CN115725326A

CN115725326A - Catalytic conversion method and device for producing ethylene, propylene and light aromatic hydrocarbon

Info

Publication number: CN115725326A
Application number: CN202111016666.2A
Authority: CN
Inventors: 马文明; 成晓洁; 朱根权
Original assignee: Sinopec Research Institute of Petroleum Processing; China Petroleum and Chemical Corp
Current assignee: Sinopec Research Institute of Petroleum Processing; China Petroleum and Chemical Corp
Priority date: 2021-08-31
Filing date: 2021-08-31
Publication date: 2023-03-03

Abstract

The invention discloses a catalytic conversion method for producing ethylene, propylene and light aromatic hydrocarbon, which comprises the following steps: feeding the preheated heavy raw material into a first reactor to contact and react with a heavy oil catalyst to obtain first reaction oil gas and a first reactant; feeding the light raw material into a second reactor to contact and react with a light oil catalyst to obtain second reaction oil gas and a second spent catalyst; carrying out heat exchange on the second reaction oil gas and the heavy raw material to obtain a preheated heavy raw material; introducing a first agent to be regenerated into a first regenerator for regeneration to obtain a first regenerating agent and first regenerated flue gas; introducing part or all of the first regeneration flue gas into a second regenerator, and returning the first regenerant to the first reactor; and introducing the second regenerant into a second regenerator for regeneration to obtain a second regenerant and second regenerated flue gas, and returning the second regenerant to the second reactor. The method can realize the zone control of the reaction regeneration of the heavy raw material and the light raw material, and improve the yield of ethylene, propylene and light aromatic hydrocarbon.

Description

Catalytic conversion method and device for producing ethylene, propylene and light aromatic hydrocarbon

Technical Field

The application relates to the field of petroleum processing, in particular to a catalytic conversion method and a catalytic conversion device for producing ethylene, propylene and light aromatic hydrocarbons.

Background

In recent years, with the continuous development of the chemical industry, the demand of light olefins such as ethylene and propylene and light aromatic hydrocarbons (benzene, toluene and xylene, abbreviated as BTX) is on the trend of rapid increase. The existing methods for producing low-carbon olefins mainly comprise a steam cracking technology, a catalytic cracking technology, a technology for preparing propylene by propane dehydrogenation, a technology for preparing methanol from coal and then preparing olefins, and the like. BTX is mainly derived from light hydrocarbon reforming processes and steam thermal cracking processes. With the adoption of new light raw materials in the steam cracking process, the distribution of products will change, for example, ethane is adopted as the steam cracking raw material, the proportion of ethylene in the products is obviously improved compared with naphtha, and the yield of low-carbon olefin and light aromatic hydrocarbon is reduced. The catalytic cracking process can treat heavy raw oil and can generate more low-carbon olefins and light aromatics, and is an effective supplementary measure for the steam cracking technology. However, the conventional catalytic cracking process has low yields of light olefins and light aromatics, wherein the yield of light olefins is not more than 15% of the raw oil, and thus it is difficult to meet the market demand, and therefore it is necessary to develop a catalytic cracking technology capable of processing heavy raw oil and increasing the yields of light olefins and light aromatics.

Chinese patent document CN1031834A discloses a catalytic conversion method for producing low-carbon olefins. The method uses petroleum fractions and residual oil with different boiling rangesOr crude oil is taken as a raw material, a mixture containing Y zeolite and pentasil zeolite is taken as a catalyst, a fluidized bed or a moving bed is taken as a reactor, and the reaction conditions are as follows: the temperature is 500-650 ℃, the pressure is 0.15-0.30MPa, and the weight hourly space velocity is 0.2-20 hours ^-1 The catalyst after reaction is returned to the reactor for recycling after being burnt and regenerated, and the catalyst has the catalyst oil ratio of 2-12. Compared with the conventional catalytic cracking and steam cracking, the method can obtain more propylene and butylene.

Chinese patent document CN1566267A discloses a catalytic thermal cracking method for producing petroleum hydrocarbon from ethylene and propylene. The method comprises the steps of feeding preheated petroleum hydrocarbon raw materials into a riser reactor, contacting the preheated petroleum hydrocarbon raw materials with a hot catalyst containing pentasil zeolite, and reacting under the catalytic thermal cracking condition. Wherein the pentasil zeolite contains phosphorus and transition metal. The process can increase the yield of light olefins, particularly ethylene and propylene.

Chinese patent document CN104560149A discloses a catalytic conversion method for producing butene. The method is provided with 4 reactors, except for adopting a reactor configuration of double risers and a fluidized bed, the outside of a settler is also provided with a fluidized bed reactor for cracking gasoline fraction, reaction products enter a riser reactor to continuously carry out cracking reaction, and a catalyst after reaction returns to the reactor for recycling after being burnt and regenerated. The method takes the mixture containing the Y zeolite and the beta zeolite as the catalyst, and can obtain higher yields of the propylene and the butylene.

Chinese patent document CN102206509A discloses a hydrocarbon catalytic conversion method for producing propylene and light aromatic hydrocarbons. The method adopts a combined reactor form of a double-riser and a fluidized bed reactor, wherein heavy hydrocarbons and a cracking catalyst containing modified beta zeolite contact and react in a first reactor, C4 hydrocarbons and/or light gasoline fractions and the cracking catalyst containing modified beta zeolite contact and react in a second reactor and then are introduced into a third reactor for continuous reaction, and the third reactor is a fluidized bed reactor and creates conditions for secondary cracking reaction of the gasoline fractions, thereby improving the yield of propylene and light aromatic hydrocarbons.

Chinese patent document CN102690680A discloses a catalytic conversion method and apparatus for producing propylene. The heavy raw material and catalyst using Y-type zeolite as main active component are contacted and reacted in the first riser reactor, and the light hydrocarbon and catalyst using selective zeolite whose pore size is less than 0.7nm as main active component are contacted and reacted in the second riser reactor. The device adopts a combined reactor configuration formed by a double-riser reactor and a fluidized bed, and a stripper is divided into two independent stripping zones by a partition plate. The method can realize the reaction regeneration of two different catalysts in one reaction device, and has simple structure.

The technology improves the yield of the low-carbon olefin and the light aromatic hydrocarbon by adjusting the formula of the catalyst, adopting a combined reactor form combining a lifting pipe and a fluidized bed and adopting different catalysts aiming at different hydrocarbon raw materials, but the independent zone control of different reaction regeneration systems is difficult to realize, the yield of the ethylene, the propylene and the light aromatic hydrocarbon needs to be further improved, and the heat of the device is not coupled and utilized to reduce the energy consumption.

Disclosure of Invention

The purpose of the present disclosure is to provide a method for effectively adjusting the reaction environment of heavy raw materials and light raw materials, and separate zone control of different reaction regeneration systems, so as to realize the coupling utilization of heat and improve the yields of ethylene, propylene and light aromatics.

To achieve the above object, a first aspect of the present disclosure provides a catalytic conversion process for producing ethylene, propylene and light aromatic hydrocarbons, the process comprising:

s1, feeding a preheated heavy raw material into a first reactor, and contacting the heavy raw material with a heavy oil catalyst containing a Y-type molecular sieve and a beta-type molecular sieve to perform a first catalytic cracking reaction to obtain first reaction oil gas and a first catalyst to be generated;

s2, feeding the light raw material into a second reactor, contacting with a light oil catalyst containing an MFI structure molecular sieve, and performing a second catalytic cracking reaction to obtain a second reaction oil gas and a second spent catalyst; introducing the second reaction oil gas into a raw material preheater, and exchanging heat with the heavy raw material to obtain the preheated heavy raw material;

s3, introducing the first catalyst to be regenerated into a first regenerator after steam stripping to perform a first regeneration reaction to obtain a first regenerated catalyst and first regenerated flue gas; introducing part or all of the first regeneration flue gas into a second regenerator, returning first regenerated catalyst to the first reactor; and introducing the second spent catalyst into the second regenerator after steam stripping for a second regeneration reaction to obtain a second regenerated catalyst and second regenerated flue gas, and returning the second regenerated catalyst to the second reactor.

