer should be taken as the required minimum number of beds. According to the invention, the reaction mixture passing through the second stage reactor may flow through each catalyst bed in axial flow or preferably in radial flow. The radial flow configuration has a large cross-sectional area in the flow direction and a small bed thickness, resulting in a small pressure drop through each catalyst bed. Radial flow types can be used with either fixed or moving beds. The effluent from the second reaction zone is a gas stream containing ethylene, water, unconverted methanol, and unconverted dimethyl ether. Small amounts of light paraffin and other hydrocarbons having 3 to 10 carbon atoms are also present in low concentrations. The effluent is passed through a condenser to form three phases: an aqueous phase containing most of the unreacted methanol, a liquid phase containing C ₅ + hydrocarbons, and a gaseous phase containing the produced ethylene, other light hydrocarbons, and unconverted dimethyl ether. be done. Contacting methanol with a dehydration catalyst results in an equilibrium mixture. 2CH ₃ OHCH ₃ OCH ₃ +H ₂ ...(1) Zeolite catalyst converts methanol and dimethyl ether to olefin products. 2CH ₃ OH→CH ₂ ＝CH ₂ +2H ₂ O CH ₃ OCH ₃ →CH ₂ ＝CH ₂ +H ₂ O …(2) …(3) Since all these reactions are exothermic, the methanol converted in the second reaction zone generates more heat there. Only a portion (<90%) of methanol and DME is converted to hydrocarbons in this operation, resulting in high ethylene selectivity. Conversion of methanol and DME is 90
%, the increased amount is converted to gasoline boiling range hydrocarbons, such as aromatics. It will be appreciated that if a highly olefinic product is desired, the conversion reaction will be limited to producing the desired amount of ethylene. Conversion rates of about 30 to about 90% are commonly used. Because a partial conversion operation must be performed to obtain the desired olefin product, the effluent contains significant amounts of unconverted methanol and dimethyl ether. It is therefore good industrial practice to recover the unreacted feed from the product stream and return it to the initial operating step to improve the economics of the operation. This process is used in known processes for converting low aliphatic alcohols to high olefin products. Unreacted alcohol and unreacted ether are separately recovered and recycled to the operation, but both recycled streams are in each case mixed with fresh charge. However, according to a preferred form of the invention, if the recovered alcohol is recycled to the new charge to the operation, i.e. to treatment in the first reaction zone containing the dehydration catalyst, and the recovered ether is recycled to the second reaction zone containing the zeolite catalyst. Significant benefits can be obtained if the equilibrium mixture obtained from the first reaction zone is recycled to be processed in the first reaction zone. Many operational benefits are obtained by separating the recycled ether from the fresh charge and mixing it with the alcohol-ether-water equilibrium mixture while mixing the recycled alcohol with the feed alcohol.
As is clear from equation (1), recycling of alcohol and ether increases the conversion of alcohol to ether and water in the dehydration reactor. Also, if more of the feed is converted to the olefin product by directing the recycle streams to each as described above, the secondary ethers, which do not release as much heat as in the case of alcohols, are
Since it enters the reaction zone, the heat load in this reaction zone is reduced. The heat of reaction generated is:

Table: After conversion in the two reaction zones, the effluent from the second reaction zone is divided into three phases which are sent to a condenser. This effectively separates essentially all of the unreacted methanol and unreacted dimethyl ether into distinct separate phases. Two immiscible liquid phases, an aqueous phase and a hydrocarbon phase, are obtained along with a gaseous phase. Unreacted alcohol enters the water phase and unreacted ether enters the gas phase. Since the aqueous phase contains a large amount of water, most of which is not good to add to the fresh charge, the recovered alcohol mixed with this small amount of water is recycled after the water content of this phase is reduced. Methods such as evaporation and steam stripping are used to remove this water. Since unconverted ether is present in the gas phase along with ethylene and other hydrocarbon products from the operation, it is necessary to remove the ether from this gas phase and recycle it. This can be removed by passing the gas phase through a compressor to increase the pressure and then contacting the gas phase with an absorption liquid effective to selectively absorb dimethyl ether. Methanol or water can be conveniently used as absorption liquids because both selectively absorb dimethyl ether from ethylene and other product gases. It is preferred to separate the dimethyl ether from the absorption liquid before the absorption liquid enters methanol conversion and the DME is recycled to avoid undesirable side effects. Separation can be carried out by passing the absorption liquid containing dimethyl ether through a discharge column that can be easily separated by reducing pressure. The invention will now be explained in detail by means of examples with reference to the accompanying drawings. FIG. 1 is a schematic diagram for implementing the method of the present invention. FIG. 2 is the same diagram as FIG. 1, but using a different interstage cooling method. FIG. 3 is the same diagram as FIG. 1, but using radial flow through the catalyst bed. FIG. 4 is the same diagram as FIG. 2, but using radial flow through the catalyst bed. FIG. 5 is a diagram of four different schemes for passing radial flow through the catalyst bed. FIG. 6 is the same as FIG. 3 but with a catalyst regeneration device. FIG. 7 is a diagram similar to that of FIG. 1, but with a separation and recycling system for unconverted reactants. Figures 8, 9 and 10 plot the temperature versus catalyst bed length in the second step of the process to illustrate the temperature profile of the operation. First, in FIG. 1, the feed methanol containing about 50% water by weight enters the operation along line 2. This is mixed with a recycle stream rich in unreacted methanol from line 4 and a recycle stream rich in unreacted dimethyl ether (DME) from line 6. The mixed stream is heated to a reaction temperature of 285 to 370°C in a heat exchanger 8 and sent through a line 10 to a dehydration reactor 12, where methanol,
It is converted into an essentially equilibrium mixture of DME and water.
