CN101440014A - Method for producing light olefins - Google Patents

Abstract

A method for producing low-carbon olefins. The heavy raw material is introduced into the first reactor, and the regenerated catalyst is contacted and reacted. After the reaction, the oil gas and the catalyst enter the second reactor to continue the reaction; the light raw material is introduced into the third reactor, and the Regenerated catalyst contact reaction, after the oil gas discharged from the third reactor is separated from the raw catalyst, the separated raw catalyst is introduced into the second reactor; after the oil gas and the raw catalyst after the reaction in the second reactor are separated, the separated The raw catalyst is regenerated by burning and then returned to the first reactor and the third reactor for recycling. The oil and gas separated from the second and third reaction zones are introduced into the catalytic cracking separation and stabilization system, and the target product, low-carbon olefins, is obtained through fractionation. The method provided by the invention can flexibly adjust the operating conditions of light raw material re-smelting, optimize product distribution, increase the yield of low-carbon olefins; reduce the hydrogen transfer reaction activity in the second reactor, and increase the yield of low-carbon olefins.

Description

A method for producing low-carbon olefins

technical field

The present invention relates to a method for the catalytic conversion of hydrocarbon oils in the absence of hydrogen, and more particularly relates to a method for converting heavy raw materials into low-carbon olefins such as propylene and ethylene.

Background technique

Low-carbon olefins such as ethylene and propylene are important organic chemical raw materials, among which propylene is a synthetic monomer for products such as polypropylene and acrylonitrile. With the rapid growth of demand for derivatives such as polypropylene, the demand for propylene is also increasing year by year. The demand in the world propylene market has increased from 15.2 million tons 20 years ago to 51.2 million tons in 2000, with an average annual growth rate of 6.3%. It is estimated that the demand for propylene will reach 86 million tons by 2010, with an average annual growth rate of about 5.6%.

The methods for producing propylene are mainly steam cracking and catalytic cracking (FCC), wherein steam cracking uses naphtha and other light oils as raw materials to produce ethylene and propylene through thermal cracking, but the yield of propylene is only about 15% by weight, while FCC Heavy oil such as vacuum gas oil (VGO) is used as raw material. At present, 66% of the world's propylene comes from the by-products of steam cracking to produce ethylene, 32% comes from the by-products of gasoline and diesel produced by FCC in refineries, and a small amount (about 2%) is obtained from propane dehydrogenation and ethylene-butene metathesis.

If the petrochemical industry follows the traditional route of steam cracking to produce ethylene and propylene, it will face several major constraints such as shortage of light raw material oil, insufficient production capacity and high cost.

Catalytic cracking has been paid more and more attention due to its advantages such as wide adaptability of raw materials and flexible operation. In the United States, almost 50% of the propylene market demand comes from FCC units. The improvement technology of catalytic cracking to increase the production of propylene is developing rapidly.

US 6045690 discloses a method for producing low-carbon olefins. In this method, the regenerated catalyst is divided into two or more streams and enters a downflow reactor, but this method cannot solve the problem of uneven temperature and density distribution in the reactor.

US 4257875 discloses a two-stage catalytic conversion method, which divides the regenerated catalyst into two streams and enters two riser reactors respectively, but this method also cannot solve the problem of uneven temperature and density distribution of the reactors.

US 4578183 discloses a catalytic conversion method, which divides the regenerated catalyst into two strands in the riser, one strand enters the first reactor, and the second strand enters the second reactor, but the method also has relatively low dry gas production rate high question.

CN1898362A discloses a method for producing low-carbon olefins and aromatics. The raw materials are contacted with catalytic cracking catalysts and reacted at a reaction temperature of 400-800°C and a weight hourly space velocity of 0.1-750h ^-1 to separate unborn catalysts and reaction oil and gas , the catalyst to be used is returned to the reactor after being regenerated, the reaction is carried out in at least two reaction zones, the reaction temperature of at least one reaction zone in the reaction zone downstream of the first reaction zone is higher than the reaction temperature of the first reaction zone, and Its weight hourly space velocity is lower than that of the first reaction zone. The method introduces a regenerated catalyst with high temperature and high activity in the second reaction zone, so that the hydrogen transfer reaction in the second reaction zone is relatively large, and the yield of low-carbon olefins is reduced.

Contents of the invention

The purpose of the present invention is to provide a method for producing low-carbon olefins from heavy raw materials with a higher yield of low-carbon olefins on the basis of the prior art.

