JPS58146515A - Conversion of methanol to olefin - Google Patents

Abstract

(57) [Summary] This bulletin contains application data before electronic filing, so abstract data is not recorded.

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. Due to limited capacity, the increasing demand for olefin feedstocks has periodically led to petrochemical feedstock shortages. A source of ethylene from another non-petroleum source is one means of keeping pace with ethylene and semi-olefin requirements.

Methanol and/or dimethyl ether obtained from coal, gas or biological sources can be converted to more complex hydrocarbons such as olefins and aromatics by the use of catalytic zeolites such as ZSM-5 zeolite. It is known that conversion is possible. Ethylene is an olefinic hydrocarbon 1a+ obtained by this catalytic conversion method. This reaction is extremely exothermic and the olefins initially formed are likely to undergo further reaction to form aromatic hydrocarbons useful in the production of automobile gasoline.

There is much patent literature on the conversion of methanol and/or dimethyl ether to light olefins. Therefore H
A process for producing olefins by catalytic conversion from aliphatic ethers using zSM-5 zeolite is described in U.S. Pat. No. 4s 94,1.
It is described in No. 06. US patent number! 1,979,4
No. 72 describes that the conversion reaction of lower alcohols and their ethers with a mixture of antimony oxide and ZM-5 zeolite produces a mixture of ethylene, propylene and mononuclear aromatics.

U.S. Patent No. 4,025,572 announces that ZSM-5 can be mixed with an inert diluent to improve ethylene selectivity, while U.S. Patent No. 4,025,575 provides similar results. It is obtained by using low pressure partial pressure of material. Ethylene selectivity is at least about 1 micron large crystal form 28M
-5 by using only zeolite (U.S. Pat. No. 4,02
5,571) or by mixing the above zeolites with additive metals (U.S. Pat. No. 4,148.835). Better selectivity is achieved by dispersing amorphous silica within the crystal structure of the zeolite catalyst. can be obtained by (
U.S. Patent Nos. 4.060,568 and 4.100°2
No. 19].

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 highly exothermic depending on the particular reactants. For example, the heat generated for the conversion of lower alcohols to hydrocarbon products can be estimated as tFi.

Methanol 2500-4600 ethanol
460-1450 Gropanol
Although it is desirable for the 35-840 reaction to be exothermic as it eliminates the need for an external heat source, the large amount of heat generated requires a substantial investment in a complex reactor and considerable internal cooling equipment. From the table above, it can be seen that methanol conversion is significantly exothermic at this point, while ethanol is 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 high temperatures in the catalyst bed to complete the reaction, whereas ethanol is much 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 favorably produces a narrow product distribution. It is important that the conversion of methanol to useful products be 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 more effectively use reaction contact volumes and ancillary equipment for reactant conversions at higher pressures. However, the high pressure using charged methanol requires a large amount of 1.2
．． 4.5-Tetramethylbenzene (durene) tends to produce undesirable by-products, but low pressures, for example below 450 kP'a, are due to the formation of light olefins.

Various methods have been used to control the heat generated in this process. For example, US patent no. No. 4931.349 (using light hydrocarbon diluent), No. 4.052;
-25-inch suppression reaction and dense fluidized catalyst bed].

U.S. Patent No. 4. No. D 35,430 describes a method for controlling the heat generated in the conversion of methanol to gasoline boiling range products. After converting the feedstock into an equilibrium mixture in a dehydrating catalyst bed, it is passed through a series of zeolite catalyst beds of increasing size to produce methanol, dimethyl ether and
Also, Fi) interstage cooling with light hydrocarbons to remove reaction heat. The temperature rise in any bed does not exceed about 28°C and the total temperature rise does not exceed about 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. 2J19,620, 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 the number of fixed catalyst beds, with the main stream passing through only one bed.

Reactant flow through each bed may be axial or radially inward. US Pat. No. 2,475,855 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 graded hydrogen 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 described method 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, it has been observed that olefin selectivity can only be improved by partial conversion of the feed alcohol ether; however, this approach also involves significant economic losses, as methanol costs, for example, require collection and recycling of unreacted reactants. is also often recognized.

The present invention combines methanol enrichment or its corresponding ether in the presence of a zeolite catalyst by arranging 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 during 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 bed can be minimized by passing the reaction mixture radially through each bed. Additionally, by recycling unconverted alcohol to the dehydration step and unconverted ether to the conversion step. It has also been discovered that significant economic improvements in operation are possible.

In accordance with the present invention, a methanol-containing feed is contacted at elevated temperature with a dehydration catalyst to produce an ether-rich surface product and the product is characterized in that: (1) a pore size greater than 5 angstroms, a silica to alumina molar ratio of at least 12; or (-) most pore sizes are 6
The catalyst is brought into contact with the catalyst under heating through a cooling zone and a catalyst bed made of any crystalline aluminosilicate zeolite having a pore window size of 1.5 Angstroms or less and formed by an 8-membered ring containing an oxygen atom. to achieve a predetermined conversion rate to olefinic hydrocarbons,
The decrease in temperature of the reaction mixture as it passes through each cooling zone is substantially equal to the increase in temperature in the immediately preceding catalyst bed, and the increase in temperature of the reaction mixture as it passes through each catalyst bed increases the conversion of methanol to ethylene. A process is provided for the conversion of methanol to an olefinic hydrocarbon slag product in which the number of catalyst beds is not more than an order of magnitude greater than the sensitivity parameter and the number of catalyst beds is at least equal to the ratio of the total adiabatic temperature rise to the sensitivity parameter for a given conversion reaction. It will be done.