Optionally, the heavy raw material is selected from one or more than one of vacuum wax oil, atmospheric residue, vacuum residue, coking wax oil, deasphalted oil, furfural refined raffinate oil, coal liquefied oil, oil sand oil, shale oil, distillate oil obtained through F-T synthesis and animal and vegetable oil; the light feedstock comprises at least one of C4 hydrocarbons and a light gasoline fraction; the C4 hydrocarbon is a mixture of C4 alkane and C4 alkene; the C4 hydrocarbon has a content of C4 olefins of greater than 50 wt%, preferably greater than 60 wt%; the light gasoline fraction is light gasoline fraction rich in olefin, and the olefin content in the light gasoline fraction is 20-95 wt%, preferably 50-95 wt%; the weight ratio of the light feedstock to the heavy feedstock is from 0.02 to 0.6, preferably from 0.04 to 0.3; the weight ratio of said C4 hydrocarbons to said heavy feedstock is from 0 to 0.30, preferably from 0.02 to 0.15; the weight ratio of the light gasoline fraction to the heavy feedstock is from 0 to 0.30, preferably from 0.02 to 0.15.

Optionally, the first reactor comprises a first riser reactor and/or a first fluidized bed reactor, preferably a combined first riser reactor and first fluidized bed reactor; in the combined reactor, the first fluidized bed reactor is positioned above the first riser reactor and is communicated with the first riser reactor.

Optionally, the reaction temperature of the first riser reactor is from 500 to 580 ℃, preferably from 520 to 560 ℃; the agent-oil ratio is 2-25, preferably 5-20; the reaction time is 0.5 to 10 seconds, preferably 1 to 5 seconds; the reaction temperature of the first fluidized bed reactor is 480-560 ℃, and preferably 500-540 ℃; the weight hourly space velocity is 1-20 hours ^-1 Preferably 2 to 8 hours ^-1 (ii) a The density of the catalyst is 50-350kg/m ³ Preferably 100 to 250kg/m ³ (ii) a The bed height of the first fluidized bed reactor is 1/2-4/5 of the height of the first fluidized bed reactor, and preferably 1/2-3/4 of the height of the first fluidized bed reactor; the pressure in the first fluidized bed reactor is 0.1-0.4MPa, preferably 0.15-0.3MPa.

Optionally, the second reactor is a combined reactor of a second riser reactor and a second fluidized bed reactor, and the second fluidized bed reactor is located above the second riser reactor and is communicated with the second riser reactor.

Optionally, the reaction temperature of the second riser reactor is 630-690 ℃, preferably 650-670 ℃; the agent-oil ratio is 20-80, preferably 40-60; the reaction time is 0.5 to 15 seconds, preferably 1 to 10 seconds; the reaction temperature of the second fluidized bed reactor is 590-650 ℃, preferably 610-630 ℃; the weight hourly space velocity is 1-20 h ^-1 Preferably 2 to 8 hours ^-1 (ii) a The density of the catalyst in the second fluidized bed reactor is 50-350kg/m ³ Preferably 100 to 250kg/m ³ (ii) a The bed height of the second fluidized bed reactor is 1/2-4/5 of the height of the second fluidized bed reactor, and preferably 1/2-3/4 of the height of the second fluidized bed reactor; the pressure in the second fluidized bed reactor is 0.1-0.4MPa, preferably 0.15-0.3MPa.

Optionally, the temperature of the preheated heavy feedstock is in the range of from 100 to 350 ℃, preferably from 150 to 300 ℃.

Optionally, the first regeneration reaction comprises complete regeneration or incomplete regeneration, preferably incomplete regeneration; the regeneration temperature in the first regenerator is 620-720 ℃, preferably 650-680 ℃; the second regeneration reaction comprises complete regeneration or incomplete regeneration, preferably complete regeneration; the regeneration temperature in the second regenerator is 650-750 ℃, preferably 680-730 ℃.

Optionally, the heavy oil catalyst comprises the Y-type molecular sieve, the beta-type molecular sieve, a first clay, and a first binder; the content of the Y-type molecular sieve is 20-60 wt%, preferably 30-50 wt%, based on the total weight of the heavy oil catalyst; the content of the beta-type molecular sieve is 1-40 wt%, preferably 1-20 wt%; the content of the first clay is 10 to 70 wt%, preferably 15 to 45 wt%; the content of the first binder is 10 to 40 wt%, preferably 20 to 35 wt%; the Y-type molecular sieve is selected from at least one of HY type molecular sieve, USY type molecular sieve, REUSY type molecular sieve, REY type molecular sieve, REHY type molecular sieve, DASY type molecular sieve and REDASY type molecular sieve; the beta-type molecular sieve is a beta-type molecular sieve modified by phosphorus and a transition metal M, wherein M is selected from at least one of Fe, co, ni, cu, mn, zn and Sn; the first clay is at least one selected from kaolin, montmorillonite and bentonite; the first binder is at least one selected from silica sol, aluminum sol and pseudo-boehmite, and preferably, the first binder is double-aluminum binder of the aluminum sol and the pseudo-boehmite.

Optionally, the light oil catalyst comprises the MFI structure molecular sieve, a second clay, and a second binder; the content of the MFI structure molecular sieve is 20-60 wt%, preferably 30-50 wt%, based on the total weight of the light oil catalyst; the content of the second clay is 10 to 70 wt%, preferably 15 to 45 wt%; the content of the second binder is 10-40 wt%, preferably 20-35 wt%; the MFI structure molecular sieve is selected from at least one of ZRP molecular sieves, ZRP molecular sieves containing phosphorus, ZRP molecular sieves containing rare earth, ZRP molecular sieves containing phosphorus and alkaline earth metal and ZRP molecular sieves containing phosphorus and transition metal, and the ZRP zeolite containing phosphorus and rare earth is preferred; the second clay is at least one of kaolin, montmorillonite and bentonite; the second binder is at least one selected from silica sol, aluminum sol and pseudo-boehmite, and preferably, the second binder is double-aluminum binder of the aluminum sol and the pseudo-boehmite.

Optionally, the method further comprises: and separating the second reaction oil gas after heat exchange to obtain C4 hydrocarbon containing olefin and a light gasoline fraction, and returning part or all of the C4 hydrocarbon containing olefin and the light gasoline fraction to the second reactor.

A second aspect of the present disclosure provides an apparatus for a catalytic conversion process for producing ethylene, propylene and light aromatic hydrocarbons, the catalytic conversion apparatus comprising a first reactor, a second reactor, a first regenerator, a second regenerator and a feedstock preheater, the first reactor and the first regenerator being connected by a first catalyst delivery line; the second reactor is connected with the second regenerator through a second catalyst conveying pipeline; the first reactor comprises a preheated heavy raw material inlet, the second reactor comprises a light raw material reaction oil gas outlet, and the raw material preheater comprises a reaction oil gas inlet before heat exchange, a reaction oil gas outlet after heat exchange, a heavy raw material inlet before preheating and a heavy raw material outlet after preheating; the light raw material reaction oil gas outlet is connected with the pre-heat exchange reaction oil gas inlet, and the preheated heavy raw material outlet is connected with the preheated heavy raw material inlet of the first reactor; the first regenerator and the second regenerator are arranged in parallel or coaxially, preferably in parallel; the first regenerator includes a first regenerated flue gas outlet, the second regenerator includes a first regenerated flue gas inlet, and the first regenerated flue gas outlet is connected to the first regenerated flue gas inlet.

