CN113735676B

CN113735676B - Method for high-selectivity catalytic pyrolysis of high-yield propylene and high-yield gasoline

Info

Publication number: CN113735676B
Application number: CN202010473053.0A
Authority: CN
Inventors: 朱金泉; 朱根权; 王新; 首时; 刘守军; 杨义华; 侯典国
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: 2020-05-29
Filing date: 2020-05-29
Publication date: 2023-10-13
Anticipated expiration: 2040-05-29
Also published as: CN113735676A

Abstract

The present disclosure relates to a method and system for high selectivity catalytic cracking to produce more propylene and increase gasoline yield. The method comprises the following steps: reacting heavy raw materials in a heavy oil cracking reactor and an enhanced catalytic conversion reactor, and separating to obtain a first product; separating the first product from the second product and the third product to obtain different fractions and the like; introducing the C4 liquefied gas fraction into a multi-branched olefin synthesis reactor for reaction, and separating to obtain an isobutene-lean C4 material flow and a multi-branched olefin product; introducing the isobutene-lean C4 stream and the light gasoline fraction into a high-cracking olefin synthesis reactor for reaction to obtain a high-cracking olefin product and an olefin-lean recycle stream; reacting the olefin-lean recycle stream, the light gasoline fraction and the high-cracking olefin product in a light hydrocarbon cracking reactor and an enhanced catalytic conversion reactor, and separating to obtain the second product; injecting the multi-branched olefin product into the upper part of a heavy oil cracking reactor or an enhanced catalytic conversion reactor, and separating to obtain a third product; mixing some fractions to obtain the product gasoline. The method can greatly reduce the reaction severity, greatly improve the yield of the catalytic cracking propylene and increase the yield of the high-octane gasoline.

Description

Method for high-selectivity catalytic pyrolysis of high-yield propylene and high-yield gasoline

Technical Field

The present disclosure relates to a method for high selectivity catalytic cracking to produce more propylene and increase gasoline yield.

Background

In the automobile gasoline standard GB17930-2016 of the people's republic of China, the technical requirement on the automobile gasoline is that the olefin content is not more than 15% (volume fraction), the aromatic hydrocarbon content is not more than 35% (volume fraction), the benzene content is not more than 0.8% (volume fraction), and the oxygen content is not more than 2.7% (mass fraction). According to the requirements of national delivery and change commission and the like, the ethanol gasoline for vehicles is promoted and used nationwide by 2020, and according to the current promotion of E10 ethanol gasoline (containing 10% ethanol), the vehicle gasoline market is expected to be saturated. At the same time, MTBE is limited to be added into motor gasoline, and a new outlet is required to be searched for as a raw material C4 olefin of MTBE.

Modern petroleum processing technology pursues high yields of high value products (e.g., ethylene, propylene, C8 aromatics), and also places greater emphasis on reducing unit feedstock processing energy consumption and reducing carbon emissions. The C4 olefin is catalytically cracked to produce more ethylene and propylene. For example, in the 2 nd period of volume 36 in 2005 of journal "petroleum refining and chemical industry", the study on the preparation of propylene and ethylene by catalytic cracking of butene "reports that C4 olefins can be used for producing ethylene and propylene by catalytic cracking.

There are also patent disclosures of a reaction process for producing ethylene and propylene by a superposition reaction and a catalytic cracking reaction using C4 or liquefied gas as a raw material and a catalyst thereof.