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 may 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, can react with olefinic hydrocarbons, in particular in a reactor with specifications that allow effective regulation of the heat output during the conversion of the ether-rich intermediate to olefinic hydrocarbons. , especially 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 an olefinic product, the feed stream to the operation may include small amounts of other alcohols, For example, the methanol may include ethanol, propatool, and isopropanol. In the first step of the process, methanol is contacted with a dehydration catalyst to produce water and an ether-rich intermediate product.

The dehydration catalyst may 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 simple ethers and various 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. These dehydration reactions themselves are generally known for the use of 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 with the methanol charge no intramolecular dehydration is possible and therefore the dehydration reaction proceeds exothermically only to form dimethyl ether.

The process of the present invention consists of two successive catalyst contacting steps, with both steps co-exothermic. The exotherm in the first step is limited by the inhibition of the conversion of methanol to an 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 reduced to about 270 to about 430°C depending on the second stage zeolite catalyst by indirect heat exchange with the circulating heat exchange fluid. Temperatures are well controlled. 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 highly exothermic and occurs rapidly in the presence of certain crystalline zeolites, particularly ZSM-5 type crystalline zeolites and small-pore crystalline zeolites.

Generally, the ZSM-5 type zeolite used in the present invention is a crystalline zeolite having a silica to alumina ratio of 12 or greater and a constraint index (C,I, ) of from 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 No. 4,894.
, No. 106, no.

Preferred zeolites are 28M-5 type zeolites, such as ZSM-5, zSM-11, ZSM-12, ZSM-2.
5, ZSM-45 and ZSM-5B, and ZSM-
5 is particularly preferred.

ZSM-5 is described in U.S. Patent No. 1702.886 and ZSM-11 is described in U.S. Patent No. 4709.97.
No. 9, ZSM-12 was published in U.S. Patent No. 5,852,449.
ZSM-23 in U.S. Pat. No. 4.07 &842; ZSM-35 in U.S. Pat. No. 4.016,245;
ZSM-38 is also described in U.S. Pat. No. 4,046,859.

Particularly preferred catalysts are large crystals, i.e., U.S. Pat. No. 4,025
, 571 and 4,148,835, and other particularly preferred catalysts are those comprising a ZSM-5 type zeolite with a crystal size of at least 1 micron, as described in U.S. Pat. , for example ZSM-5, which contains amorphous silica dispersed within a zeolite crystal structure. This type of catalyst is described in U.S. Patent No. 4.060.5.
No. 68 and No. 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 may be either naturally occurring or artificial, such as erionite, chabazite, zeolite T, zeolite) ZK-
5 and ZSM-54. Zeolite T is described in U.S. Pat. No. 2,95α952 as 1. Zeolite ZK-5 is described in U.S. Pat. No. 4,421,195, and ZSM-34 is described in U.S. Pat. No. 4,079,095 and 4,079,096. Since the zeolite has a pore diameter larger than 3 angstroms and smaller than 6 angstroms, it is able to approach the free space within the crystal and exit again. This zeolite is further characterized by pore windows similar to those made by eight evil monkeys, which have oxygen atoms behind them. Of course, this 8-membered ring forms the anionic skeleton of crystalline aluminosilicate.
It is created by the regular arrangement of the tetrahedrons, and the oxygen atoms themselves are bonded to silicon or aluminum atoms at the center of the tetrahedrons. This zeolite has pores with a uniform diameter of approximately 59 angstroms.
It may be substantially circular, such as SM-5, or somewhat elliptical, such as erionite, which has pores of about 56 x 5.2 angstroms. In each case, the small pore zeolites have pore diameters that are primarily smaller than 6 angstroms. Pore size of these zeolites as well as other suitable zeolites 1d Advancesin Ch@m1str
y 5eries, vol. 101, 155-170 (1
Zeolite Structures'' by Vl, M, Mayer, and IL Olson in 971).

According to the invention, the conversion reaction of the ether-rich effluent from the first step to high olefinic hydrocarbon products is carried out in a continuous multistage adiabatic reactor with interstage cooling. Cooling may be done by direct or indirect methods. 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 medium heat of conversion of the charge to high olefinic products is released.

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 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 parameters can be calculated from Equation 20: RTo2 (where θ represents the sensitivity parameter, Rq gas constant, To represents the initial temperature, and E represents the activation energy). Kel θ for the conversion of methanol to ethylene is approximately equal to 18°C.

When the exothermic chemical reaction is carried out in an adiabatic reactor (provided that x
A represents the conversion rate, ΔT represents the adiabatic temperature rise, and Δ
H represents the heat of reaction and cp represents the specific heat of the feed stream in units of specific heat of the feed stream per mole of reactant entering). Solving for ΔT, the specific conversion and the total adiabatic temperature rise for an exothermic reaction at a given feed composition can therefore be calculated from the feed specific heat and the heat of reaction.