According to the technical scheme, in the catalytic conversion method and the catalytic conversion device, the heavy raw material and the light raw material react in different reaction regeneration systems respectively, so that the independent zone control of the reaction regeneration of different heavy raw materials and light raw materials can be realized, the primary cracking of the heavy raw material and the secondary cracking of the light raw material are enhanced by adopting different catalysts and reaction conditions, and the yield of ethylene, propylene and light aromatic hydrocarbon is improved; meanwhile, the first regeneration flue gas generated by the first regenerator provides heat for the regeneration of the light oil catalyst so as to reach the temperature required by the regeneration of the light oil catalyst and provide heat for the catalytic conversion reaction of the light raw material. The second reaction oil gas generated by catalytic cracking of the light raw material provides heat for preheating of the heavy raw material, and the energy consumption of the device can be greatly reduced through coupling utilization of the heat.

Additional features and advantages of the present disclosure will be set forth in the detailed description which follows.

Drawings

The accompanying drawings, which are included to provide a further understanding of the disclosure and are incorporated in and constitute a part of this specification, illustrate embodiments of the disclosure and together with the description serve to explain the disclosure without limiting the disclosure. In the drawings:

FIG. 1 is a schematic flow diagram of an apparatus suitable for use in the catalytic conversion process of the present disclosure;

FIG. 2 is a particular embodiment of an apparatus suitable for use in the catalytic conversion process of the present disclosure;

FIG. 3 is another specific embodiment of an apparatus suitable for use in the catalytic conversion process of the present disclosure;

FIG. 4 is a preferred embodiment of an apparatus suitable for use in the catalytic conversion process of the present disclosure;

FIG. 5 is yet another specific embodiment of an apparatus suitable for use in the catalytic conversion process of the present disclosure.

Description of the reference numerals

1 first reactor 2 first regenerator

3 second regenerator 4 second reactor

5 raw material preheater

1-1 first riser reactor 1-2 first stripper

1-3 first fluidized bed reactor 1-4 first settler

101 heavy feed 102 cracked heavy oil

103 first pre-lift gas 104 first lift gas

105 first stripping baffle 106 heavy oil spent catalyst conveying pipeline

107 first cyclone separator 108 second cyclone separator

109 first plenum 110 first separation system line

2-1 first regenerator 201 first regenerator main wind inlet line

202 heavy oil regenerated catalyst conveying pipeline 203 and third cyclone separator

204 fourth cyclone 205 second plenum

206 first regenerated flue gas

3-1 second regenerator 301 second regenerator main air inlet line

302 light oil regenerated catalyst transfer line 303 fifth cyclone

304 sixth cyclone separator 305 fourth plenum

306 second regeneration flue gas

4-1 second riser reactor 4-2 second stripper

4-3 second fluidized bed reactor 4-4 second settler

401 first part of light raw material 402 second part of light raw material

403 second pre-lift gas 404 second lift gas

405 second stripping baffle 406 spent light oil catalyst transfer line

407 seventh cyclone 408 eighth cyclone

409 third plenum 410 second separation system line

501 reaction oil gas before heat exchange and 502 reaction oil gas after heat exchange

503 heavy feedstock before preheating 504 heavy feedstock after preheating

Detailed Description

The following describes in detail specific embodiments of the present disclosure. It should be understood that the detailed description and specific examples, while indicating the present disclosure, are given by way of illustration and explanation only, not limitation.

A first aspect of the present disclosure provides a catalytic conversion process for producing ethylene, propylene and light aromatics, the process comprising:

s1, feeding the preheated heavy raw material into a first reactor, and contacting the heavy raw material with a heavy oil catalyst containing a Y-type molecular sieve and a beta-type molecular sieve to perform a first catalytic cracking reaction to obtain first reaction oil gas and a first catalyst to be generated;

According to the catalytic conversion method, the heavy raw material and the light raw material are respectively reacted in different reaction regeneration systems, and different catalysts and reaction conditions are adopted to strengthen the primary cracking reaction of the heavy raw material and the secondary cracking reaction of the light raw material, so that the yield of ethylene, propylene and light aromatic hydrocarbon is improved; meanwhile, the regeneration flue gas generated by the first regenerator provides heat for the regeneration of the light oil catalyst so as to reach the temperature required by the regeneration of the light oil catalyst and provide heat for the catalytic conversion reaction of the light raw material. The reaction oil gas provides heat for preheating the heavy raw material, and the energy consumption of the device can be greatly reduced.

According to the present disclosure, the heavy feedstock undergoes primarily cracking reactions in the first reactor, converting from macromolecular reactants to small molecular products. The heavy raw material can be one or a mixture of more than one of vacuum wax oil, atmospheric residue oil, vacuum residue oil, coking wax oil, deasphalted oil, furfural refined raffinate oil, coal liquefied oil, oil sand oil, shale oil, distillate oil obtained by F-T synthesis or animal and vegetable oil.

The light feedstock of the present disclosure can be introduced at the same location of the second riser reactor or at different locations, preferably at different locations. The weight ratio of the light feedstock to the heavy feedstock in this disclosure is from 0.02 to 0.6, preferably from 0.04 to 0.3. The light feedstock may include at least one of C4 hydrocarbons and a light gasoline fraction; the C4 hydrocarbon may be a mixture of C4 alkanes and C4 alkenes; the C4 olefin content of the C4 hydrocarbon may be greater than 50 wt%, preferably greater than 60 wt%. The C4 hydrocarbons in this disclosure include C4 hydrocarbons produced by the catalytic conversion unit of this disclosure during a reaction or by other units, preferably C4 hydrocarbons produced by the catalytic conversion unit of this disclosure. The weight ratio of C4 hydrocarbons injected into the second reactor to heavy feedstock injected into the first reactor in this disclosure is from 0 to 0.30, preferably from 0.02 to 0.15. The light gasoline fraction is an olefin-rich light gasoline fraction having an olefin content of 20 to 95% by weight, preferably 50 to 95% by weight. The light gasoline fraction includes light gasoline fractions produced by the catalytic conversion unit of the present disclosure or light gasoline fractions produced by other units; the light gasoline fraction produced by other devices can be one or more than one of light gasoline fraction, coking light gasoline fraction, visbreaking light gasoline fraction and light gasoline fraction produced by other oil refining or chemical engineering processes, and is preferably the light gasoline fraction produced by the device. The weight ratio of the light gasoline fraction injected into the second riser reactor to the heavy feedstock injected into the first reactor in the present disclosure is from 0 to 0.30, preferably from 0.02 to 0.15.

According to the present disclosure, the first reactor may comprise a first riser reactor and/or a first fluidized bed reactor, preferably a combined first riser reactor and first fluidized bed reactor. When the first reactor is a combined reactor of the first riser reactor and the first fluidized bed reactor, different reaction environments can be created for the primary cracking reaction of the heavy raw material and the secondary cracking reaction of the intermediate product, the conversion of the heavy raw material is more facilitated, and the low-carbon olefin and the aromatic hydrocarbon can be produced in more. Wherein, in the combined reactor, the first fluidized bed reactor may be positioned above and in communication with the first riser reactor.