This equilibrium mixture exits the dehydration reactor, passes through line 14, is adjusted to 270-370°C in heat exchanger 16, and then passes through line 18 to multi-stage ZSM-5 fixed bed adiabatic reactor 2.
Enter 4. The mixture then contacts the first stage catalyst bed to produce hydrocarbons and to increase the temperature of the reactant and product mixture due to the exotherm of the conversion reaction. The mixture exits the first stage catalyst bed and is cooled by indirect heat exchange to be substantially isothermal with the reactant mixture entering the first stage catalyst bed. The cooled reactant and product mixture then enters the second stage catalyst bed where further reactant conversion occurs and further heat is generated, again raising the temperature of the reactant and product mixture. The mixture exits the second stage catalyst bed and is cooled by indirect heat exchange to approximately the same temperature as it entered the first stage catalyst bed. This conversion, exothermic and cooling operation is repeated until the final catalyst bed (stage N) is reached. The mixture exiting the penultimate stage (N-1st stage) is cooled by indirect heat exchange and enters the Nth stage catalyst bed. Upon exiting the Nth stage catalyst bed, methanol and
Once the desired conversion of DME has been achieved, the heated product and reactant streams exit the reactor via line 32 to heat exchanger 3.
4 and passed through line 36 to product recovery device 38
to go into. The cooled product and reactant mixture contains (1) a hydrocarbon liquid phase, (2) a large portion of unreacted methanol and
There are three phases: an aqueous phase containing a small amount of DME, and (3) a gaseous phase containing most of the produced light olefin hydrocarbons and most of the unreacted DME. The aqueous phase becomes a recycled water stream through line 4 by various separation methods such as steam stripping, evaporation and distillation. It contains a higher concentration of methanol than the raw aqueous phase resulting from the dehydration process. This circulation also contains most of the DME that was initially in the aqueous phase. It may not be necessary to use all of the above separation methods to obtain the desired composition of the recycled water stream. This treatment of the aqueous phase combines the water produced in the conversion of methanol to hydrocarbons with line 2
Most of the water initially present in the methanol feed is recovered and these materials are removed from the system via line 40, leaving only a small amount (<0.5% by weight) of methanol and
This results in an aqueous product containing DME. The gas phase remaining after condensation is compressed and sent to an absorber to recover unreacted DME. Suitable absorbents are methanol, feed methanol and water, each with varying efficiencies. DME rich solution is line 42
and 6 to either the dehydration reactor or to the multi-stage fixed bed reactor via lines 42 and 44.
Stripped to provide DME-rich flow. The DME-laden portion from the absorber is sent via line 46 to an olefin recovery unit similar to that used in conventional olefin plants for ethylene, propylene, butenes, and gasoline partial recovery. Alternatively, the gas phase can be compressed and sent to a distillation unit for recycling to recover the various hydrocarbons and DME. The hydrocarbon liquid phase can be stabilized for light olefin product recovery. The stabilized liquid can be mixed with gasoline boiling range components recovered from the olefin recovery unit to form the final gasoline or gasoline blend stock. Figure 1 shows two methods of processing DME recirculation streams. A diagram is shown in which all unreacted DME is supplied to the dehydration reactor. But if there is line 42
It is desirable to feed the unreacted DME therein directly to the multistage fixed bed reactor via lines 44 and 18. The advantages of this last operation are (1) increased conversion of methanol to DME and water in the dehydration reactor, and (2) less heat than methanol when a larger portion of the feed is converted to hydrocarbons. This is the load heat reduction in the multi-stage reactor due to the DME entering as non-generated DME.