Patent CN102531824B discloses a process for preparing propylene and ethylene from liquefied gas containing butene: (1) The liquefied gas exchanges heat with the reacted oil-gas mixture and/or is directly heated to the preheating temperature of 150-450 ℃; (2) Liquefied gas enters the upper section of the reactor from the top of the reactor, olefin components in the liquefied gas undergo a superposition reaction under the action of a superposition catalyst to generate macromolecular hydrocarbon, and meanwhile, the temperature of the oil gas is increased due to exothermic reaction; (3) The oil-gas mixture generated by the superposition reaction enters the lower section of the reactor, and is subjected to cracking reaction under the action of a cracking catalyst to generate a hydrocarbon mixture containing target products of propylene and ethylene, and the hydrocarbon mixture flows out from the bottom of the reactor and enters a subsequent separation system to recover propylene, ethylene and aromatic hydrocarbon oil generated by the reaction. In the process technology for preparing propylene and ethylene from liquefied gas, two or more sections of catalyst beds are adopted, so that raw material liquefied gas is contacted with two or more catalysts in sequence, and a superposition reaction and a cracking reaction are carried out in sequence, thereby greatly improving the conversion rate of the raw material liquefied gas and the selectivity of propylene and ethylene. However, the reactor provided in the technology has an upper section and a lower section arranged in series, the reaction conditions of the upper section and the lower section can interfere with each other, the high pressure and the low temperature are required for the superposition reaction, and the low pressure and the high temperature are required for the cracking reaction. Meanwhile, when the liquefied gas containing C4 passes through the upper section of the reactor once, the complete conversion is difficult to achieve.

The patent CN100537721C discloses a catalytic conversion method for increasing propylene yield, which comprises the steps of injecting preheated raw oil into a main riser of a double-riser reaction regeneration system, contacting with a hot catalyst for catalytic cracking reaction, separating a reaction product, and recycling a raw catalyst after steam stripping and regeneration; and injecting the liquefied gas product separated by the gas separation system into an auxiliary riser to contact with a hot catalyst, sequentially carrying out superposition reaction, catalytic cracking and alkane dehydrogenation reaction in two reaction areas in the auxiliary riser, separating reaction products, and recycling the regenerated catalyst. The method provided by the invention can further convert the liquefied gas product after propylene removal into propylene, and obviously improve the propylene yield on the premise of not increasing the liquefied gas yield. The auxiliary lifting pipe is a first reaction zone, a second reaction zone, an outlet zone and a horizontal pipe which are coaxial with each other in sequence from bottom to top, the horizontal pipe is connected with the settler, and the lower parts of the first reaction zone and the second reaction zone are respectively connected with a catalyst inlet pipe; the operating conditions are as follows: the reaction temperature of the main riser reactor is 450-650 ℃; the weight ratio of the catalyst to the raw oil is 1-25, the reaction time is 0.5-30 seconds, and the pressure (absolute pressure) in the main riser reactor is 0.1-0.4 MPa; the temperature of the first reaction zone of the auxiliary riser reactor is 150-450 ℃, the reaction time is 0.5-2.0 seconds, the weight ratio of the catalyst to the raw material gas is 1-30, the temperature of the second reaction zone is 450-650 ℃, the reaction time is 3-20 seconds, the weight ratio of the catalyst to the raw material gas is 3-60, and the pressure (absolute pressure) of the auxiliary riser reactor is 0.1-0.4 MPa.

Also, in patent CN100448954C, a catalytic conversion method for increasing propylene yield is disclosed, preheated raw oil is injected into a main riser of a double riser reaction regeneration system, contacted with a hot catalyst to perform catalytic cracking reaction, reaction products are separated, and the regenerated raw oil is recycled; injecting the liquefied gas product after propylene separation into an auxiliary riser, contacting with a hot catalyst, sequentially carrying out olefin superposition, superposition product cracking and alkane dehydrogenation reaction, separating reaction products, and recycling after regeneration of a generating agent; the catalyst is a mixture of two catalysts: the weight ratio of the dry basis of the first cracking catalyst and the second cracking catalyst is 10-70:30-90. The method provided by the invention can further convert the liquefied gas product after propylene removal into propylene, and obviously improve the propylene yield on the premise of not increasing the liquefied gas yield. The technology is characterized in that the olefin superposition, the superposition product cracking and the alkane dehydrogenation reaction are completed at the same time in the auxiliary riser. The auxiliary lifting pipe is provided with a first reaction zone (150-450 ℃) to promote the light olefin to carry out the superposition reaction and a second reaction zone (450-650 ℃) to promote the further pyrolysis of the superposition product and the dehydrogenation of propane to generate propylene.