When carried out in an adiabatic catalyst bed, the temperature rise within the bed in an exothermic reaction that is not sensitive to variations in flow rate, catalyst stage, feed composition or charge temperature must be limited to a 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 bedding required is JT total, equal to . 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 bedding.

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 paraffins and other hydrocarbons having 3 to 1 o carbon atoms are also present in low concentrations. The effluent is passed through a condenser to produce three phases: an aqueous phase containing most of the unreacted methanol, a liquid phase containing C5+ hydrocarbons, and a gaseous phase containing the produced ethylene, other light hydrocarbons, and unconverted dimethyl ether. .

Contacting methanol with a dehydration catalyst changes the equilibrium mixture.

2CH,OH,=ICIOCHm+H, (1) Zeolite catalyst converts methanol and dimethyl ether to olefin products.

2CHsOH−+CHz =CHt+2HtO(
2) CH30CH3→CH2=C&+H2O(3
) All of these reactions are exothermic, so the methanol converted in the second reaction zone generates more heat there. The methanol and DME are only partially converted to hydrocarbons (<901) in this operation, resulting in high ethylene selectivity. Once the methanol and DME conversion exceeds 90%, an increasing 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. About 30 to about 90 inches of converted grass is commonly used. Since a partial conversion operation must be performed to obtain the desired olefin product, the effluent should be recovered from the stream with a significant amount of unconverted methanol and returned to the original operating step to improve the economics of the operation. is good industrial practice. This process is used in known processes for converting low aliphatic alcohols to high olefin products. Unreacted alcohol and unreacted ether are collected separately and recycled to the operation;
Both recycle 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 reaction mixture treated in the first reaction zone is recycled to the equilibrium mixture obtained from the first reaction zone. Many operational benefits can be obtained by separating the recycled ether from the fresh charge and mixing it with the narcol-ether-water equilibrium mixture while mixing the recycled alcohol with the tack feed alcohol. As is clear from equation (1), the recycling of dialcohol and ether results in an increased conversion of alcohol to ether and water in the dehydration reactor.

Also, by directing the recycle streams to each as described above, if much of the feed is converted to olefin product, it will enter the second reaction zone as ether, which does not release as much heat as in the case of alcohol. The heat load on the area is reduced. The reaction heat generated is as follows: 1 Methanol → DME -4,92
Methanol-ethylene-2,83DME-
+Ethylene L/N-(L8 Conversion Ridge in 2 Reaction Zones The effluent from the 2 reaction zones is divided into 6 phases which are sent to a condenser. This removes substantially all of the unreacted methanol and unreacted dimethyl ether. Effectively partitioned into distinct phases. Two immiscible liquid phases, aqueous and hydrocarbon, are obtained along with a gas phase. Unreacted alcohol enters the aqueous phase and unreacted ether enters the gas phase. The aqueous phase contains a large amount of Since it contains water, most of which can be added to the fresh charge, the recovered alcohol mixed with a small amount of water is recycled after the water content of this phase has been reduced.

Methods such as evaporation and steam stripping are used to remove this water. Since the unconverted ether is 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. After the pressure is released, the gas phase can then be removed by contacting with an absorption liquid effective for selective absorption of 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 dimethyl ethyl from the absorbent before the absorbent enters methanol conversion and the DME is recycled to avoid undesirable side effects. Separation can be carried out by passing the absorbent solution containing dimethyl ether through a releasing group that can be easily separated under reduced pressure.

The invention will now be explained in more detail by way of example 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 shows four different methods of passing radial flow through the catalyst bed. FIG. 6 shows the same as FIG. 3 but with a catalyst regeneration device.

FIG. 7 is a view similar to that of FIG. 1, but with a view of the unconverted reactant separation and recycling system.

Figures 8.9 and 10 illustrate the temperature profile of the operation by plotting temperature versus catalyst bed length in the second step of the process.

First, in FIG. 1, the feed methanol containing about 50% water by weight enters the operation through line 2. This is mixed with a recycle stream rich in unreacted methanol in line 4 and a recycle stream rich in unreacted dimethyl ether (DME) in line 6.

The mixed stream is heated to a reaction temperature of 285 to 570°C in heat exchanger 8.
It is heated to , passes through fJilo, and is sent to the dehydration reactor 12.
It is now converted into an essentially equilibrium mixture of methanol, DME and water. This equilibrium mixture leaves the dehydration reactor at line 1
4 and enters a multi-stage ZSM-5 fixed bed adiabatic reactor 24 via a ridgeline 18 which is adjusted to 270-370°C in a heat exchanger 16.

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 which again increases 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. Second from the end (
The mixture leaving stage N-1) is cooled by indirect heat exchange and enters the first stage catalyst bed. Upon exiting the stage N catalyst bed, the desired conversion of methanol and DME is changed and the mixture reacts with the heated product. The body stream exits the reactor via @32, is cooled in heat exchanger 34, and enters product recovery unit 584C via #36. The cooled product and reactant mixture is separated into (1) hydrocarbon liquid phase; , (2) an aqueous phase containing most of the unreacted methanol and a small amount of DME, and (3) a gas phase containing most of the produced light olefin hydrocarbons and most of the unreacted DME. The water vapor is recycled through 114 by various separation methods such as IJ tubing, evaporation and distillation. The dough contains a higher concentration of methanol than the raw aqueous phase resulting from the dehydration process.