According to the present disclosure, the reaction temperature of the first riser reactor may beIs 500-580 deg.C, preferably 520-560 deg.C; the agent-oil ratio can be 2-25, preferably 5-20; the reaction time may be from 0.5 to 10 seconds, preferably from 1 to 5 seconds; the reaction temperature of the first fluidized bed reactor can be 480-560 ℃, and is preferably 500-540 ℃; the weight hourly space velocity can be 1-20 h ^-1 Preferably 2 to 8 hours ^-1 (ii) a The catalyst density can be 50-350kg/m ³ Preferably 100 to 250kg/m ³ (ii) a The height of the first fluidized bed reactor can be 1/2-4/5 of the height of the first fluidized bed reactor, and is preferably 1/2-3/4 of the height of the first fluidized bed reactor; the pressure in the first fluidized bed reactor may be in the range of 0.1 to 0.4MPa, preferably 0.15 to 0.3MPa.

According to the present disclosure, the second reactor may be a combined reactor of a second riser reactor and a second fluidized bed reactor, and as a preferred embodiment of the present disclosure, the second fluidized bed reactor is located above and communicates with the second riser reactor.

According to the present disclosure, the reaction temperature of the second riser reactor may be 630 to 690 ℃, preferably 650 to 670 ℃; the agent-oil ratio can be 20-80, preferably 40-60; the reaction time may be from 0.5 to 15 seconds, preferably from 1 to 10 seconds; the reaction temperature of the second fluidized bed reactor may be 590-650 ℃, preferably 610-630 ℃; the weight hourly space velocity can be 1-20 hours ^-1 Preferably 2 to 8 hours ^-1 (ii) a The catalyst density in the second fluidized bed reactor can be 50-350kg/m ³ Preferably 100 to 250kg/m ³ (ii) a The bed height of the second fluidized bed reactor can be 1/2-4/5 of the height of the second fluidized bed reactor, and is preferably 1/2-3/4 of the height of the second fluidized bed reactor; the pressure in the second fluidized bed reactor may be in the range of 0.1 to 0.4MPa, preferably 0.15 to 0.3MPa.

According to the present disclosure, the temperature of the preheated heavy feedstock may be in the range of from 100 to 350 ℃, preferably from 150 to 300 ℃.

According to the present disclosure, the first regeneration reaction may comprise complete regeneration or incomplete regeneration, preferably incomplete regeneration; the regeneration temperature in the first regenerator may be from 620 to 720 ℃, preferably from 650 to 680 ℃; the second regeneration reaction may comprise complete regeneration or incomplete regeneration, preferably complete regeneration; the regeneration temperature in the second regenerator may be in the range 650 to 750 c, preferably 680 to 730 c. In the method disclosed by the disclosure, part or all of the regeneration flue gas generated in the first regenerator is introduced into the second regenerator to provide heat for the regeneration of the light oil catalyst, so that the coupled utilization of heat is realized, and part of the regeneration flue gas is preferably introduced into the second regenerator.

According to the present disclosure, the heavy oil catalyst may include the Y-type molecular sieve, the beta-type molecular sieve, a first clay, and a first binder; the content of the Y-type molecular sieve may be 20 to 60 wt%, preferably 30 to 50 wt%, based on the total weight of the heavy oil catalyst; the content of the beta-type molecular sieve may be 1 to 40 wt%, preferably 1 to 20 wt%; the first clay may be present in an amount of 10 to 70 wt%, preferably 15 to 45 wt%; the content of the first binder may be 10 to 40 wt%, preferably 20 to 35 wt%; the Y-type molecular sieve is selected from at least one of HY type molecular sieve, USY type molecular sieve, REUSY type molecular sieve, REY type molecular sieve, REHY type molecular sieve, DASY type molecular sieve and REDASY type molecular sieve; the beta type molecular sieve can be a beta type molecular sieve modified by phosphorus and a transition metal M, wherein M can be at least one selected from Fe, co, ni, cu, mn, zn and Sn; the beta-type molecular sieve modified by phosphorus and transition metal M can be prepared by various methods, for example, phosphorus and transition metal M can be introduced during the synthesis of the beta-type molecular sieve, or phosphorus and transition metal M can be introduced after the synthesis of the beta-type molecular sieve by steps of ammonium exchange, phosphorus modification, transition metal M modification, calcination treatment and the like. The first clay can be at least one selected from kaolin, montmorillonite and bentonite; the first binder may be at least one selected from the group consisting of silica sol, alumina sol and pseudo-boehmite, and preferably, the first binder may be a double alumina binder of alumina sol and pseudo-boehmite.

According to the present disclosure, the light oil catalyst may include the MFI structure molecular sieve, a second clay, and a second binder; the MFI structure molecular sieve may be present in an amount of 20 to 60 wt.%, preferably 30 to 50 wt.%, based on the total weight of the light oil catalyst; the content of the second clay may be 10 to 70% by weight, preferably 15 to 45% by weight; the content of the second binder may be 10 to 40% by weight, preferably 20 to 35% by weight; the MFI structure molecular sieve can be selected from one or more of ZRP molecular sieve, ZRP molecular sieve containing phosphorus, ZRP molecular sieve containing rare earth, ZRP molecular sieve containing phosphorus and alkaline earth metal and ZRP molecular sieve containing phosphorus and transition metal, preferably ZRP zeolite containing phosphorus and rare earth; the second clay can be at least one selected from kaolin, montmorillonite and bentonite; the second binder may be at least one selected from the group consisting of silica sol, alumina sol and pseudo-boehmite, and preferably, the second binder is a double alumina binder of alumina sol and pseudo-boehmite.

According to the present disclosure, the method may further include: the second reaction oil gas passing through the heavy raw material preheater is further separated into products such as dry gas, cracked gas, gasoline, light oil, oil slurry and the like in a product separation system, wherein the cracked gas can obtain a polymer grade propylene product and C4 hydrocarbon rich in olefin after subsequent product separation and refining; gasoline is firstly cut into light and medium gasoline distillation sections, and part or all of C4 hydrocarbon rich in olefin is directly returned to the second riser reactor for reaction; and returning part or all of the light gasoline to the second riser reactor for reaction.

The catalytic conversion apparatus of the present disclosure may specifically include: the system comprises a first riser reactor 1-1, a first fluidized bed reactor 1-3, a first regenerator 2-1, a second regenerator 3-1, a second riser reactor 4-1, a second fluidized bed reactor 4-3 and a raw material preheater 5-1. The first riser reactor 1-1 is connected with the first regenerator 2-1 through a heavy oil regenerated catalyst conveying pipeline 202, and the first stripper 1-2 is connected with the first regenerator 2-1 through a heavy oil spent catalyst conveying pipeline 106, and is used for catalytic conversion reaction of heavy raw materials and regeneration of heavy oil catalysts; the second riser reactor 4-1 is connected to the second regenerator 3-1 through a light oil regenerated catalyst transfer line 302, and the second stripper 4-2 is connected to the second regenerator 3-1 through a light oil spent catalyst transfer line 406, for catalytic conversion of the light feedstock and regeneration of the light oil catalyst.