These documents mainly focus on how to carry out the polymerization reaction of liquefied gas or butene containing olefin to produce gasoline fraction or diesel fraction, and how to further catalyze and crack the polymerization product of liquefied gas or C4 olefin to produce ethylene and propylene has less report. Moreover, the superposition products of C4 olefins are mainly C8 olefins, and the propylene yield and propylene selectivity of the direct catalytic cracking are not high, because the C8 olefins synthesized from the C4 stack are easier to be further catalytically cracked into 2C 4 olefins, and the synthetic C8 olefins are also easier to undergo aromatization reaction in the catalytic cracking process to produce coke, and the coke yield is higher. The mass fraction of olefins in the synthesized C8 olefins is 80-100%, and the direct use as a gasoline blending component necessarily results in an increase in the olefin content of the finished gasoline, so that further economic and reasonable treatment is required.

Therefore, a combined process of light olefin polymerization and selective catalytic cracking to produce more propylene needs to be developed.

Disclosure of Invention

The present disclosure provides a method for high selectivity catalytic cracking to produce more propylene and increase gasoline yield, which is characterized in that,

the method is carried out in a catalytic cracking device comprising at least three reactors: a heavy oil cracking reactor, a light hydrocarbon cracking reactor and an enhanced catalytic conversion reactor;

The method comprises the following steps:

contacting a heavy feedstock in a heavy oil cracking reactor with a first catalytic cracking catalyst to perform a first catalytic cracking reaction to obtain a first cracking reaction mixture comprising the first catalytic cracking catalyst; subjecting the first cleavage reaction mixture to a further catalytic reaction in an enhanced catalytic conversion reactor;

contacting at least a portion of the olefin lean recycle stream, at least a portion of the light gasoline fraction, at least a portion of the high cracking olefin product in a light hydrocarbon cracking reactor with a second catalytic cracking catalyst to perform a second catalytic cracking reaction to obtain a second reaction mixture comprising the second catalytic cracking catalyst; subjecting the second reaction mixture to a further catalytic reaction in an enhanced catalytic conversion reactor;

simultaneously, the multi-branched olefin product is injected into the upper part of a heavy oil cracking reactor or an enhanced catalytic conversion reactor to carry out catalytic conversion reaction to obtain a third reaction mixture;

carrying out gas-agent separation on the obtained first reaction mixture, second reaction mixture and second reaction mixture in the settler to obtain a first carbon deposit catalyst, a second carbon deposit catalyst, a first product, a second product and a third product;

Introducing the first product, the second product and the third product into a catalytic product separation system for separation to obtain a dry gas fraction, a C3 liquefied gas fraction, a C4 liquefied gas fraction, a light gasoline fraction, a heavy gasoline fraction, a diesel oil fraction and a heavy oil fraction;

introducing the C4 liquefied gas fraction into a multi-branched olefin synthesis reactor for a first synthesis reaction to obtain a first synthesis product oil gas, and further introducing the first synthesis product oil gas into a first synthesis product separation system for separation to obtain an isobutene-poor C4 material flow and the multi-branched olefin product;

introducing at least a part of the isobutene-lean C4 material flow and/or at least a part of the light gasoline fraction into a high-cracking olefin synthesis reactor for a second synthesis reaction to obtain a second synthesis product oil gas, and further introducing the second synthesis product oil gas into a second synthesis product separation system for separation to obtain the olefin-lean recycle material flow and the high-cracking olefin product;

the multi-branched olefin products are returned to a heavy oil cracking reactor or an enhanced catalytic conversion reactor for reaction; the high-cracking olefin product returns to the light hydrocarbon cracking reactor for reaction;

mixing a portion of the light gasoline fraction, a portion of the olefin-lean recycle stream, and the heavy gasoline fraction can result in a product gasoline.