The method for producing light olefins provided by the invention comprises: introducing the heavy raw material into the first reactor, contacting and reacting with the catalytic cracking catalyst from the regenerator, after the reaction, the oil gas and the catalyst enter the second reactor to continue the reaction; The high-quality raw material is introduced into the third reactor and reacted with the catalytic cracking catalyst from the regenerator. After the reaction in the third reactor, the oil gas and the raw catalyst are separated by gas-solid separation equipment, and the separated raw catalyst is introduced into the second reactor. ; After the oil and gas reacted in the second reactor and the catalyst to be used are separated, the separated catalyst to be used is returned to the first reactor and the third reactor for recycling after being burnt and regenerated in the regenerator. The oil and gas separated from the reactor and the third reactor are introduced into the catalytic cracking separation and stabilization system, and the target product light olefins are obtained through fractionation; the reaction temperature of the second reactor is higher than that of the first reactor, and the weight hourly space velocity is lower than that of the first reactor. The weight hourly space velocity of the first reactor, the reaction temperature of the third reactor is higher than the reaction temperature of the second reactor.

The advantage of the method provided by the invention is:

The method provided by the invention can flexibly adjust the operating conditions of the third reactor, optimize the product distribution of light raw material re-refining, increase the yield of low-carbon olefins; reduce the hydrogen transfer reaction activity in the second reactor, and further increase the low-carbon The yield of olefins, especially the yield of propylene is increased.

Description of drawings

Fig. 1 is the schematic flow sheet of the method provided by the present invention;

Fig. 2 is the schematic flow sheet of the method for producing more light olefins adopted in the comparative example.

Detailed ways

Method provided by the invention is implemented like this:

The heavy raw material is introduced into the first reactor and contacted with the catalytic cracking catalyst from the regenerator at a reaction temperature of 500-650°C, preferably 540-600°C, and a weight hourly space velocity of 0.1-750h ^-1 , preferably 1-500h ^-1 , the reaction pressure is 0.10～1.0MPa (absolute pressure), the weight ratio of catalytic cracking catalyst and raw material is 2～100, reacts under the condition of preferred 5～50, after reaction oil gas and catalyst enter the second reactor to continue reaction, The reaction temperature of the second reactor is 10-100°C higher than that of the first reactor, preferably 20-60°C higher, and the ratio of the weight hourly space velocity of the second reactor to the first reactor is 1: (1.1-750), preferably 1: (1.1～300); introduce the light raw material into the third reactor, contact with the catalytic cracking catalyst from the regenerator, the reaction temperature is 600～750°C, preferably 650～700°C, and the residence time is 1～20 seconds, preferably 1 to 10 seconds, the reaction pressure is 0.10 to 1.0 MPa (absolute pressure), the weight ratio of catalyst to raw material is 4 to 100, preferably 5 to 80, and the reaction temperature of the third reactor is higher than that of the second reactor. The reaction temperature of the reactor, the oil gas and the raw catalyst after the reaction in the third reactor are separated by gas-solid separation equipment, and the separated raw catalyst is introduced into the second reactor; the oil gas and the raw catalyst after the reaction of the second reactor are passed through After separation, the separated spent catalyst is burned and regenerated in the regenerator to restore its activity, and then returned to the first reactor and the third reactor for recycling, and the oil and gas separated from the second reactor and the third reactor are introduced into the catalytic cracking separation Stabilize the system, and obtain the target product low-carbon olefins through fractional distillation.

In the method provided by the invention, the heavy raw material is selected from vacuum wax oil, atmospheric gas oil, coking wax oil, deasphalted oil, vacuum residue, atmospheric residue, re-refined oil, oil slurry, diesel oil one or a mixture of several. Among them, vacuum gas oil, atmospheric gas oil, coking gas oil, deasphalted oil, vacuum residual oil, atmospheric residual oil, and diesel oil are whole fractions or partial fractions without hydrogenation, or whole fractions after hydrogenation or partial fractions. When two or more raw materials enter the reactor, they may enter the reactor at the same position or at different positions.

In the method provided by the invention, the light raw material is selected from gasoline and/or C4-C8 hydrocarbon mixtures. Among them, C4-C8 hydrocarbon mixtures are preferred.

Described gasoline is selected from the catalytic cracking gasoline that obtains in the method provided by the present invention, also can be from the catalytic cracking gasoline outside this device, straight-run gasoline, coker gasoline, thermal cracking gasoline, thermal cracking gasoline, hydrogenated gasoline One or a mixture of more than one of them.