Lift gas is introduced to the first riser reactor 1-1 via a first pre-lift gas line 103. The lift gas used is well known to those skilled in the art and may be selected from one or more of steam, nitrogen, dry gas, preferably steam. First stripper 1-2 wraps at least a portion of first riser reactor 1-1 and is in communication with first regenerator 2-1 via heavy oil spent catalyst transfer line 106. The first fluidized bed reactor 1-3 is located above the first riser reactor 1-1 and connected to the first riser reactor 1-1. The first settler 1-4 is communicated with the first fluidized bed reactor 1-3, reaction oil gas in the first fluidized bed reactor 1-3 can directly enter the first settler 1-4, is separated by the first cyclone separator 107 and the second cyclone separator 108, then enters the first gas collection chamber 109, and is led out of the reactor through the separation system pipeline 110. The inlets of the first cyclone separator 107 and the second cyclone separator 108 are positioned at the upper part of the first precipitator 1-4, the catalyst outlets of the first cyclone separator 107 and the second cyclone separator 108 enable the catalyst therein to enter the first cyclone separator 1-2, and the oil gas outlets of the first cyclone separator 107 and the second cyclone separator 108 are connected and then communicated with an oil gas separation system through a first gas collection chamber 109 and a first separation system pipeline 110. Spent catalyst from the first stripper 1-2 enters the first regenerator 2-1 through the heavy oil spent catalyst transfer line 106 for coke burning regeneration, thereby converting the spent catalyst into a regenerated catalyst. The regenerated catalyst in the first regenerator 2-1 is returned to the pre-lifting section of the first riser reactor 1-1 for recycling through the heavy oil regenerated catalyst conveying line 202, wherein the conveying speed of the catalyst can be adjusted through valves on the heavy oil regenerated catalyst conveying line 202 and the heavy oil spent catalyst conveying line 106. The stripper 1-2 is provided with a first stripping baffle 105 and a stripping gas distribution ring for reducing the falling speed of the spent catalyst and making the stripping gas distribution more uniform, so that the residual reaction oil gas on the spent catalyst is fully removed from the spent catalyst. The bottom of the first regenerator 2-1 is provided with a first regenerator main air inlet pipeline 201, which can introduce regeneration gas into the first regenerator 2-1, and the flue gas generated after regeneration enters a second gas collection chamber 205 through a third cyclone 203 and a fourth cyclone 204, and then is introduced into a second regenerator 3-1 to provide heat for the regeneration of light oil catalyst.

A lift gas is introduced to the second riser reactor 4-1 via second pre-lift gas line 403. The lift gas used is well known to those skilled in the art and may be selected from one or more of steam, nitrogen, dry gas, preferably steam. The second stripper 4-2 wraps at least a portion of the second riser reactor 4-1 and is in communication with the second regenerator 3-1 via a spent light oil catalyst transfer line 406. The second fluidized bed reactor 4-3 is located above the second riser reactor 4-1 and is connected to the second riser reactor 4-1. The second settler 4-4 is communicated with the second fluidized bed reactor 4-3, the reaction oil gas in the second fluidized bed reactor 4-3 can directly enter the second settler 4-4, is separated by the seventh cyclone 407 and the eighth cyclone 408, then enters the third gas collection chamber 409, and is led out of the reactor through the second separation system pipeline 410. Inlets of the seventh cyclone 407 and the eighth cyclone 408 are positioned at the upper part of the second precipitator 4-4, catalyst outlets of the seventh cyclone 407 and the eighth cyclone 408 enable catalyst therein to enter the second precipitator 4-2, and oil and gas outlets of the seventh cyclone 407 and the eighth cyclone 408 are connected and then communicated with an oil and gas separation system through a third gas collection chamber 409 and a second separation system pipeline 410. The spent catalyst from the second stripper 4-2 enters the second regenerator 3-1 through a light oil spent catalyst transfer line 406 for scorching regeneration, thereby converting the spent catalyst into a regenerated catalyst. The regenerated catalyst in the second regenerator 3-1 is returned to the pre-lift section of the second riser reactor 4-1 for recycling through the light oil regenerated catalyst line 302, wherein the catalyst transfer rate can be adjusted by valves on the light oil regenerated catalyst transfer line 302 and the light oil spent catalyst transfer line 406. The second stripper 4-2 is provided with a second stripping baffle 405 and a stripping gas distribution ring for reducing the falling speed of the spent catalyst and making the stripping gas distribution more uniform, so that the residual reaction oil gas on the spent catalyst is sufficiently removed from the spent catalyst. The bottom of the second regenerator 3-1 is provided with a second regenerator main air inlet pipeline 301, which can introduce regeneration gas into the second regenerator 3-1, and the flue gas generated after regeneration enters a fourth gas collection chamber 305 through a fifth cyclone 303 and a sixth cyclone 304, and is led out of the system after treatment.

As a preferred embodiment of the present disclosure, the method of the present disclosure may further include rapidly separating the oil gas from the reacted soot catalyst using a separation device. Thus, the dry gas yield can be reduced, the propylene is inhibited from being converted after being produced, and the product separation system used in the present disclosure can be various separation systems in the prior art.

According to the present disclosure, as shown in fig. 1, a first reactor 1 is connected to a first regenerator 2 through a heavy oil spent catalyst transfer line 106 and a heavy oil regenerated catalyst transfer line 202 for catalytic conversion reaction of heavy raw materials and regeneration of heavy oil catalyst; the second reactor 4 is connected to the second regenerator 3 through a light oil regenerated catalyst transfer line 302 and a light oil spent catalyst transfer line 406, and is used for catalytic conversion reaction of the light raw material and regeneration of the light oil catalyst.

As a specific embodiment of the present disclosure, as shown in fig. 2, the first reactor 1 is a first riser reactor 1-1, the first regenerator 2-1 and the second regenerator 3-1 are arranged in parallel, the second regenerator 3-1 is connected to the first regenerator 2-1 through a regeneration flue gas conveying line, the regeneration reaction performed in the first regenerator 2-1 is incomplete regeneration, and a part of the generated regeneration flue gas enters the second regenerator 3-1.

As another specific embodiment of the present disclosure, as shown in FIG. 3, a first regenerator 2-1 and a second regenerator 3-1 are coaxially arranged, the second regenerator 3-1 is located at the upper part of the first regenerator 2-1 and is connected with the first regenerator 2-1 through a regeneration flue gas conveying line, the regeneration reaction in the first regenerator 2-1 is incomplete, and part of the generated regeneration flue gas enters the second regenerator 3-1.

As a preferred embodiment of the present disclosure, as shown in fig. 4, the first reactor 1 is a combined reactor of a first riser reactor 1-1 and a first fluidized bed reactor 1-3, the first regenerator 2-1 and a second regenerator 3-1 are arranged in parallel, the second regenerator 3-1 is connected to the first regenerator 2-1 through a regeneration flue gas conveying line, the regeneration reaction performed in the first regenerator 2-1 is incomplete, and a part of the generated regeneration flue gas enters the second regenerator 3-1. This embodiment can effectively improve the catalytic cracking efficiency of heavy raw materials by configuring the first reactor 1 as a combined reactor of the first riser reactor 1-1 and the first fluidized bed reactor 1-3, and can effectively realize the coupling utilization of heat by adopting the parallel arrangement of the first regenerator 2-1 and the second regenerator 3-1.

As still another specific embodiment of the present disclosure, as shown in fig. 5, the first reactor 1 is a combined reactor of a first riser reactor 1-1 and a first fluidized bed reactor 1-3, the first regenerator 2-1 and a second regenerator 3-1 are coaxially arranged, the second regenerator 3-1 is located at the upper part of the first regenerator 2-1 and is connected to the first regenerator 2-1 through a regeneration flue gas delivery line, the regeneration reaction performed in the first regenerator 2-1 is incomplete regeneration, and a part of the generated regeneration flue gas enters the second regenerator 3-1.