This load heat reduction increases the stability of the multistage reactor. Further details of this recycling method are provided below with respect to FIG. Instead of the interstage indirect cooling shown in Figure 1, the desired cooling can be achieved by direct cooling between stages with (1) feed liquid, (2) water, or (3) low pressure steam (e.g., 450 KPa operating steam). I can do it. For this alternative method, only the multistage reactor of this system is shown in FIG. 2, but the rest of the system is the same as that of FIG. In FIG. 2, the methanol, DME and water mixture from the dehydration step enters reactor 124 via line 118 and the final product exits via line 132. Cooling fluids obtained from various sources enter directly into the interstage direct cooling zone via line 190. A number of liquids obtained from the feed to the operation can be used for cooling. The charge to dehydration reactor 12, ie the mixed stream entering from lines 2, 4 and 6, can be used. Other possible feed liquid sources for direct cooling are methanol from line 2, recycled aqueous phase from line 4, recycled DME from lines 6 and 42 or condensed product from the dehydration reactor (line 14). (cooled liquid). Other direct cooling fluid sources are water and low pressure (450KPa)
It is a flow of operation. The major difference between interstage direct cooling using water or steam and one working stream is that the working stream contains reactants that are injected and will be converted to olefins in a subsequent stage of direct cooling. The following preferred operating conditions for the multistage adiabatic reactor described above are based on results obtained from small scale reactors. It is well known that high operating pressures reduce the yield of light olefins. Therefore, the inlet pressure of the multistage reactor should be less than 520 KPa, preferably 380 KPa. Although it is economically preferable to include all stages in one container, it is not critical. For interstage cooling, liquid injection devices and spargeers are used in the mixing zone between stages. The fixed catalyst bed in each stage has a diameter of 1.5 to 6 m, preferably 2.5 to 4.5 m. Due to pressure drop and economic considerations, the space velocity (WHSV on a total feed basis) should be in the range of 0.2 to 2. Preferred operating temperatures are from about 270°C to about 370°C for ZSM-5 type zeolites and from about 315°C to about 485°C for small pore zeolites. Catalyst aging can be compensated for by increasing the inlet temperature at each stage. Catalysts consisting of large crystal (at least about 1 micron) ZSM-5 zeolite may be used in multi-bed reactors as described in U.S. Pat. Nos. 4,025,571 and 4,148,835. In order to reduce the hexane decomposition activity (alpha value) of the ZSM-5 type catalyst, steam treatment may be performed in advance. Under the operating conditions described here, the alpha value must be greater than 20. The catalyst also includes U.S. Patent Nos. 4,060,568 and 4,100,219.
Ethylene selectivity-improving components, such as intracrystalline silica, may be added as described in No. Of course other suitable catalysts can also be used. The total conversion of methanol and DME should be limited to 30 to about 90%. At high conversions, even in the preferred temperature range, a significant amount of methanol will be converted to aromatics. To ensure partial conversion, the reactor geometry and operating conditions do not result in unstable or sensitive reactor operation (i.e., slight variations in operating conditions may cause excessive temperature increases and therefore excessive conversion of the desired light olefin product). That is important. For stable operation of the reactor, the temperature rise in each bed must be limited to approximately the value of the sensitivity parameter, ie, in most cases about 10 to about 25°C, often about 18°C. The minimum number of fixed beds can be calculated from the total adiabatic temperature rise and sensitivity parameters as described above, and is usually between 2 and 15.
Often 4 to 10. Generally, the greater the number of beds, the less sensitive the system is to fluctuations in feed temperature to any fixed bed. However, economic considerations suggest that the use of a smaller number of beds within a constant temperature fluctuation will reduce the effluent from any bed to approximately
It will not exceed 370°C, or about 485°C in small pore zeolites. The methanol feed to this process has a nominal water content of 0
From about 70% by weight, preferably from about 20 to about 50% by weight
You can change it with Any "equivalent" feed of methanol and water can be used. a zeolite having substantially the same composition as the equilibrium mixture obtained by contacting the methanol and water feed with the dehydration catalyst given the appropriate recycle stream recovered from the methanol and water feed and the products of the operation in question; When producing the feeds to the catalytic reactor, the methanol and water feeds can be said to be "equivalent" to a given methanol and water feed. The diagram shown in FIG. 3 is the same as that of FIG. 1 above, but in this case with a radial flow design of the reaction mixture in the individual catalyst beds of the multi-bed reactor. This reactor geometry allows the gaseous reaction mixture to pass radially through each catalyst bed, with the catalyst within each bed being enclosed in a perforated annular cage. When such a cage is used, the reaction mixture flows from the annular space in the reactor through the outer surface of the perforated cage toward the center of the catalyst. Since the top and bottom of the catalyst cage are closed, the reactant and product mixture cannot flow axially through the catalyst bed nor can the mixture take a shortcut to the hump center tube for the next cooling stage. As the reaction mixture passes through the catalyst, at least a portion thereof is converted to olefinic products, generating an exotherm and raising the temperature of the mixture. The inner surface of each cage is also a perforated plate that separates the central tube. The reaction mixture thus flows through the internal holes into the central tube which leads the heated mixture to a cooling stage. The reaction mixture passes through lines 220 and 222 to outer ring 22.
4 and flows toward the center and comes into contact with the first stage catalyst bed 226. The temperature of the reactant and product mixture in which hydrocarbons are produced increases in the flow direction due to the exotherm of operation. The mixture exits the first stage 224 through the central tube 228 and is cooled by indirect heat exchange with the coolant in the coil 230 to substantially the same temperature as the first stage feed temperature. The cooled reactant and product mixture then reenters the second stage through the outer ring. The reactant and product mixture is contacted with a second stage catalyst bed for further conversion of the reactants and generation of additional heat. The same conversion, heat generation and cooling operations as above are repeated until the final catalyst N stage is reached. Upon exiting the N catalyst bed, the desired conversion of methanol and DME is achieved. The hot product and reactant streams exit the reactor through line 232 and are cooled and partially condensed in heat exchanger 234 in line 236.