In one embodiment, the reaction conditions of the first synthesis reaction are: the reaction temperature is 300-450 ℃, the pressure is 0.5-2.0 MPa, and the mass airspeed is 1-5 h ^-1 The method comprises the steps of carrying out a first treatment on the surface of the The reaction conditions of the second synthesis reaction are as follows: the reaction temperature is 180-280 ℃, the pressure is 0.8-2.0 MPa, and the mass airspeed is 0.9-4.0 h ^-1 。

In one embodiment, the first synthesis reaction catalyst includes 1 to 12 mass% NiO, 45 to 82 mass% amorphous aluminum silicate, and 10 to 50 mass% alumina; the second synthesis reaction catalyst includes 1 to 20 mass% of NiO, 40 to 80 mass% of HZSM-5 zeolite, and 10 to 50 mass% of alumina.

In one embodiment, the multi-branched olefin product has a distillation range of from 90 to 150 ℃, preferably from 100 to 130 ℃, and is predominantly trimethylpentene; the high-cracking olefin product has a distillation range of 130-330 ℃, preferably 140-290 ℃ and is mainly C9-C12 olefin.

In one embodiment, the light gasoline fraction has a distillation range of 9 to 150 ℃, more preferably 9 to 100 ℃, still more preferably 9 to 60 ℃; the light gasoline fraction has an olefin content of 20 to 90 wt%, preferably 35 to 90 wt%, based on the total weight of the light gasoline fraction.

In one embodiment, the weight ratio of the multi-branched olefin product to the heavy feedstock is from 0.01 to 0.3:1, preferably 0.05 to 0.15:1, a step of; the weight ratio of the high-cracking olefin product to the heavy raw material is 0.01-0.4: 1, preferably 0.05 to 0.2:1.

in one embodiment, the heavy feedstock is at least one selected from the group consisting of petroleum hydrocarbon oils, synthetic oils, coal liquefaction oils, oil sand oils, and shale oils, preferably petroleum hydrocarbon oils, which are at least one selected from the group consisting of atmospheric wax oil, vacuum wax oil, coker wax oil, deasphalted oil, hydrogenated tail oil, atmospheric residuum, vacuum residuum, and crude oil, hydrogenated wax oil, hydrogenated residuum.

In one embodiment, the operating conditions of the first catalytic cracking reaction include: the reaction temperature is 480-600 ℃; the reaction time is 0.5-10 seconds; the weight ratio of the agent to the oil is 5-15: 1, a step of; the weight ratio of water to oil is 0.05-1: 1.

in one embodiment, the operating conditions of the second catalytic cracking reaction include: the reaction temperature is 520-750 ℃; the reaction time is 0.1-3 seconds; the weight ratio of the agent to the oil is 6-40: 1, a step of; the weight ratio of water to oil is 0.1-1: 1.

in one embodiment, the enhanced catalytic conversion reactor is a fluidized bed reactor, and the operating conditions of the enhanced catalytic conversion reactor are: the reaction temperature is 450-750 ℃, preferably 510-560 ℃; weight hourly space velocity of 1-30 h ^-1 。

In one embodiment, when the multi-branched olefin product is injected into the heavy oil cracking reactor, the injection point should be located after the injection point of the heavy feedstock, with the total length from the injection point of the heavy feedstock to the outlet of the heavy oil cracking reactor being 100%, the injection point of the heavy feedstock being 0%, and the multi-branched olefin product being injected into the heavy oil cracking reactor at a position of 50% -100%;

wherein when the multi-branched olefin product is injected into the enhanced catalytic conversion reactor, it is preferred that its injection point is located at the bottom of the enhanced catalytic conversion reactor;

when the multi-branched olefin product is injected into a heavy oil cracking reactor or an enhanced catalytic conversion reactor, the operating conditions are: the reaction temperature is 450-600 ℃, preferably 460-550 ℃; weight hourly space velocity of 1-30 h ^-1 。

In one embodiment, the first and second catalytic cracking catalysts each contain a shape selective zeolite having an average pore size of less than 0.7 nanometers, the shape selective zeolite being at least one selected from the group consisting of zeolite having an MFI structure, ferrierite, chabazite, cyclote, erionite, a-type zeolite, column zeolite, and turbid zeolite.