In the method provided by the invention, the catalyst active component is selected from medium-pore zeolite and optional large-pore zeolite, and the large-pore zeolite is Y-type or HY-type zeolite containing or not containing rare earth, or zeolite containing or not containing rare earth. The ultra-stable Y-type zeolite and the medium-porous zeolite are one or more of zeolites with MFI structure or high-silica zeolites with five-membered ring structure prepared by other methods.

The catalyst is composed of zeolite, inorganic oxide and optional clay, wherein based on the total weight of the catalyst, it contains 10-50w% of zeolite, 5-90w% of inorganic oxide and 0-70w% of clay.

The inorganic oxide as a binder is selected from silicon dioxide (SiO ₂ ) and/or aluminum oxide (Al ₂ O ₃ ).

The clay as a matrix (ie carrier) is selected from kaolin and/or Halloysite.

In the method provided by the present invention, the first reactor is a riser, an ascending conveying line, a descending conveying line or a fluidized bed reactor, preferably a riser reactor or a fluidized bed reactor, more preferably a riser reaction device; the second reactor is a fluidized bed reactor; the third reactor is a riser reactor. The superficial linear velocity of the fluidized bed reactor is 0.2-2.5 m/s, and the catalyst density is 150-700 kg/m ³ .

In order to realize fluidized operation, a lifting medium may be injected at the bottom of the reaction zone, the lifting medium is selected from steam or dry gas, preferably steam. The weight ratio of water vapor to raw material is 0.05-1.0.

In the method provided by the invention, the first reactor and the second reactor are connected in series, the second reactor is downstream of the first reactor, and the reaction temperature of the second reactor is higher than the reaction temperature of the first reactor. 10-100°C, preferably 20-60°C. According to the common sense of those skilled in the art, the reaction temperature of a tubular reactor such as a riser refers to the outlet temperature, and the reaction temperature of a bed reactor such as a fluidized bed refers to the average temperature of the bed. The ratio of the weight hourly space velocity of the second reactor to the weight hourly space velocity of the first reactor is 1:(1.1-750), preferably 1:(1.1-300).

In the method provided by the present invention, the third reactor is connected in parallel with the first reactor and the second reactor. The spent catalyst and reaction oil gas at the top of the third reactor are separated by conventional settling and cyclone separators, and the spent catalyst separated from the top of the third reactor enters the bottom or middle of the second reactor to provide heat for the second reactor and catalyst; the separated oil and gas enters the catalytic cracking fractionation stabilization system.

The spent catalyst separated from the top of the second reactor returns to the reactor after stripping or non-stripping and coke regeneration; the separated oil gas enters the catalytic cracking fractionation stabilization system. After fractionation and passing through other separation systems, low-carbon olefin products are obtained. The low-carbon olefins are ethylene, propylene and optional butene, that is, the low-carbon olefins are ethylene, propylene, or ethylene, propylene and butene.

The method for separating ethylene from reaction oil gas is the same as the method for separating ethylene from catalytic cracking dry gas well known to those of ordinary skill in the art, and the method for separating propylene and optional butene from reaction oil gas is the same as that known to those of ordinary skill in the art The separation of propylene and optionally butenes from catalytically cracked liquefied gas is the same. The method for separating aromatics from the pyrolysis gasoline fraction of reaction oil and gas is the same as the method for separating aromatics from steam cracking gasoline known to those skilled in the art, that is, solvent extraction. Before separating aromatics from the catalytic cracking gasoline obtained by the method, the C ₅ -C ₈ in gasoline are separated first and used as recycle material.

In the method provided by the invention, the spent catalyst is introduced into the regenerator, and reacts with an oxygen-containing medium to burn off all or most of the coke on the spent catalyst, so that the activity of the catalyst can be recovered. The regenerated catalyst regenerated by burning is returned to the first reactor and the third reactor for recycling. There may also be some incompletely burned coke on the regenerated catalyst. This patent does not limit the amount of carbon on the to-be-regenerated catalyst, as long as the regenerated catalyst can have appropriate activity.

The catalyst regeneration conditions of the regenerator are as follows: the temperature is 600-800° C., the pressure is 0.1-0.6 MPa (absolute pressure), and the residence time is 60-720 seconds.