The product then enters product collector 238 for further processing as described above with respect to FIG. A method for recycling unreacted DME to the multi-stage reactor as shown in Figure 1 above is described below with reference to Figure 7. Instead of the indirect interstage cooling shown in the figure, the desired cooling effect can also be achieved by direct cooling. This direct cooling can be accomplished in a manner similar to that described above with respect to FIG. 2, but is illustrated in FIG. 4 for the radial flow configuration of FIG. Four radial flow directions are available for any stage in the diagrams shown in FIGS. 3 and 4. FIG. 5 shows these directions for one fixed catalyst bed. The radiation flow is shown in Figure 5.
It may flow inwardly from the outer surface, as shown in FIGS. 5c and 5b, or outwardly from the center, as shown in FIGS. 5c and 5d. It may also be cocurrent as in Figures 5a and 5c, or countercurrent as in Figures 5b and 5d. Although the Figure 5a type is shown in Figures 3, 4 and 6, any of the other three types can be used if that type is particularly good. It will be appreciated that while the flow form of Figure 5c can be adapted with minor modifications to the piping of equipment using the flow form of Figure 5a, the flow forms of Figures 5b and 5d require significant piping changes. Even in the basic design, Figure 5
Flow types a and 5c do not have any unusual design or piping problems, but those in Figures 5b and 5d require external piping to the reaction vessel to achieve the desired radial flow type. Since the zeolite catalyst used in this process must be periodically regenerated, the catalyst is moved by gravity through a series of catalyst beds and recovered from the last bed to another area for regeneration before the first bed of the system. We can prepare the equipment to be returned to you. FIG. 6 is a combination diagram of a radial flow and moving bed multi-stage reactor. Since the figure is the same in all respects as described above with respect to FIG. 3, except for the catalyst regeneration equipment, only the catalyst regeneration method in the figure will be described.
The zeolite catalyst regenerated in the regenerator 450 of FIG. 6 is sent along a line 452 along with the necessary additional feed catalyst entering from a line 454 through a hopper 456 that supplies catalyst to the first stage catalyst basket for a multi-stage reaction. Enter the top 424 of the machine. The catalyst moves down by gravity flow through passage 458 to the next stage. The catalyst generally flows from stage to stage in overall co-current relationship with the reaction mixture. The reaction mixture passes through each catalyst bed in a radial flow, with a “cocurrent relationship throughout”.
and similar phrases mean that the fresh feed is first contacted with fresh or freshly regenerated catalyst and through each stage is gradually converted by the catalyst to the desired product, the catalyst being gradually and constantly deactivated until finally the desired product is converted to the desired product. Both the reaction mixture at conversion and the deactivated catalyst are used to describe the flow as it exits the lower end of the multistage reactor. Catalyst stream from stage N is collected in hopper 460 and sent by line 462 to catalyst regenerator 450. Separately, the catalyst from each stage may be sent directly to the catalyst regenerator 450. In this case, the catalyst regeneration of each stage is carried out separately. Catalyst recycling through the moving bed can be continuous or intermittent depending on the rate of deactivation, which in turn depends on the particular catalyst used and operating conditions.
A moving bed multi-stage radial flow reactor has all the advantages of a fixed bed radial flow reactor and also provides a means of continuous reactor operation. Fixed bed multi-stage reactors, on the other hand, require a separate regeneration movement cycle. Alternatively, the feed stream can be introduced from the bottom of the multistage reactor so that the feed and zeolite catalyst flow in generally countercurrent relationship to each other throughout. This means that the “highest concentration” feedstock contacts the “lowest” active catalyst and the “lowest concentration” feedstock contacts the “highest” active catalyst.