In one embodiment, each of the heavy oil cracking reactor and the light hydrocarbon cracking reactor is one selected from a riser reactor, a downer reactor, a fluidized bed reactor, a riser-downer combined reactor, a riser-fluidized bed combined reactor, and a downer-fluidized bed combined reactor.

In one embodiment, the first and second carbon catalysts are stripped in a stripping zone of the settler, and the stripped first and second carbon catalysts are regenerated in a regenerator to obtain first and second regenerated catalysts, respectively; and sending the first regenerated catalyst into the heavy oil cracking reactor to serve as the first catalytic cracking catalyst, and sending the second regenerated catalyst into the light hydrocarbon cracking reactor to serve as the second catalytic cracking catalyst.

In one embodiment, the method further comprises: separating the olefin lean recycle stream into a C4 hydrocarbon fraction and an alkane-rich light gasoline fraction, combining the alkane-rich light gasoline fraction with the heavy gasoline fraction, a portion of the light gasoline fraction, and a portion of the high cracker olefin product to yield a low olefin content gasoline product.

In one embodiment, the regeneration is performed at a temperature of 600 to 800 ℃.

In one embodiment, the temperature of the first regenerated catalyst is from 560 ℃ to 800 ℃; the temperature of the second regenerated catalyst is 560-800 ℃.

In one embodiment, the first char catalyst is stripped in a first stripping zone of the settler and the second char catalyst is stripped in a second stripping zone of the settler.

In another aspect, the present invention provides a system for high selectivity catalytic cracking of high propylene yield and high gasoline yield, comprising:

a reactor system, the reactor system comprising:

the heavy oil cracking reactor is provided with a plurality of heavy oil cracking reactor raw material inlets, a heavy oil cracking reactor catalyst inlet and a heavy oil cracking reactor product outlet; the heavy oil cracking reactor feed inlet comprises a heavy feed inlet, and optionally a multi-branched olefin product inlet;

the light hydrocarbon cracking reactor is provided with a plurality of light hydrocarbon cracking reactor raw material inlets, a light hydrocarbon cracking reactor catalyst inlet and a light hydrocarbon cracking reactor product outlet; and

the enhanced catalytic conversion reactor is provided with an enhanced catalytic conversion reactor raw material inlet and an enhanced catalytic conversion reactor product outlet; the material inlet of the enhanced catalytic conversion reactor is communicated with the product outlet of the heavy oil cracking reactor, and the product outlet of the light hydrocarbon cracking reactor is communicated with the middle lower part of the enhanced catalytic conversion reactor, so that material flows from the heavy oil cracking reactor and the light hydrocarbon cracking reactor enter the enhanced catalytic conversion reactor;

A catalytic product separation system provided with a feedstock inlet and a plurality of fraction outlets; the raw material inlet of the catalytic product separation system is communicated with the product outlet of the enhanced catalytic conversion reactor; the plurality of fraction outlets of the catalytic product separation system comprise a dry gas fraction outlet, a C3 liquefied gas fraction outlet, a C4 liquefied gas fraction outlet, a light gasoline fraction outlet, a heavy gasoline fraction outlet, a diesel fraction outlet and a heavy oil fraction outlet;

the multi-branched olefin synthesis reactor is provided with a raw material inlet of the multi-branched olefin synthesis reactor and a product outlet of the multi-branched olefin synthesis reactor, and the raw material inlet of the multi-branched olefin synthesis reactor is communicated with a C4 liquefied gas fraction outlet of the catalytic product separation system;

a first synthetic product separation system provided with a first synthetic product separation system raw material inlet and a plurality of first synthetic product separation system separation product outlets; the raw material inlet of the first synthesis product separation system is communicated with the product outlet of the multi-branched olefin synthesis reactor; the plurality of first synthesis product separation system separation product outlets comprises an isobutylene-lean C4 stream outlet and a multi-branched olefin product outlet;