The beneficial effects of the method provided by the invention are:

Since gasoline and C4-C8 hydrocarbons are separately recuperated in the third reactor, the operating conditions can be flexibly adjusted and optimized to increase the yield of propylene. Simultaneously because the spent agent after the reaction of the third reactor is introduced into the second reactor to continue the reaction, compared with the regenerated catalyst introduced from the regenerator, the catalyst activity is reduced, and when the temperature of the second reactor is increased, The hydrogen transfer reaction activity in the second reactor is reduced, which can further increase the yield of light olefins, especially the yield of propylene.

The method provided by the present invention will be further described below in conjunction with the accompanying drawings, but the present invention is not limited thereby.

Fig. 1 is a schematic flow chart of the method provided by this patent. As shown in Figure 1, the first reactor A is a riser reactor, the second reactor B is a fluidized bed reactor, the first reactor and the second reactor are connected in series, and the third reactor C is a riser reactor , the third reactor C is connected in parallel with the first reactor A and the second reactor B.

Part of the regenerated catalyst enters the bottom of the pre-lift section 3 at the lower part of the first reactor A through the pipeline 1 and the control valve 2, and the pre-lift steam enters the bottom of the pre-lift section 3 through the pipeline 4, and is accelerated along the riser under the lifting effect of the pre-lift medium Movement: The raw oil is injected into the riser A through the pipeline 5 together with the atomized steam from the pipeline 6, contacts with the regenerated catalyst and reacts upward along the riser A. The oil and gas from the outlet of the riser A enter the second reactor B together with the catalyst, and mix with the catalyst coming from the line 8 at the bottom of the fluidized bed B and react in the second reactor B. The second reactor can also be supplemented with regenerated catalyst via line 27.

Another part of the regenerated catalyst enters the bottom of the pre-lift section 11 of the third reactor C through the pipeline 9 and the control valve 10, and the pre-lift steam enters the bottom of the pre-lift section 11 through the pipeline 14, and is accelerated upward along the riser under the lifting effect of the pre-lift medium Movement: The C4-C8 hydrocarbon mixture of this device is injected into the riser C through the pipeline 12 together with the atomized steam from the pipeline 13, and reacts in the third reactor C in contact with the regenerated catalyst. The oil gas and catalyst at the outlet of reactor C enter the cyclone 7 through the pipeline 15, and the separated catalyst enters the second reactor B through the dipleg of the cyclone and the pipeline 8 to provide catalyst and heat for it; the separated oil gas passes through the pipe first 18 Enter settler D.

The oil gas and catalyst at the outlet of the second reactor B also enter the settler D through the pipeline 16. After being separated by the cyclone, there is almost no catalyst in the oil gas. The oil gas enters the separation system F at the rear through the pipeline 19 and is separated in the separation system F. Ethylene, propylene, C4-C8 hydrocarbons, gasoline, diesel oil and oil slurry are output. The oil slurry is fully re-refined through pipeline 5, part of diesel oil is also re-refined through pipeline 5, and part of C4-C8 hydrocarbons enters the third reactor through pipeline 12. Reaction in C. The catalyst separated by the settler D is contacted with the water vapor from the pipeline 20 in the stripping section 28, and the oil gas carried by the catalyst is stripped out, and the stripped carbon-bearing catalyst (called spent agent) passes through the pipeline 21 and the control valve 22 Enter regenerator E.

In the regenerator E, the spent catalyst is in contact with the air coming through the pipeline 23 and the air distributor 24, and the coke on the spent catalyst is burned, and the activity of the catalyst is restored, and at the same time, the coke releases a large amount of heat for the reaction. The flue gas carrying the catalyst is separated from the catalyst through the cyclone 25, and the clean flue gas enters the flue gas energy recovery system through the pipeline 26 to recover the heat of the flue gas. The regenerated catalyst enters the reaction system through pipeline 1 and pipeline 9 respectively.

The following examples will further illustrate this method, but do not limit the present invention thereby.

The heavy raw material used in the examples and comparative examples is vacuum wax oil F1, and its properties are shown in Table 1; the light raw material used is F2, and its properties are shown in Table 3. The catalyst used is MMC-2 catalyst, which is produced by Qilu Catalyst Factory of Sinopec Co., Ltd., and the properties of the catalyst are shown in Table 2.

comparative example

The comparative example illustrates the effect of the method for producing light olefins and aromatics disclosed in CN1898362 in producing light olefins.