The usual advantages of countercurrent flow, such as contacting the active catalyst and thus eliminating the excessively high reaction rate problems that often occur in the first bed of cocurrent systems, are obtained. Interstage indirect cooling is illustrated in FIG. However, direct interstage cooling is also possible as described above for stationary catalyst beds in connection with FIG. In Figure 6, the catalyst activity decreases as the number of stages increases, with the first stage having the most active catalyst and the Nth stage having the least active catalyst. The input temperature increases with the number of stages, ie the Nth stage is operated at the highest temperature. However, for stable and insensitive reactor operation, it is desirable to control the product and reactant mixture temperatures in separate stages close to the charging temperature of each stage, as in cocurrent. Based on the results obtained with the small-scale isothermal reactor, the following preferred operating conditions for the multistage radial flow reactor have been developed. Since low pressure operation is important to obtain a high yield of light olefins, the charging pressure of the multistage reactor should be 380 KPa or less, preferably 275 KPa or less. Containing all zeolite catalyst stages within one vessel is usually economically preferred but not critical. Interstage cooling includes liquid injection systems and diffusers with partitions in the mixing zone between stages. The diameter of this reactor is approximately 1.2
From about 4.5 m to about 4.5 m, preferably from about 1.8 to about 2.8 m is required. The thickness of the radial catalyst bed is about 0.3 to about 1.2
m, preferably about 0.6 m. The center tube diameter should be about 0.3 to about 1.0 m. According to the pressure drop and economic considerations, the space velocity (on a total feed basis,
WHSV) should be about 0.5 to about 5.0, preferably about 1.5 to about 3.5/hour. The pressure drop obtained when the reaction mixture is delivered in radial flow mode to the zeolite catalyst bed is considerably lower than when operating in axial flow mode at the same conversion of the same amount of catalyst and reactants. In fact, the pressure drop in the radial flow reactor is a factor of 3, which is about an order of magnitude lower than when using a conventional axial packed bed reactor operated at the same WHSV. Conversely, for a given catalyst capacity and permissible pressure drop, the throughput of a radial flow reactor is significantly higher than using a conventional axial flow packed bed reactor. Radial flow large diameter fixed beds (pancake reactors) with L/D ratios less than 1/2 are preferred since it is more economical to build taller smaller diameter reactors. Preferred operating temperatures range from about 270°C to about 370°C for ZSM-5 type zeolites and from about 315°C to about 485°C for small pore zeolites.
Compensation for catalyst aging in this radial flow reactor can be adjusted by varying the charging temperature of the reactant and product streams to each stage. As mentioned above, catalyst activity in a moving bed reactor decreases progressively from top to bottom (ie, from stage 1 to stage N), with relative activity depending on the rate of catalyst circulation through the system. Optimum conditions are achieved when the catalyst circulation rate is very low and there is a significant difference in the catalyst activity of the first and Nth stages. In this case the relative cooling of the subsequent stages can be reduced so that the feed temperature of each stage increases with increasing number of stages. The temperature in this configuration generally varies from 285°C in the first stage to as high as 355°C in the final stage. In Figure 6, the reactants and catalyst transfer are cocurrent;
That is, the fresh feed encounters the most active catalyst. Considering the stability of the reactor, a desirable modification of this method is countercurrent operation, in which case the feed flow enters from the bottom of the reactor and the N
Encounter with the lowest active catalyst in the stage. Reactant concentration decreases as the reactant and product streams advance upward and encounter more active catalyst. Suitable catalysts, conversion rates and feed compositions used in the radial flow reactor are generally the same as those described above for use in the axial flow reactor shown in FIGS. 1 and 2. 50:50 (weight:weight) under typical operating conditions
The methanol-water feed requires 6 stages for 50% conversion. As the reaction proceeds in the reactor, the reactant concentration decreases and the amount of catalyst at each stage increases with the direction of reactant flow. (Temperature rise per bed is almost constant), and when indirect cooling is used between stages in a fixed bed radial flow reactor (for example, as shown in Figure 3), the relative ratio of desolvation amount for 6 beds is 1.0: 1.10: 1.21: 1.32 :1.49：
It is 1.78. Since the radial dimension of the catalyst cage is constant, the above ratio is the ratio of the heights of each stage. Practical limits for minimum and maximum height are determined by production capacity and feed flow failure distribution. The axial length is about 0.6 to about 3 m, preferably about 0.9 to about 1.2 m. The relative proportions of catalyst in each stage of the moving bed reactor configuration are determined by catalyst deactivation and are specific to a given catalyst. Now in FIG. 7, fresh methanol feed containing up to about 50% water by weight enters from line 502 and is mixed with an aqueous product recycle stream rich in unreacted methanol from line 504. This recycle stream is recovered in the subsequent steps described below. The mixed stream passes through line 506 to heat exchanger 5
08 to the reaction temperature (285-370℃) and line 5
10 to a dehydration reactor 512 having a fixed bed 514 of gamma alumina. The mixture is converted by the catalytic action of the alumina into an essentially equilibrium mixture of methanol, dimethyl ether (DME) and water and exits the reactor at line 516. A recycle stream of unreacted DME recovered from the absorption-desorption operation described below passes through line 518 and is mixed in line 520 with the equilibrium mixture from line 516. The resulting mixture is heated in a heat exchanger 522 to a reaction temperature suitable for the second reaction zone (270 to
370℃) or cooled. The mixture then passes through line 524 to reactor 52 with a ZSM-5 catalyst.