This recycle stream also contains most of the DME that was originally in the aqueous phase. It may not be necessary to use all of the separation methods described above to obtain the desired composition of the recirculated water stream.This treatment of the aqueous phase is produced in the conversion of hydrocarbons to methanol in the feed water and line 2 feed methanol. Most of the water that was initially present was recovered, and these liquids were removed from the system at 1/40 l/#40 to form an aqueous product containing only small amounts of methanol and DME. Become.

The gas phase remaining after condensation is compressed and sent to the absorber to remove unreacted D.
ME is collected. Suitable absorbents are methanol, feed methanol and water, each with varying efficiencies.

DME rich solution is l! It is stripped to obtain an ME-enriched stream suitable for return to either a dehydration reactor via lines 42 and 6 or a multistage fixed bed reactor via lines 42 and 44.

The DME-containing fraction from the absorber is sent via 11546 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 to recover the hydrocarbons and DME for recycling as before.

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.

FIG. 1 shows two methods for handling DME recirculation and reflux.

ing. However, in some cases, the unreacted DME in line 42 is
4 and 18, the advantages of the most advantageous operation of feeding directly into the multistage fixed bed reaction are (1) an increase in the conversion rate of methanol to DME and water in the dehydration reactor; 2) Load heat reduction in the multistage reactor due to a larger portion of the feed entering as DME, which does not generate as much heat as methanol when converted to hydrocarbons. 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 FIG.
The desired cooling can be achieved by direct cooling with (1) the feed liquid, (2) water, or (3) low pressure steam (e.g. 450 KPa operating steam). This alternative method is illustrated in Figure 2 for the multistage reaction of this system. Only the reactor is shown in Figure 1, but the other parts of the system are the same as those in Figure 1. In Figure 2, the mixture of methanol, DME and water from the Matsugi process passes through #118 and is transferred to reactor 124.
The final product enters and exits through #132. Cooling fluids obtained from various sources enter directly into the interstage direct cooling zone via line 190. The majority of the liquid obtained from the feed to the operation can be used for cooling. Charge to the dehydration reactor 12,
That is, a 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 a6 and 42.
or the condensate product from the dehydration reactor (cooled liquid from line 14). Other direct cooling fluid sources are water and low pressure (45
0KPa) This is the operating flow. The main 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 stage subsequent to the 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 contain all stages in one container, it is not possible to overlap them. 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 fn, preferably 2.5 to 4.5 m. Space velocity (WH8V based on total feed) is 0.2 due to pressure drop and economic considerations.
It should be in the range of ~2. The preferred operating temperature is ZSM-
from about 270 to about 370°C for type 5 zeolites and from about '315 to about 48°C for small pore zeolites.
The temperature is 5°C. Catalyst aging can be compensated for by increasing the inlet temperature at each stage. U.S. Pat. No. 4,025,571 and 4.
It is advantageous to use a catalyst consisting of large crystal (at least about 1 micron) ZSM-5 zeolite in a multi-bed reactor as described in No. 14a835. To reduce the hexane decomposition activity (alpha value) of the ZSM-5 type catalyst, steam treatment may be performed or preliminarily.Under the operating conditions described herein, the alpha value must be 20 or more. be. The catalyst may also be supplemented with ethylene selectivity improving components, such as intracrystalline silica, as described in U.S. Pat. Nos. 4,060,568 and 4.10 to 219. Of course, other suitable catalysts may be used, but the total rotation of methanol and DME may be limited to 30 to about 90 inches. At high conversions, even in the preferred temperature range, large amounts of methanol will be converted to aromatics. To ensure zero partial conversion, the reactor geometry and operating conditions must be operated in an unstable or sensitive reaction-like manner (i.e., operating in a sensitive manner). It is important to avoid excessive temperature rises caused by slight variations in conditions (and therefore excessive conversion of the desired light olefin product). 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, and in most cases about 18°C. The minimum number of fixed beds can be calculated from the total adiabatic temperature rise and the sensitivity parameter as 10F, with a general value of 2 to 15, and a maximum of 4 to 10. Generally, the more bedding, 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 result in the effluent from any bed being about 370°C in the ZSM-5 zeolite, or about 4°C in the small-pore zeolite.
The water content of the methanol feed to this process will be nominally 0 to about 70°C.
0% by weight, preferably from about 20 to about 50% by weight. 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 appropriate reflow reflux, recovered from the methanol and water feed and the product 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, except that in this case the design is such that the reaction mixture flows radially in the individual catalyst beds of the multistage reactor junction. This reactor geometry allows the gaseous reaction mixture to pass radially through each catalyst bed, with the catalyst within each bed being contained within a perforated annular tube.

Using such a cage, the reaction mixture flows from the annular space in the reactor through the outer surface of the perforated cage toward the catalyst body.