The flow process of the comparative example is shown in Figure 2. In the medium-scale test device, the first reactor A and the second reactor B are connected in series, and the regenerated catalyst from the regenerator E is introduced into the first reactor A through the pipe 1 and the control valve 2. The bottom of the pre-lift section accelerates upward along the riser reactor under the lift of the pre-lift medium introduced through the pipeline 4, and the raw material oil F1 is injected into the riser A through the pipeline 5 and the atomized steam from the pipeline 6, and the regenerated catalyst contact reaction. The oil gas and catalyst at the outlet of the first reactor enter the second reactor B to continue the reaction, while the regenerant in the regenerator E is replenished to the bottom of the second reactor through the pipeline 27 . The oil gas and catalyst after the reaction in the second reactor enter the settler D through the pipeline 16 for separation, and the separated spent agent enters the regenerator E through the pipe 21 first to be burnt and regenerated; the separated oil gas enters the split flow stabilization system through the pipeline 19 F is separated, and the product that separates obtains is measured product distribution through gas chromatography. The reaction temperature of the first reactor is 600°C, the reaction temperature of the second reactor is 580°C, the weight hourly space velocity of the first reactor is 180hr ^-1 , and the weight hourly space velocity of the second reactor is 3hr ^-1 . The operating conditions and product distribution of the first and second reactors are shown in Table 4.

Example

Examples illustrate the effect of the method provided by the invention in producing light olefins.

As shown in Figure 1, the test was carried out in a medium-scale test device, the first reactor and the second reactor were connected in series, and the third reactor was connected in parallel with the first reactor and the second reactor. There is no diesel oil and oil slurry re-refining, the raw material oil F1 is used as the raw material of the first reactor, and the C4 re-refined raw material F2 is used as the raw material of the third reactor. Among them, the reaction temperature of the first reactor is 560°C, and the weight hourly space velocity is 180hr ^-1 ; the reaction temperature of the second reactor is 600°C, and the weight hourly space velocity is 3hr ^-1 , the distance between the second reactor and the first reactor The ratio of gravity to space velocity is 1:60. The reaction temperature of the third reactor is 680° C., and the supplementary heat for the second reactor accounts for 24% of the reaction heat of the whole reaction system. After the products are separated, only carbon tetraolefins are recycled to the third reactor, and other recycled materials are not recycled. The operating conditions and product distribution of the first, second and third reactors are shown in Table 4.

Table 1 Raw oil

Table 2 Catalyst

catalyst C1 Chemical composition, w% RE ₂ O ₃ 0.56 Al ₂ O ₃ 54.00 physical properties Specific surface, m ² /g 120 Pore volume, cm ³ /g 0.17 Apparent density, g/ ^cm3 0.91 Sieve, w% 0～20μm 0.8 0～40μm 10.4 0～80μm 70.8 0～110μm 88.5 0～149μm 97.8 >149μm 2.2 APS, μm 64.3 Microreactive activity, w% 60

Table 3 Composition of raw material F2 carbon four olefins

composition w% Butene-1 15.53 Isobutylene 43.67 trans-butene 24.02 cis-butene 16.78

Table 4 Operating conditions and product distribution

project comparative example Example catalyst C1 C1 The third reactor raw material - F2 First Reactor Raw Materials F1 F1 Feed amount of the first reactor, kg/h 6 6 Reaction pressure of the first reactor, MPa (absolute pressure) 0.22 0.22 Average temperature of the first reactor, °C 580 560 Ratio of agent to oil in the first reactor, w/w 12 10 Space velocity of the first reactor, hr ^-1 180 180 The amount of atomized steam in the first reactor, kg/hr 1.8 1.8 The preheating temperature of the raw material in the first reactor, ℃ 330 330 Average temperature of the second reactor, °C 600 600 Agent to oil ratio in the second reactor, w/w 17 17 Second reactor WHSV space velocity, hr ^-1 3 3 The average temperature of the third reactor, ℃ - 680 Feed amount of C4 in the third reactor back refining, kg/h - 1.2 Atomized steam volume of carbon four in the third reactor, kg/h - 0.12 The preheating temperature of raw materials in the third reactor, ℃ - 150 Agent-oil ratio of the third reactor, w/w - 35 Space velocity of the third reactor WHSV, hr ^-1 - 3.5 Product distribution, w% dry gas 12.3 13.2 liquefied gas 48.1 42.8 gasoline 18.5 22.1 diesel fuel 9.5 9.7 oil slurry 2.1 2.1 coke 9.5 10.1 total 100.0 100.0 Ethylene yield, w% 7.5 8.3 Propylene yield, w% 24.1 28.4

As can be seen from Table 4, compared with the comparative example, the propylene and ethylene yields of the method provided by the present invention are 28.4w% and 8.3w%, respectively increased by 4.3 percentage points and 0.8 percentage points.