It flows during 6. Reactor 526 has multiple catalyst beds and interstage cooling of the reaction mixture. The floor may be a fixed bed or a moving bed. The reaction mixture may flow axially or radially through each of these beds. As mentioned above, interstage cooling can be done by direct or indirect methods. Effluent gas from reactor 526 consisting of unconverted methanol, unconverted dimethyl ether, ethylene, other hydrocarbons, and water vapor enters condenser 534 via line 532 where methanol, water vapor, and C ₅ + hydrocarbons are condensed. do. The condensed mixture and non-condensed gases pass through line 536 to separator 538 which separates the mixture into three phases. A typical composition (wt %) of the effluent leaving reactor 526 and separated into three phases in separator 538 is as follows: Methanol Dimethyl Ether Water CO, CO ₂ , H ₂ Total Methane Ethane Ethylene Propane Propylene Isobutane n-Butane Butenes C ₅ ⁺ 8.7 8.1 70.0 <0.1 0.1 <0.1 3.3 0.6 2.1 0.6 0.2 1.0 5.1 The bottom layer in separator 538 contains water produced by the conversion of methanol to hydrocarbons, which is present in the feed. The aqueous layer consists of unreacted water, unreacted methanol, and a small amount of dimethyl ether. The second layer between the water and gas layers is C ₅ ⁺ hydrocarbons, ie hydrocarbons consisting of high molecular weight aliphatic and aromatic hydrocarbons. The top layer consists of light hydrocarbons saturated with gas and water vapor, primarily ethylene and a substantial portion of unreacted dimethyl ether. First, the aqueous layer is drawn off and sent through line 540.
Since there is an unnecessary amount of water in this aqueous phase, a portion of it is sent via line 542 to a steam stripper 544. The steam introduced into the stripper 544 removes unreacted methanol and traces of DME from the aqueous phase portion flowing down through the stripper. The separated product is passed through line 546 to join the remaining aqueous phase portion in line 504 and to line 502 where it is mixed with fresh methanol feed as described above. unreacted methanol and
The DME separated water is discharged from stripper 544 via line 548. It will be appreciated that the amount of aqueous phase portion directed to the stripper recycled in line 504 may be varied to limit the amount desired for efficient operation of the process. First, at least the volume of water phase directed to the stripper resulting from the conversion of methanol and DME to hydrocarbons should be disposed of through line 548. Operational variables such as the nature of the absorbent used for DME recovery will dictate more or less aqueous phase stripping. This stripping operation can be adjusted as the operation and product require. The liquid hydrocarbon phase containing _C5 ⁺ hydrocarbons is transferred to separator 5
38 via line 550 and sent to the next device for collection or further processing. For example, this portion can be used as a useful portion for gasoline blending. The third phase in separator 538 contains substantially all of the ethylene produced, as well as light hydrocarbons and trace amounts of CO, _CO2 , and _H2.
and a gas phase containing relatively large amounts of dimethyl ether. This gas phase exits through line 552 and is compressed by compressor 554 to approximately 1500 KPa.
56 to a gas absorption column 558 having a column bed 560 packed with, for example, Bar saddles, Laschichi rings or Paul rings. Water flows from line 562 to absorption tower 5
The water admitted to 58 flows countercurrently with the gas ascending the column, effectively removing dimethyl ether from the gas. The recovered DME dissolved in water exits column 558 in line 564 and enters discharge column 566 where the dimethyl ether is separated from the water under reduced pressure. Water passing through the discharge tower 566 is recycled via line 562 to the absorption tower 558 for reuse. Water is added to the absorption system via line 568 if necessary. The separated dimethyl ether is recycled from discharge column 566 via line 518 to line 520 where it is recycled to the equilibrium mixture of methanol, dimethyl ether and water exiting dehydration reactor 512 for conversion to the ethylene-rich product described above. will be merged with A partially purified ethylene stream leaves absorption column 558 via line 570 and is sent to gas plant 572. Ethylene of the desired purity is delivered from the gas plant via line 57.
Heavy hydrocarbons such as propylene and butenes recovered via line 4 exit the gas plant via line 576. Gas plants also serve for CO, CO ₂ , H ₂ , CH ₄ and C ₂ H ₆ removal. The following examples illustrate the invention. Example 1 The conversion reaction of methanol to olefinic hydrocarbons is highly exothermic and, unlike ammonia synthesis or methanol production, there are no thermodynamic equilibrium constraints. Therefore, high temperatures do not limit the conversion reaction or temperature increase in a given catalyst bed. In reactions that are not overly sensitive and allow for control over variations in flow rates, catalyst conditions, feed composition, and charge temperature, the temperature rise in each bed should be limited to approximately the value of the sensitivity parameter, ie, about 18°C. The table below shows the allowable ΔT/
The effects of the floor are shown together. The considered condition is 50% water and 50% methanol in the dehydration reactor.
Stable 50% total conversion with 4, 6 and 8 beds of zeolite catalyst in the post-conversion reactor charged with feed of . The corresponding temperature rise per floor is 27, respectively.
The temperatures were 18 and 13.6°C. Using the above equation, the total adiabatic temperature rise for olefin production from 50% conversion of this methanol feed is about 108°C. This is the initial feed charge temperature if all catalysts were in a single bed.
This means that the outflow temperature will be approximately 401°C compared to 293°C. Such a temperature increase would be undesirable as it would lead to operational instability and production of substantial amounts of high boiling products. Furthermore, the gradual increase in feed charge temperature required to compensate for catalyst aging further exacerbates this unstable operation. The table shows that with multistage beds and indirect interstage cooling, olefinic products can be obtained without conversion to aromatics at outlet temperatures of 320°C (4th stage), 311°C (6th stage) and 307°C (8th stage). ing.