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 and generates an exotherm, raising the temperature of the lacrimal compound. The inner surface of the famous calligraphy 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 enters the first stage in outer ring 224 via lines 220 and 222 and flows toward the center to contact first stage catalyst bed 226 . Hydrocarbons are produced and the temperature of the reactant and product mixture 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 250 to substantially the same temperature as the first stage feed temperature. The cooled reactant and product mixture then enriches the second stage from 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 can be changed. The hot product and reactant streams exit the reactor into 1#23
2, is cooled in a heat exchanger 234, and is partially condensed.
6 and enters product recovery unit 238 for further processing as described above with respect to FIG. A method for recycling unreacted DME to the multi-stage reactor as in FIG. 1 above is described below with respect to FIG.

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 radial flow may flow inwardly from the outer surface as shown in Figures 5a and 5b, or outwardly from the center as shown in Figures 50 and 5d. Also, Figures 51 and 5C
It may be co-current or counter-current as shown in Figures 5b and 5d. The ti/Sa type is shown in Figures 3.4 and 6, but any of the five Sacred types can be used if that type is particularly good. It will be appreciated that while the flow form of Figure 5C can be applied 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, the flow shapes in Figures 51 and 5c have many unusual design or piping problems, while the flow shapes in Figures 5b and 5d require external piping to the reaction vessel to achieve the desired radial flow shape.

The zeolite catalyst used in this process must be periodically regenerated, so the catalyst is moved by gravity through a series of catalyst beds, from the last bed to Gaesong for regeneration, and then transferred to the first bed of the system. We can arrange for you to return. Figure 6 is a combination diagram of a radial flow and moving bed multi-stage reactor. Figure 6 is the same in all respects as described above for Figure 3, except for the catalyst regeneration equipment, so only the catalyst regeneration method in the figure is shown. ? Describe.

The zeolite catalyst regenerated by the regenerator 450 in FIG.
J 452 and enters the top 424 of the multistage reactor through a hopper 456 that supplies catalyst to the first stage catalyst box where it is routed with any additional feed of catalyst required entering from line 454. The catalyst moves down by gravity flow and passes through passage 45B 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 configuration, but 'cocurrent flow throughout' and similar phrases mean that the fresh feed first contacts the fresh or fresh catalyst and gradually catalyzes the catalyst through each stage. 9 is used to describe a flow in which the catalyst is gradually and continuously deactivated, and finally, at a given desired conversion, both the reaction mixture and the deactivated catalyst exit the lower end of the multistage reactor.
The catalyst stream from the stage is collected in a hopper 460 and line 462
Alternatively, the catalyst from each stage may be directly sent to the catalyst regenerator 450. In this case, the catalyst regeneration of each stage is carried out separately. Catalyst circulation in the moving bed can be continuous or intermittent, depending on the deactivation rate, depending on the specific catalyst used and the operating conditions, or can be carried out continuously or intermittently depending on the deactivation rate.
The reactor has all the advantages of a fixed bed radial flow reactor and also provides a means of continuous reaction-like operation. Fixed bed multi-stage reactors, on the other hand, require a separate regeneration 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 is the lowest concentration when the feedstock is in contact with the lowest active catalyst.
The usual advantages of co-current flow, such as the feedstock being contacted with the "highest" active catalyst, thus eliminating the single high reaction rate problems often encountered in the first bed of co-flow systems, are replaced.

Interstage indirect cooling is illustrated in FIG. However, direct interstage cooling is also possible as described above for a stationary catalyst bed in connection with FIG. In Figure 6, the catalyst activity decreases as the number of stages increases, with the first stage receiving the most active catalyst and the Nth stage receiving the least active catalyst, so the charging temperature of the second stage is usually higher than that of the first stage. charging temperature increases with the number of stages, i.e. N
The stages are operated at maximum temperature. However, for stable, insensitive, reaction-like operation, it is desirable to control the intrastage product and reactant mixture temperatures close to the charge 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 for multi-stage reaction is 380″KPa or less, preferably 275″KPa or less.
It should be less than KPm. Containing all zeolite catalyst stages within one vessel is usually economically preferred but not critical. Interstage cooling requires a liquid injection system and preferably about 1.8 to about 2.8 m with a diaphragm in the mixing zone between stages. The thickness of the radial catalyst bed is about α3 to about 1.2
m, preferably about 0.6 m. The center tube diameter should be about α3 to about 1.0 m. According to the pressure drop and economic considerations, the space velocity (on a total feed basis, WHS V
) is about α5 to about SO1, preferably about 1.5 to about 55
/ It is necessary to take the time. .

The pressure drop obtained when the reaction mixture is delivered in radial flow mode to a bed of zeolite catalyst is much lower than when operating in axial flow mode at the same amount of catalyst and the same overall conversion of reactants. In reality, the pressure drop inside the radial flow reactor is a coefficient of 3, which is the same as WH8V.
It is also about an order of magnitude smaller than when using an ordinary axial flow packed bed reactor operated in . Conversely, for a given catalyst capacity and permissible pressure drop, the output of a radial flow reactor is significantly higher than that of a conventional axial packed bed reactor. Since it is more economical to build a higher and smaller diameter reaction junction, the L/D
A radial flow large diameter fixed bed (pancake reactor) with a ratio less than 2 is preferred.

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.

Catalytic activity in the above-described moving element reaction decreases gradually from top to bottom (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 subsequent stages can be reduced, so that the feed temperature of each stage increases as the number of stages increases.The temperature in this configuration generally varies from 285°C in wJ1 stage to as high as 355°C in the maximum P stage. .