The present disclosure is further illustrated by the following examples. The raw materials used in the examples are all available from commercial sources. In the embodiment and the comparative example of the disclosure, the gas product is tested by a petrochemical analysis method RIPP 77-90 method, the coke content is determined by a petrochemical analysis method RIPP 107-90 method, the composition of the organic liquid product is determined by an SH/T0558-1993 method, the cut points of the gasoline and the diesel are 221 ℃ and 331 ℃ respectively, and the light aromatic hydrocarbon in the gasoline is determined by a petrochemical analysis method RIPP 82-90.

In the following examples, the conversion of the feedstock oil and the yield of cracked products were calculated according to the following formulas:

the RIPP petrochemical analysis method used in the invention is selected from the editions of "petrochemical analysis method (RIPP test method)", yanggui and the like, and the science publishing company, 1990.

The reagents used below were all chemically pure reagents except those specifically mentioned.

The Y-type molecular sieve is produced by Qilu catalyst factories and has the industrial grades as follows:

DASY, physical parameters: the unit cell constant is 2.443nm, na ₂ The O content was 0.85% by weight;

the beta molecular sieve used was produced by the Qilu catalyst works, siO ₂ /Al ₂ O ₃ ＝25；

The MFI structure molecular sieve is produced by Qilu catalyst factories, and the industrial grades are as follows:

ZRP-1: wherein SiO is ₂ /Al ₂ O ₃ ＝30，Na ₂ O content of 0.17 wt%, rare earth oxide RE ₂ O ₃ Is 1.4 wt%, with lanthanum oxide being 0.84 wt%, cerium oxide being 0.18 wt%, and the other rare earth oxides being 0.38 wt%.

Specific properties of the stock oils used in the examples and comparative examples are shown in Table 1.

The catalysts used in the examples and comparative examples were in-process catalysts, the compositions and properties of which are shown in Table 2.

TABLE 1

TABLE 2

Item	CAT-1	CAT-2	CAT-3
				Molecular sieve type	USY+β	ZRP-1	USY+β+ZRP-1
Elemental composition,% (w)
				Al ₂ O ₃	54.6	50.9	52.8
SiO ₂	40.9	43.7	42.6
				Micro-reverse activity,% (W)	69	63	67
Specific surface area, m ² /g	151	169	158
				Pore volume, ml/g	0.459	0.524	0.496
Bulk ratio, g/ml	0.89	0.91	0.9
				Particle size distribution
0-20μm	0.4	0.3	0.4
				0-40μm	15.8	16.2	15.9
0-80μm	65.2	68.4	68.7
				0-105μm	88.7	86.3	89.2
＞105μm	11.3	13.7	10.8

Example 1

The test was carried out on a medium-sized test apparatus as shown in FIG. 2. The apparatus comprises two riser reactors, a fluidized bed reactor and two regenerators. The inner diameter of the first riser reactor 1-1 is 20mm, the height is 4000mm, the inner diameter of the second riser reactor 4-1 is 16mm, the height is 3000mm, the inner diameter of the second fluidized bed reactor 4-3 is 80mm, and the height is 500mm. The first regenerator 2-1 has an inner diameter of 150mm and a height of 800mm, and the second regenerator 3-1 has an inner diameter of 100mm and a height of 600mm. The first regenerator 2-1 is arranged in parallel with the second regenerator 3-1.

Introducing the preheated raw oil into the bottom of a first riser reactor 1-1, contacting with a catalyst CAT-1, and carrying out a first catalytic cracking reaction to obtain an oil mixture, and separating the oil mixture by a cyclone separator to obtain first reaction oil gas and a first catalyst to be generated. The first catalyst to be regenerated enters a first stripper 1-2 and then enters a first regenerator 2-1 for regeneration to obtain a first regenerated catalyst and first regenerated flue gas, the first regenerated catalyst returns to the first riser reactor 1-1 for recycling, and the first regenerated flue gas is introduced into a second regenerator. The first reaction oil gas is introduced into a fractionation system for separation.

C4 hydrocarbon is introduced into the bottom of a second riser reactor 4-1, contacts with a catalyst CAT-2 and carries out a second catalytic cracking reaction, a generated oil mixture is introduced into a fluidized bed reactor 4-3, and the reacted oil mixture is separated by a cyclone separator to obtain second reaction oil gas and a second spent catalyst. The second spent catalyst enters a second stripper 4-2 and then enters a second regenerator 3-1 for regeneration, the regenerated second regenerated catalyst returns to a second riser reactor 4-1 for recycling, and the second reaction oil gas is introduced into a fractionation system for separation. The mass ratio of C4 hydrocarbons to feed oil in this example was 0.4. The reaction conditions and results are shown in Table 3.

Example 2

The test was carried out on a medium-sized test apparatus as shown in FIG. 3. The apparatus comprises two riser reactors, a fluidized bed reactor and two regenerators. The inner diameter of the first riser reactor 1-1 is 20mm, the height is 4000mm, the inner diameter of the second riser reactor 4-1 is 16mm, the height is 3000mm, the inner diameter of the fluidized bed reactor 4-3 is 80mm, and the height is 500mm. The first regenerator 2-1 has an inner diameter of 150mm and a height of 800mm, and the second regenerator 3-1 has an inner diameter of 100mm and a height of 600mm. The first regenerator 2-1 is arranged coaxially with the second regenerator 3-1, and the second regenerator 3-1 is located at the upper part of the first regenerator 2-1.

Introducing the preheated raw oil into the bottom of a first riser reactor 1-1, contacting with a catalyst CAT-1 and carrying out a first catalytic cracking reaction to obtain an oil agent mixture, and separating the oil agent mixture by a cyclone separator to obtain first reaction oil gas and a first catalyst to be generated. The first catalyst to be regenerated enters a first stripper 1-2 and then enters a first regenerator 2-1 for regeneration to obtain a first regenerated catalyst and first regenerated flue gas, the first regenerated catalyst returns to the first riser reactor 1-1 for recycling, and the first regenerated flue gas is introduced into a second regenerator. The first reaction oil gas is introduced into a fractionation system for separation.

C4 hydrocarbon is introduced into the bottom of a second riser reactor 4-1, contacts with a catalyst CAT-2 and reacts, a generated oil agent mixture is introduced into a fluidized bed reactor 4-3, the reacted oil agent mixture is separated through a cyclone separator to obtain second reaction oil gas and a second catalyst to be regenerated, the second catalyst to be regenerated enters a second stripper 4-2 and then enters a second regenerator 3-1 for regeneration, the regenerated second catalyst returns to the second riser reactor 4-1 for recycling, and the second reaction oil gas is introduced into a fractionation system for separation. Wherein the mass ratio of the C4 hydrocarbon to the raw oil is 0.4. The reaction conditions and results are shown in Table 3.

Example 3

The test was carried out on a medium-sized test apparatus as shown in FIG. 4. The apparatus includes two riser reactors, two fluidized bed reactors, and two regenerators. The first riser reactor 1-1 has an inner diameter of 20mm and a height of 4000mm, and the second riser reactor 4-1 has an inner diameter of 16mm and a height of 3000mm. The first fluidized bed reactor 1-3 has an inner diameter of 90mm and a height of 600mm, and the second fluidized bed reactor 4-3 has an inner diameter of 80mm and a height of 500mm. The first regenerator 2-1 has an inner diameter of 150mm and a height of 800mm, and the second regenerator 3-1 has an inner diameter of 100mm and a height of 600mm. The first regenerator 2-1 is arranged in parallel with the second regenerator 3-1.