Table: For a given feed dilution (50/50 in this case), the more stages used (and the correspondingly smaller temperature rise per bed), the more stable the reactor performance. This point is demonstrated by varying the first stage charge temperature by 5, 7 and 10°C. The interstage cooling is indirect, and the coolant in this indirect heat exchanger is 288℃
was maintained. From the conversion rate obtained as shown in the table, the stability of the 8-stage type is not much higher than that of the 6-stage type, and it requires a considerably large investment.
Elevated temperature/floor is best. The sensitivity parameter for this reaction is approximately the same as the temperature rise obtained in each bed of a six-stage system when processing a 50/50 methanol-water feed at 50% total conversion. The temperature state of this six-stage system is shown in FIG. FIG. 9 illustrates the temperature state of the four-stage reactor shown in the table with a 7°C fluctuation. The "X" temperature state is obtained at reference conditions (no fluctuations). From this figure 7
°C fluctuation is caused by excessive temperature rise (temperature
370℃ or higher), producing a considerable amount of unnecessary products. EXAMPLE This example illustrates the effect of feed composition and total conversion on the minimum number of zeolite catalyst beds required to achieve stable operation while producing high olefinic products. The table summarizes the minimum number of ZSM-5 zeolite catalyst stages required for various feed compositions at 50% and 80% conversion. The feed composition in this table is methanol entering the multi-bed zeolite catalyst system;
It is expressed as weight percent methanol in water entering the dehydration reactor giving an equilibrium mixture of DME and water.

[Table] 50/50 (wt%) at typical 50% conversion rate
Six stages are required for the methanol water feed. As the reaction progresses in the reactor, the reactant concentration decreases,
The amount of catalyst in each stage therefore increases progressively in the direction of reactant flow. (Constant temperature rise per bed) The relative proportion of the 6 beds with indirect interstage cooling in the above example is
1.0:1.10:1.21:1.32:1.44:1.78. Typical temperature conditions for a six-stage reactor are shown in FIG.
Note the uniform temperature rise per bed, the interstage cooling of the reaction mixture back to the first stage charge temperature, and the increase in catalyst bed thickness as the reaction mixture passes through the six successive stages. Similarly, the 10-stage method that achieves 80% conversion rate is as follows: 1.0: 1.09: 1.20: 1.33: 1.5: 1.72:
Ten beds of 2.00:2.41:3.00:3.98 would be required. Various variations of interstage cooling are possible as described above. The table shows various preferred relative amounts of coolant in a six stage system. In water and feed cooling, the bed thickness increase may be essentially the same as detailed above. For steam cooling, the thickness fraction is larger.

Table: Example This example shows the effect of a high water content feed, except for the feed composition, which is the same as the four-stage system of the example. Figure 10 shows 67/33 (weight) H ₂ O/
The effect of water dilution and temperature variation when charging MeOH feed is shown. The total conversion rate of this system under stable operating conditions is 50%. The temperature rise per bed is again approximately 18°C. The temperature state of this system is shown in Figure 10.
is shown. The "X" curve is for the same basic temperature rise for each of the four beds. The temperature curve obtained for a 10°C variation of the charging temperature to the first stage is also shown in this figure. It should be noted that the starting temperature of this four-stage system is lower than the six-stage system of FIG. Figure 10 shows that compared to the case where the basic fluctuation is 10°C,
Although the stage outflow temperature is significantly increased, it is gradually lower than the first stage outflow temperature obtained in the 7° C. fluctuation of the four-stage system of the example. In fact, with a high water content feed, a 10°C variation will not cause the first stage effluent temperature to exceed a limit of 370°C. Figure 2 and Figure 1
0, it can be seen that the high water content methanol feedstock can handle larger fluctuations in the first stage charging temperature of a continuous multi-bed system, while avoiding high temperatures that tend to cause side reactions. EXAMPLE The pressure drop of the reaction fluid across the catalyst bed is determined using Ergin's equation: dp/dZ=G ₂ /D _pgc 〓・1−ε/ε ³ [1.75+150(1−ε) μ/φD _p G] Can calculate. Applying the ideal gas law to the density change due to pressure drop and integrating the above equation, (In the formula, G = mass flow rate φ = shape factor D _p = catalyst particle diameter ρ = gas density g _c = constant ε = porosity μ = viscosity Po = charging pressure M = molecular weight T = average temperature of gas R = gas constant Z = distance in the flow direction). This equation can be used to compare the pressure drop in an axial flow reactor to that in a radial flow reactor. The feed to both reactors was the same methanol-water mixture and the following operating conditions were used: WHSV Charging Temperature Porosity 3/hr (overall) 293°C 0.38 A fixed bed reactor with a 3-foot packing height according to the above equation. The pressure drop is 2.76KPa, but the pressure drop of a radial flow reactor with the same capacity and a thickness of 0.3m is
It becomes 0.21KPa. From the above results, it is clear that the conversion of methanol to olefins in a radial flow reactor can reduce the pressure drop by an order of magnitude compared to that obtained in an axial flow reactor. Examples Reaction: 2CH ₃ OHCH ₃ OCH ₃ +H ₂ O is an equilibrium limited reaction. Due to the Lusciaterie principle, addition of dimethyl ether to methyl alcohol reduces the conversion. This effect is caused by a temperature of 315℃ and a pressure of
This is illustrated by introducing two different feed streams into a 310 kPa dehydration reactor. In the case of , the mole amount of the feed that enters the dehydration reactor without recycling DME The mole amount that comes out MeOH 50g 1.56 0.45 DME 0g 0.0 0.56 H ₂ O 50g 2.78 In the case of 3.33, the mole amount of the DME recycled feed that enters the dehydration reactor The mole amount that comes out Quantity MeOH 50g 0.78 0.47 DME 25g 0.54 0.70 H ₂ O 50g 2.78 2.94 MeOH Conversion rate = 1-Incoming C/Outgoing C = 1-0.45/1.56=71.2% In case of =1-0.47/0.78=39.7% This implementation The example shows that adding dimethyl ether to the feed stream entering a dehydration reactor reduces the methyl alcohol conversion made in the reactor.