In FIG. 6, reactant and catalyst transfer are mutual, ie, fresh feed meets the most active catalyst. From reactor stability considerations, a desirable modification of this system is countercurrent operation, in which the feed stream enters from the bottom of the reactor and encounters the lowest active catalyst in the υNth stage. The reactant concentration decreases as the reactant and product streams advance upward and encounter increasingly more active catalysts.

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.

An so:5o (z amount: weight) methanol-water feed at typical operating conditions requires 6 stages for 50% conversion.

As the reaction progresses in the reactor, the reactant concentration decreases and therefore 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 proportion of desolvation amount in 6 beds is to: 1.1o:
1.21:1.32:1.49:1.78. Since the radial dimension of the catalyst cage is constant, the above ratio is equal to the height ratio of each stage. Practical limits for minimum and maximum height are determined by production capacity and feed flow failure distribution. The shaft length is about α6 to about 3 m, preferably about α9 to about 1.2 m.

The relative proportions of catalyst in each stage of a moving bed reaction-like configuration are determined by catalyst deactivation and are specific to a given catalyst.

Now in FIG. 7, the fresh methanol feed containing up to about 50% water by weight enters from line 502 and is mixed with the aqueous product recycle stream rich in unreacted methanol from line 504. This recycled flow is recovered in the process with mulberry described below. Mixed flow is line 5
The reaction temperature (285-37
0° C.) and sent via line 510 to a dehydration reactor 512 having a fixed bed 514 of gamma alumina. The mixture is converted into an essentially equilibrium mixture of methanol, dimethyl ether (DME) and water by the catalytic action of alumina, line 5.
Exit the reactor at 16. The recycle stream of unreacted DME recovered from the absorption-desorption operation described below passes through line 518 and is mixed with the equilibrium mixture from line 516 in #52°. The resulting mixture is heated or cooled as needed in a heat exchanger 522 to a reaction temperature (270-370'C) suitable for the second reaction zone, and then the mixture passes through line 524 to a reactor 526 with a ZSM-5 catalyst. flows inside. 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.

The 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 C5+ hydrocarbons are condensed. The condensed mixture and the non-condensed gas pass through Gen 536 to separator 5.
At 38, the mixture is divided 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 a7 dimethyl ether a1 water 7αOCO,co;
, H, total < α1 methane
α1 Ethane α1 Ethylene 53 Propane α6 Propylene 2.1 Isobutane
α6 n-butane a2 Butenes 1.0Cs+
5.1 The bottom layer in separator 538 is an aqueous layer consisting of the water produced by the conversion of methanol to hydrocarbons, the water contained in the feed, and unreacted methanol and a small amount of dimethyl ether. The second layer between the water and gas layers is C,+ 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 undesired amount of water in this aqueous phase, a portion of it is passed through line 542 to steam stripper 544.

S) The steam introduced into the IJ tube 544 removes unreacted methanol and a small amount of DME from the aqueous phase portion flowing down through the stripper. Separated product is #546
It joins the remaining aqueous phase portion in line 504 and is sent to line 502 where it is mixed with freshly supplied methanol as described above. The water from which unreacted methanol and DME have been separated is discarded via line 548 from IJ tube 544. 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 methanol and DME are present in the amount of aqueous phase directed to the stripper.
The amount of water resulting from the conversion to hydrocarbons should be disposed of through line 548. Operational variables such as the nature of the absorption liquid used for DME recovery may dictate more or less aqueous phase stripping; this stripping operation can be adjusted as the operation and product require.

The liquid hydrocarbon phase containing C,+ hydrocarbons is removed from separator 538 via line 550 and sent to subsequent equipment for recovery or further processing. For example, this portion can be used as a useful portion for gasoline blending.

The third phase in separator 538 is a gaseous phase containing substantially all of the ethylene produced, as well as light hydrocarbons, trace amounts of Co, Co, , H2, and a relatively large amount of dimethyl ether.

This gas phase exits through 1s 552, is compressed to about 1500 KPa by compressor 554, and is sent via line 556 to a gas absorption column 558 having each bed 560 packed with, for example, Bahr saddles, Raschig terms, or Paul rings. #562
Water is introduced into absorption column 558 and the water flows countercurrently with the gas ascending the column, effectively removing dimethyl ether from the gas. DME recovered by dissolving in water is transferred from tower 558 #
The dimethyl ether exits at 564 and enters a releasing group 566 where vacuum is applied to separate the dimethyl ether from the water. The water that has passed through the releasing group 566 is recycled through the condenser 562 to the absorption tower 558 for reuse. If necessary, water is added to the absorption system via condenser 568.

The separated dimethyl ether is traced from the releasing group 566 to the line 51
8 to line 520 where the ethylene enrichment is combined with the equilibrium mixture of methanol, dimethyl ether and water exiting dehydration reactor 512 for conversion to product.

The partially purified sinter stream exits absorption tower 558 via line 570 and is sent to gas plant 572. Ethylene of the desired purity is recovered from the gas plant via line 574 and heavy hydrocarbons such as propylene and butenes exit the gas plant via line 576. The gas factory also has C01C02, H2,
The following example illustrates the invention.