Introducing C4 hydrocarbon into the bottom of a second riser reactor 4-1, introducing light gasoline fraction (distillation range is 40-80 ℃, and olefin content is 65 wt%) into the middle of the second riser reactor 4-1, contacting and reacting with a catalyst CAT-2, introducing a generated oil mixture into a fluidized bed reactor 4-3, and separating the reacted oil mixture through a cyclone separator to obtain second reaction oil gas and a second catalyst to be generated; the second spent catalyst enters a second stripper 4-2 and then enters a second regenerator 3-1 for regeneration, the regenerated second regenerated catalyst returns to a second riser reactor 4-1 for recycling, and the second reaction oil gas is introduced into a fractionation system for separation. Wherein the mass ratio of the C4 hydrocarbon, the light gasoline fraction and the raw oil is 0.2. The reaction conditions and results are shown in Table 3.

Example 4

The test was carried out on a medium-sized test apparatus as shown in FIG. 5. The apparatus includes two riser reactors, two fluidized bed reactors and two regenerators. The first riser reactor 1-1 has an inner diameter of 20mm and a height of 4000mm, and the second riser reactor 4-1 has an inner diameter of 16mm and a height of 3000mm. The first fluidized bed reactor 1-3 has an inner diameter of 90mm and a height of 600mm, and the second fluidized bed reactor 4-3 has an inner diameter of 80mm and a height of 500mm. The first regenerator 2-1 has an inner diameter of 150mm and a height of 800mm, and the second regenerator 3-1 has an inner diameter of 100mm and a height of 600mm. The first regenerator 2-1 is arranged coaxially with the second regenerator 3-1, and the second regenerator 3-1 is located at the upper part of the first regenerator 2-1.

Introducing C4 hydrocarbon into the bottom of a second riser reactor 4-1, introducing light gasoline fraction (the distillation range is 40-80 ℃, and the olefin content is 65 wt%) into the middle of the second riser reactor 4-1, contacting and reacting with a catalyst CAT-2, introducing a generated oil agent mixture into a fluidized bed reactor 4-3, and separating the reacted oil agent mixture through a cyclone separator to obtain second reaction oil gas and a second catalyst to be generated; the second spent catalyst enters a second stripper 4-2 and then enters a second regenerator 3-1 for regeneration, the regenerated second regenerated catalyst returns to a second riser reactor 4-1 for recycling, and the second reaction oil gas is introduced into a fractionation system for separation. Wherein the mass ratio of the C4 hydrocarbon, the light gasoline fraction and the raw oil is 0.2. The reaction conditions and results are shown in Table 3.

TABLE 3

Comparative example 1

The tests were carried out on a medium-sized test apparatus. The apparatus comprises a riser reactor and a fluidized bed reactor. The riser reactor had an internal diameter of 16mm and a length of 3200mm, and the fluidized bed reactor had an internal diameter of 64mm and a height of 500mm.

Introducing raw oil into the bottom of a riser reactor, contacting and reacting with a catalyst CAT-3, introducing a generated oil agent mixture into a fluidized bed reactor, separating the reacted oil agent mixture through a fast separation device to obtain reaction oil gas and a regenerated catalyst, introducing the reaction oil gas into a cyclone separator, and introducing the regenerated catalyst into a stripper for stripping.

The regenerated catalyst enters the stripper and then enters the regenerator for regeneration, the regenerated catalyst returns to the riser reactor for recycling, and the oil gas is introduced into the fractionation system for separation. The reaction conditions and results are shown in Table 4.

Comparative example 2

The tests were carried out on a medium-sized test apparatus. The apparatus comprises two riser reactors and a fluidized bed reactor. The internal diameter of the first riser reactor is 16mm, the length of the first riser reactor is 3200mm, the internal diameter of the second riser reactor is 16mm, the height of the second riser reactor is 3000mm, the internal diameter of the fluidized bed reactor is 64mm, and the height of the fluidized bed reactor is 500mm.

Introducing raw oil into the bottom of a first riser reactor, contacting with a catalyst CAT-3 and reacting, and separating the reacted oil mixture by a cyclone separator;

introducing C4 hydrocarbon into the bottom of the second riser reactor, contacting with a catalyst CAT-3 and reacting, introducing the generated oil mixture into the fluidized bed reactor, separating the reacted oil mixture by a quick separation device to obtain reaction oil gas and a regenerated catalyst, introducing the reaction oil gas into a cyclone separator, and introducing the regenerated catalyst into a stripper for stripping.

The catalyst enters a stripper and then enters a regenerator for regeneration, the regenerated catalyst returns to the riser reactor for recycling, and oil gas is introduced into a fractionation system for separation. Wherein the mass ratio of the C4 hydrocarbon to the raw oil is 0.1. The reaction conditions and results are shown in Table 4.

Comparative example 3

The test was carried out on a medium-sized test apparatus. The apparatus comprises two riser reactors and a fluidized bed reactor. The internal diameter of the first riser reactor is 16mm, the length of the first riser reactor is 3200mm, the internal diameter of the second riser reactor is 16mm, the height of the second riser reactor is 3000mm, and the internal diameter of the fluidized bed reactor is 64mm and the height of the fluidized bed reactor is 500mm.

Raw oil is introduced into the bottom of a first riser reactor, contacts with a catalyst CAT-3 and reacts, a generated oil agent mixture is introduced into a fluidized bed reactor, the reacted oil agent mixture is separated through a quick separation device to obtain reaction oil gas and a regenerated catalyst, the reaction oil gas is introduced into a cyclone separator, and the regenerated catalyst is introduced into a stripper for steam stripping.

Comparative example 4

The process of comparative example 3 was followed except that the C4 hydrocarbons obtained by fractionation were not introduced into the second riser reactor, and the light gasoline fraction (distillation range 40-80 ℃, olefin content 65 wt%) obtained by fractionation was introduced into the fluidized bed reactor 4-1, wherein the mass ratio of light gasoline to feed oil was 0.1. The reaction conditions and results are shown in Table 4.

TABLE 4

As can be seen from tables 3 and 4, compared with the comparative example, the method and apparatus provided by the present disclosure can improve the processing amount of the light raw material, realize the independent zone control of different reaction regeneration systems, and achieve the coupling utilization of the heat inside the apparatus, and obtain higher yields of ethylene, propylene and light aromatics.

The preferred embodiments of the present disclosure have been described above in detail, however, the present disclosure is not limited to the specific details of the above embodiments, and various simple modifications may be made to the technical solution of the present disclosure within the technical idea of the present disclosure, and these simple modifications all belong to the protection scope of the present disclosure.

It should be noted that the various features described in the foregoing embodiments may be combined in any suitable manner without contradiction. To avoid unnecessary repetition, the disclosure does not separately describe various possible combinations.

In addition, any combination of various embodiments of the present disclosure may be made, and the same should be considered as the disclosure of the present disclosure as long as it does not depart from the gist of the present disclosure.