Claims (14)

1. A method for producing light olefins, characterized in that it comprises: introducing heavy feedstock into the first reactor for catalytic cracking, contacting and reacting with the catalytic cracking catalyst from the regenerator, and the oil gas and catalyst after the reaction enter the second reactor Continue the reaction; introduce the light raw material into the third reactor, and contact with the catalytic cracking catalyst from the regenerator for reaction, the oil gas and unborn catalyst after the reaction in the third reactor are separated by gas-solid separation equipment, and the separated unborn catalyst Introduce into the second reactor; after the reaction of the oil and gas in the second reactor and the spent catalyst are separated, the separated spent catalyst is burned and regenerated in the regenerator and then returned to the first reactor and the third reactor for circulation Use, the oil and gas separated by the second reactor and the third reactor are introduced into the catalytic cracking separation and stabilization system, and the target product low-carbon olefins are obtained through fractionation; wherein the reaction temperature of the second reactor is higher than the reaction temperature of the first reactor, heavy The hourly space velocity is lower than the weight hourly space velocity of the first reactor, and the reaction temperature of the third reactor is higher than that of the second reactor. 2. The method according to claim 1, characterized in that the reaction conditions in the first reactor are as follows: the reaction temperature is 500-650°C, the weight hourly space velocity is 0.1-750h ^-1 , and the reaction pressure is 0.10-1.0MPa (absolute pressure), the weight ratio of catalytic cracking catalyst and raw material is 2～100; The ratio of the weight hourly space velocity of the device is 1:(1.1～750). . 3. The method according to claim 2, characterized in that the reaction conditions of the first reactor are that the reaction temperature is 540-600°C, the weight hourly space velocity is 1-500h ^-1 , and the weight ratio of catalytic cracking catalyst to raw material is 5 ~50. 4. The method according to claim 2, characterized in that the reaction temperature of the second reactor is 20-60°C higher than the reaction temperature of the first reactor. The ratio of the weight hourly space velocity of the second reactor to the weight hourly space velocity of the first reactor is 1:(1.1-300). 5. According to the method of claim 1, it is characterized in that the operating conditions of the third reactor are as follows: the reaction temperature is 600-750° C., the residence time is 1-20 seconds, and the reaction pressure is 0.10-1.0 MPa (absolute pressure) , The weight ratio of the catalyst to the raw material is 4-100. 6. The method according to claim 5, characterized in that the operating conditions of the third reactor are: the reaction temperature is 650-700°C, the residence time is 1-10 seconds, and the weight ratio of the catalyst to the raw material is 5-80 . 7, according to the method for claim 1, it is characterized in that described heavy raw material is selected from vacuum wax oil, normal pressure wax oil, coker wax oil, deasphalted oil, vacuum residual oil, normal pressure residual oil, re-refined oil , oil slurry, diesel oil or a mixture of several. 8. The method according to claim 1, characterized in that said light feedstock is selected from gasoline and/or C4-C8 hydrocarbon mixtures. 9. The method according to claim 8, characterized in that said light feedstock is selected from C4-C8 hydrocarbon mixtures. 10. The method according to claim 1, characterized in that said catalyst contains zeolite, inorganic oxide and optional clay, based on the total weight of the catalyst, contains 10-50w% of zeolite, 5-90w% of inorganic oxide, clay 0~70w%. 11. The method according to claim 10, characterized in that said catalyst active component is selected from Y-type or HY-type zeolite containing or not containing rare earth, ultra-stable Y-type zeolite containing or not containing rare earth, zeolite with MFI structure Catalyst of one or more of zeolite or high silica zeolite with five-membered ring structure prepared by other methods. 12. The method according to claim 1, characterized in that said first reactor is a riser, an ascending conveying line, a descending conveying line or a fluidized bed reactor; said second reactor is a fluidized bed Reactor; the third reactor is a riser reactor. 13. The method according to claim 12, characterized in that said first reactor is a riser reactor or a fluidized bed reactor. 14. The method of claim 12 wherein said first reactor is a riser reactor.

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