Example 1 The reaction of conversion of methanol to olefinic hydrocarbons is highly exothermic and, unlike ammonia synthesis or methanol production, there is no thermodynamic equilibrium limit and therefore high temperatures are required for the conversion reaction in a given catalyst bed. Or do not limit temperature rise.

In a non-temperature sensitive reaction where variations in flow rates, catalyst conditions, feed composition and charge temperature can be controlled, the temperature rise in each bed should be limited to approximately the sensitivity parameter value, ie about 18°C.

The table below summarizes the influence of this ΔT/bed that is allowable on the reaction contact performance.The considered condition is that the dehydration reactor is charged with feed of 50% by weight of water and 50% by weight of methanol and then converted. Stable 50% total conversion with 4.6 and 8 stage beds of zeolite catalyst in the reactor. The corresponding temperature increases per bed were 27.18 and 156°C, respectively. Using the above equation, the total iIs thermal temperature rise for olefin production from 50% conversion of this methanol feed is about 108°C. This means that if all the catalyst were in the cold bed, the outlet temperature would be about 401°C for an initial feed charge temperature of 293°C.

Such temperature increases may be desirable since they can lead to operational instability and the production of extenuating amounts of high-boiling products; furthermore, the gradual increase in feed charge temperature required to compensate for catalyst aging further reduces this instability. Table 1 shows multi-bed and indirect interstage cooling, outlet temperature 320℃ (4 stages), 311℃ (6 stages)
It is shown that at 307° C. (8 stages), olefin-based products can be converted into aromatic compounds.

Table 1 Sensitivity to temperature fluctuations Feed: 50:50 (WOW) Methanol: Water conversion: 50 CH Charge temperature: 293°C Coolant temperature: 288°C Indirect interstage cooling 4 stages 6 stages 8 stages 417 stages, 'C27181 & 6 Conversion rate 5℃ fluctuation 61 53 517℃ fluctuation 75 59 5510℃ fluctuation
100 65 55 The more stages of use (correspondingly the smaller the temperature rise per bed), the more stable the reaction contact performance.
It is shown by giving the variation in °C. Interstage cooling was indirect and the coolant in this indirect heat exchanger was maintained at 288°C. As shown in Table I, the stability of the 8-stage type is not significantly greater than that of the 6-stage type, but it is considerably thicker! Since it requires an investment in power, a 6-stage type, ie, a temperature rise of 18°C/floor, is optimal. 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 conditions of the four-stage reactor shown in Table I with a 7° C. fluctuation. The "X" temperature condition is seen at the reference condition (variation fi). This figure shows that a 7°C fluctuation results in an excessive temperature rise (temperature of 370°C or higher) in the first stage and the production of a large amount of unnecessary products.

Example 1 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.

Table 1 summarizes the minimum number of stages of fiZsM-5 zeolite catalyst required for moderate feed compositions at conversions of 50 and 80 inches. The feed composition in this table is expressed as weight percent methanol in water entering the dehydration reactor to provide an equilibrium mixture of methanol, DME and water entering the multi-bed zeolite catalyst system.

Table ■ 16 2
450 6
1084 8
15100 9
17 Six stages are required for a 50150 (wt%) methanol water feed at a typical 50% conversion. As the reaction progresses in the reactor, the reactant concentration decreases and therefore the amount of catalyst at each stage increases progressively in the direction of reactant flow. (Relative ratio of 6 beds with indirect interstage cooling in the above example with constant rise in hot water temperature per bed is 1.0: 1
．． 10: 1.21: 1.32: 1.44: 1
．． It becomes 78. 6 (Typical temperature conditions of the machine are shown in Figure 8).

Note the uniform temperature rise per bed, the interstage cooling of the reaction mixture back to the first stage charging temperature, and the increase in catalyst bed thickness as the reaction mixture passes through the series of M6 stages. The 100 steps to obtain the conversion rate are as follows: 1.0
: 1.09:1.20:t33:1.5' knee 1.7
2:2.00:2.41 = 500:598, 10 beds would be required.

Various variations of interstage cooling as described above are possible; Table M 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.

2 2.2
1.5 12.03 2
．． 3 1.5 15,
44 2.4
1.6 15,15
2.5 1.7 16
,96 2.6
1.8 189 Examples ■ This example shows the effects of a high water content feed, except that the feed-composition is the same as that of the four-stage system of Example ■. 67155 (heavy Il) Hl O to the reactor
/ Me O) (shows the effects of water dilution and temperature fluctuations when charging the feed. If the total conversion rate of this system is 50% under stable operating conditions, the temperature rise in the bed plate will increase again. The temperature state of this system is shown in Figure 10. The "X" curve is for the same basic temperature rise in each of the four beds. The temperature curve obtained for a 10° C. fluctuation is also shown in this figure. It is noteworthy that the starting temperature of this four-stage scheme is lower than that of the six-stage scheme of FIG.

FIG. 10 shows that the first stage outflow temperature is significantly higher than in the case where the 10°C fluctuation is basic, but it is much lower than the first stage outflow temperature obtained with the 7°C fluctuation in the four-stage system of Example I. ing. In fact, with high water content feeds 10°C
The fluctuations do not allow the first stage effluent temperature to exceed the limit of 370°C.