Claims

1. A catalytic conversion process for producing ethylene, propylene and light aromatics, the process comprising:

2. A catalytic conversion process according to claim 1,

the heavy raw material is selected from one or more of vacuum wax oil, atmospheric residue oil, vacuum residue oil, coking wax oil, deasphalted oil, furfural refined raffinate oil, coal liquefied oil, oil sand oil, shale oil, distillate oil obtained by F-T synthesis and animal and vegetable oil;

the light feedstock comprises at least one of C4 hydrocarbons and a light gasoline fraction;

the C4 hydrocarbon is a mixture of C4 alkane and C4 alkene; the C4 hydrocarbon has a content of C4 olefins of greater than 50 wt%, preferably greater than 60 wt%;

the light gasoline fraction is light gasoline fraction rich in olefin, and the olefin content in the light gasoline fraction is 20-95 wt%, preferably 50-95 wt%;

the weight ratio of the light feedstock to the heavy feedstock is from 0.02 to 0.6, preferably from 0.04 to 0.3;

the weight ratio of said C4 hydrocarbons to said heavy feedstock is from 0 to 0.30, preferably from 0.02 to 0.15; the weight ratio of the light gasoline fraction to the heavy feedstock is from 0 to 0.30, preferably from 0.02 to 0.15.

3. A catalytic conversion process according to claim 1,

the first reactor comprises a first riser reactor and/or a first fluidized bed reactor, preferably a combined reactor of the first riser reactor and the first fluidized bed reactor; in the combined reactor, the first fluidized bed reactor is positioned above the first riser reactor and is communicated with the first riser reactor.

4. A catalytic conversion process according to claim 3,

the reaction temperature of the first riser reactor is 500-580 ℃, preferably 520-560 ℃; the agent-oil ratio is 2-25, preferably 5-20; the reaction time is 0.5 to 10 seconds, preferably 1 to 5 seconds;

the reaction temperature of the first fluidized bed reactor is 480-560 ℃, and preferably 500-540 ℃; the weight hourly space velocity is 1-20 hours ^-1 Preferably 2 to 8 hours ^-1 (ii) a The density of the catalyst is 50-350kg/m ³ Excellence inSelecting 100-250kg/m ³ (ii) a The bed height of the first fluidized bed reactor is 1/2-4/5 of the height of the first fluidized bed reactor, and preferably 1/2-3/4 of the height of the first fluidized bed reactor; the pressure in the first fluidized bed reactor is 0.1-0.4MPa, preferably 0.15-0.3MPa.

5. A catalytic conversion process according to claim 1,

the second reactor is a combined reactor of a second riser reactor and a second fluidized bed reactor, and the second fluidized bed reactor is positioned above the second riser reactor and communicated with the second riser reactor.

6. A catalytic conversion process according to claim 5,

the reaction temperature of the second riser reactor is 630-690 ℃, preferably 650-670 ℃; the agent-oil ratio is 20-80, preferably 40-60; the reaction time is 0.5 to 15 seconds, preferably 1 to 10 seconds;

the reaction temperature of the second fluidized bed reactor is 590-650 ℃, preferably 610-630 ℃; the weight hourly space velocity is 1-20 hours ^-1 Preferably 2 to 8 hours ^-1 (ii) a The density of the catalyst in the second fluidized bed reactor is 50-350kg/m ³ Preferably 100 to 250kg/m ³ (ii) a The bed height of the second fluidized bed reactor is 1/2-4/5 of the height of the second fluidized bed reactor, and preferably 1/2-3/4 of the height of the second fluidized bed reactor; the pressure in the second fluidized bed reactor is 0.1-0.4MPa, preferably 0.15-0.3MPa.

7. Catalytic conversion process according to claim 1, wherein the temperature of the preheated heavy feedstock is in the range of 100-350 ℃, preferably 150-300 ℃.

8. A catalytic conversion process according to claim 1,

the first regeneration reaction comprises complete regeneration or incomplete regeneration, preferably incomplete regeneration; the regeneration temperature in the first regenerator is 620-720 ℃, preferably 650-680 ℃;

the second regeneration reaction comprises complete regeneration or incomplete regeneration, preferably complete regeneration; the regeneration temperature in the second regenerator is 650-750 ℃, preferably 680-730 ℃.

9. A catalytic conversion process according to claim 1,

the heavy oil catalyst comprises the Y-type molecular sieve, the beta-type molecular sieve, a first clay, and a first binder; the content of the Y-type molecular sieve is 20-60 wt%, preferably 30-50 wt%, based on the total weight of the heavy oil catalyst; the content of the beta-type molecular sieve is 1-40 wt%, preferably 1-20 wt%; the content of the first clay is 10 to 70 wt%, preferably 15 to 45 wt%; the content of the first binder is 10 to 40 wt%, preferably 20 to 35 wt%;

the Y-type molecular sieve is selected from at least one of HY type molecular sieve, USY type molecular sieve, REUSY type molecular sieve, REY type molecular sieve, REHY type molecular sieve, DASY type molecular sieve and REDASY type molecular sieve;

the beta-type molecular sieve is a beta-type molecular sieve modified by phosphorus and a transition metal M, wherein M is selected from at least one of Fe, co, ni, cu, mn, zn and Sn;

the first clay is at least one selected from kaolin, montmorillonite and bentonite; the first binder is at least one selected from silica sol, aluminum sol and pseudo-boehmite, and preferably, the first binder is double-aluminum binder of the aluminum sol and the pseudo-boehmite.

10. A catalytic conversion process according to claim 1,

the light oil catalyst comprises the MFI structure molecular sieve, a second clay and a second binder; the content of the MFI structure molecular sieve is 20-60 wt%, preferably 30-50 wt%, based on the total weight of the light oil catalyst; the content of the second clay is 10 to 70 wt%, preferably 15 to 45 wt%; the content of the second binder is 10-40 wt%, preferably 20-35 wt%;

the MFI structure molecular sieve is selected from at least one of ZRP molecular sieve, ZRP molecular sieve containing phosphorus, ZRP molecular sieve containing rare earth, ZRP molecular sieve containing phosphorus and alkaline earth metal and ZRP molecular sieve containing phosphorus and transition metal, preferably ZRP zeolite containing phosphorus and rare earth; the second clay is at least one of kaolin, montmorillonite and bentonite; the second binder is at least one selected from silica sol, aluminum sol and pseudo-boehmite, and preferably, the second binder is double-aluminum binder of the aluminum sol and the pseudo-boehmite.

11. A catalytic conversion process according to claim 1, wherein the process further comprises: and separating the second reaction oil gas after heat exchange to obtain C4 hydrocarbon containing olefin and a light gasoline fraction, and returning part or all of the C4 hydrocarbon containing olefin and the light gasoline fraction to the second reactor.

12. An apparatus suitable for use in the catalytic conversion process for producing ethylene, propylene and light aromatic hydrocarbons according to any one of claims 1 to 11, the catalytic conversion apparatus comprising a first reactor, a second reactor, a first regenerator, a second regenerator and a feedstock preheater,

the first reactor is connected with the first regenerator through a first catalyst conveying pipeline; the second reactor is connected with the second regenerator through a second catalyst conveying pipeline;

the first reactor comprises a preheated heavy raw material inlet, the second reactor comprises a light raw material reaction oil gas outlet, and the raw material preheater comprises a reaction oil gas inlet before heat exchange, a reaction oil gas outlet after heat exchange, a heavy raw material inlet before preheating and a heavy raw material outlet after preheating; the light raw material reaction oil gas outlet is connected with the pre-heat exchange reaction oil gas inlet, and the preheated heavy raw material outlet is connected with the preheated heavy raw material inlet of the first reactor;

the first regenerator and the second regenerator are arranged in parallel or coaxially, preferably in parallel; the first regenerator includes a first regenerated flue gas outlet, the second regenerator includes a first regenerated flue gas inlet, and the first regenerated flue gas outlet is connected to the first regenerated flue gas inlet.