[Brief explanation of drawings]

FIG. 1 is a schematic illustration of the implementation of the method of the invention. Figure 2
1 is a schematic diagram of the present invention using the same method as in FIG. 1 but using a different method for interstage cooling. Figure 3 is Figure 1 of the present invention.
This is the same method as above, but with a radial flow passing through the catalyst bed. FIG. 4 is a diagram of the same method as FIG. 2 of the present invention, but with radial flow passing through the catalyst bed. FIG. 5 is a diagram of four different ways of passing radial flow through the catalyst bed for use in the present invention. FIG. 6 is a schematic diagram of the same method as FIG. 3 of the present invention, but including a catalyst regenerator. Numbers in each figure above: 8, 16, 34, 208, 2
16,234,408,416,434...heat exchanger, 12,212,412...dehydration reactor, 24,
224,424...multistage reactor, 226,426
...Catalyst cage, 230,430...Interstage cooling device, 3
8,238,438...Product recovery device, 450...
Catalyst regenerator. FIG. 7 is a schematic diagram of the same method as FIG. 1 of the present invention, but including separation and recycling of unconverted reactants. In the figure, 512... dehydration reactor, 526... multistage reactor, 538... separator, 544... stripper,
554...Compressor, 558...Absorption tower, 566...Discharge tower, 572...Gas factory. FIG. 8 is a curve diagram showing the temperature characteristics in the multistage adiabatic reactor of the present invention, with the horizontal axis representing the bed length ratio and the vertical axis representing the temperature (° C.). Figure 9 shows the 50:50 ratio of the present invention.
H ₂ O: MeOH charging, 7°C fluctuation in 4-stage system,
It is a temperature characteristic diagram in indirect interstage cooling. The horizontal and vertical axes are as shown in FIG. Figure 10 shows the present invention.
67:33 _H2O :MeOH charging, 10 in 4-stage system
FIG. 3 is a temperature characteristic diagram of temperature fluctuation and indirect interstage cooling. The horizontal and vertical axes are as shown in FIG.

Claims (1)

Claims: 1. After contacting a methanol-containing feed with a dehydration catalyst at elevated temperatures to produce an ether-rich product, the product is then purified by: Any crystalline aluminosilicate having an alumina molar ratio and a constraint index of 1 to 12, or (ii) a pore window size in which the majority of the pores are smaller than 6 angstroms and are formed by eight-membered rings with oxygen atoms. contacting the ether-rich product with said catalyst under heat, passing the product alternately through a bed of zeolite and a cooling zone to achieve a desired conversion of the olefinic hydrocarbon product; The temperature decrease as the mixture passes through each cooling zone is substantially equal to the temperature increase in the immediately preceding catalyst bed, and the temperature increase as the reaction mixture passes through each catalyst bed is less than the sensitivity parameter of the conversion of methanol to ethylene. A process for the conversion of methanol to olefinic hydrocarbons, characterized in that it is not large and the number of catalyst beds is at least equal to the ratio of the total adiabatic temperature rise to a sensitivity parameter at a given conversion rate. 2. The method of claim 1, wherein the reaction mixture passes axially through each zeolite catalyst bed. 3. The method of claim 1, wherein the reaction mixture passes radially through each catalyst bed. 4. The method according to any one of claims 1 to 3, wherein each zeolite catalyst bed is a fixed bed. 5. The method according to any one of claims 1 to 3, wherein each zeolite catalyst bed is a moving bed. 6. The method of claim 5, wherein the reaction mixture stream is co-current with the catalyst stream. 7. The method of claim 5, wherein the reaction mixture stream is countercurrent to the catalyst stream. 8. Claims 1 to 7 in which the olefinic hydrocarbon product is separated into its components, unreacted methanol is recycled to the dehydration catalyst and unreacted ether is recycled to the zeolite catalyst bed. Any method described.