A comparison of FIG. 2 and FIG. 10 shows that the high water content methanol feedstock can handle larger fluctuations in the first stage charging temperature of a continuous bed system, while avoiding high temperatures that tend to cause side reactions.

Examples ■ The pressure drop of the reactant fluid across the catalyst bed can be calculated using the Ergin equation: Applying the ideal gas law to the density change due to pressure drop and integrating the above equation (where G = mass flow rate, φ = shape factor, = catalyst particle diameter, P = gas density gc = constant (
= Porosity μ = Viscosity Po = = Charging pressure M = Molecular weight
T = average temperature of gas R = gas constant
2 = 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 stations was the same methanol-water mixture and used for the following run: WH8V 3/hr (overall) Charge temperature 293°C Porosity α38 Fixed filling height of 3 ft. The pressure drop of the bed reactor is 2.76 KPa, but it has the same capacity and the thickness C
The pressure drop in a 15 m radial flow reactor is [1 L21 KPa.

From the above results, if methanol is converted to olefins using a radial flow reactor, the pressure drop seen in an axial flow reactor can be reduced.
It is clear that it can be made several orders of magnitude smaller.

Example ■ Reaction: 20H8OH, :: CH30CH3+ H
IO is an equilibrium limited reaction. From the Le Chaterie principle, addition of dimethyl ether to methyl alcohol reduces the conversion.

This effect is demonstrated by introducing two different feed streams into a dehydration reactor at a temperature of 315° C. and a pressure of 510 kPa.

MeOH50F 1.56 0.45D
ME Of (LO[156H,05(12
,78533 MeOH50f 178 [1
47DME 25f α54
0.70H, o 501 2.78
2.94 MeOH conversion rate = 1 - 0 Output C □ α45 In the case of ■ =' t s b ''' 71.2
For α47 1 = 1 - cL78 = 397 This example shows that the addition of dimethyl ether to the feed stream entering a dehydration reactor reduces the methyl alcohol conversion made in the reactor.

[Brief explanation of the drawing]

FIG. 1 is a schematic illustration of the implementation of the method of the invention. FIG. 2 is a schematic diagram of the present invention using the same method as FIG. 1 but using an alternative method for interstage cooling. FIG. 3 is a diagram of the same method as FIG. 1 of the present invention, but with 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 8, 16, and S4 in each of the figures above. 408, 416, 434 12.212.412 Dehydration reactor 24.22
4.424 Multi-stage reactor 226.426
Catalyst cage 230.430 Interstage cooling device 38.238.438 Product recovery device 450 Catalyst regenerator Figure 7
1 is a schematic diagram of the same method as in FIG. 1 of the present invention, but including separation and recycling of unconverted reactants; FIG. In the figure, 512 dehydration reactor, 526 multistage reactor, 538 separator, 544 stripper, 554 compressor, 558 absorption tower, 566
releasing group, 572 gas factory. FIG. 8 is a curve diagram showing the temperature characteristics in the multi-stage adiabatic reactor of the present invention, with the horizontal axis representing the bed length ratio and the vertical axis representing the temperature ('C). Figure 9Fi 50:50 H2O:MeO of the present invention
It is a temperature characteristic diagram with H charging, 7°C fluctuation in a 4-stage system, and indirect interstage cooling. iIA axis vertical axis upper axis 8. Figure 10 shows the 67:33 H2O:MeO of the present invention.
It is a temperature characteristic diagram with H charging, 10° C. fluctuation in a 4-stage system, and indirect inter-stage cooling. The horizontal and vertical axes are as shown in FIG. Patent applicant: Mobile Oil Corporation
＼ FIG, 1 FIG, 2 FIG, 5 FIG, 6 Continued from page 1 Priority claim @ February 5, 1982 [phase] United States (US)
■346426 @ February 5, 1982 ■United States (US) ■345984 158

Claims (1)

Claims: 1. After contacting a methanol-containing feed with a dehydration catalyst at elevated temperatures to produce an ether-rich product, said product is then subjected to (i) a pore size greater than 5 angstroms, at least 12 a silica to alumina molar ratio and a constraint index of 1 to 12, or (i) a pore window size defined by an 8-membered ring with a majority of pore sizes smaller than 6 angstroms and an oxygen atom; contacting the ether-rich product with the catalyst at elevated temperatures through alternating passes through a catalyst bed of crystalline aluminosilicate zeolite and a cooling zone to achieve a desired conversion of the olefinic hydrocarbon product; The temperature drop caused by the reaction mixture passing through each cooling zone is equivalent to the temperature rise in the immediately preceding catalyst bed! be qualitatively equal, the increase in salinity due to passage of the reaction mixture through each catalyst bed shall not be greater than the sensitivity parameter of the conversion of methanol to ethylene, and the number of catalyst beds shall be equal to the sensitivity of the total adiabatic temperature increase at a given conversion. A process for the conversion of methanol to olefinic hydrocarbons, characterized in that the ratio of methanol to auxiliary variables is at least equal to: 2. The method of claim 1, wherein the λ reaction mixture passes axially through each zeolite catalyst bed. 5. The method according to 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. & A method according to claim 5, wherein the reaction mixture flow is cocurrent with the catalyst flow. A method according to claim 5, wherein the reaction mixture flow is countercurrent with the catalyst flow. a 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.