A high-cracking olefin synthesis reactor provided with a high-cracking olefin synthesis reactor raw material inlet and a high-cracking olefin synthesis reactor product outlet; the high-cracking olefin synthesis reactor raw material inlet is communicated with the isobutene-lean C4 stream outlet;

the second synthetic product separation system is provided with a second synthetic product separation system raw material inlet and a plurality of second synthetic product separation system separation product outlets; the raw material inlet of the second synthesis product separation system is communicated with the product outlet of the high-cracking olefin synthesis reactor; the separation product outlets of the second synthetic product separation systems comprise an olefin-lean circulation flow outlet and a high-cracking olefin product outlet, and the olefin-lean circulation flow outlet and the high-cracking olefin product outlet are respectively communicated with the raw material inlet of one light hydrocarbon cracking reactor;

wherein, the raw material inlet of one light hydrocarbon cracking reactor is also communicated with the light gasoline fraction outlet of the catalytic product separation system; and

the optional multi-branched olefin product inlet of the heavy oil cracking reactor is optionally in communication with a multi-branched olefin product outlet of the first synthesis product separation system.

The inventor of the present disclosure researches and discovers that the performances of different C4 olefins (n-butene, isobutene, trans-2-butene, cis-2-butene) after superposition and the superposition products have larger difference in the catalytic cracking and the multi-production of ethylene and propylene; the products obtained by the same C4 olefin through different deep superposition reactions have larger difference in the performances of producing more ethylene and propylene through catalytic cracking; different kinds of olefins in the liquefied gas are overlapped to obtain products, and the performances of the products for producing more ethylene and propylene by catalytic cracking are different. Meanwhile, the different reaction processes of catalytic cracking have decisive influence on the performance of cracking the low-carbon olefin polymerization product to produce more ethylene and propylene.

The method for producing more propylene and increasing the yield of gasoline by high-selectivity catalytic pyrolysis provided by the disclosure overcomes the problems of high dry gas yield, poor propylene selectivity and high olefin content in the gasoline product in the heavy oil catalytic pyrolysis reaction in the existing process route. The inventor of the present disclosure has studied and found that, when catalytically cracked C4 fractions and/or light gasoline are first subjected to a polymerization reaction to produce a polymerization product, the selectivity of propylene produced by the polymerization reaction of different polymerization products is greatly different, for example, the selectivity of propylene produced by the polymerization reaction of C4 olefins is greatly different, for example, the product C8 olefins produced by the polymerization of C8 olefins are more likely to crack into butenes rather than into propylene if the branched chain number (methyl number) of C8 olefins is greater, for example, the methyl number of n-octene, methyl heptene, dimethyl-hexene and trimethyl-pentene is gradually increased, and the yield of propylene produced by the catalytic cracking is gradually reduced. Although both are C8 olefins, the propylene yield produced by the catalytic cracking of n-octene is more than twice that produced by the catalytic cracking of trimethylpentene. Therefore, C8 olefin (mainly trimethylpentene) synthesized by superposition of isobutene in C4 is more suitable to be used as a gasoline blending component, but because the mass fraction of olefin in the synthesized C8 olefin is very high, the synthesized C8 olefin is not suitable for directly blending finished gasoline, and economic and reasonable treatment is needed to reduce the olefin content. Different C4 olefins (n-butene, isobutene, trans-2-butene and cis-2-butene) are contained in the C3-removed liquefied gas, a certain C4 olefin is required to be overlapped to generate a specific type of overlapped product, and then different process routes are adopted to ensure that propylene selectivity of catalytic cracking of the overlapped product is higher.

In particular, the present disclosure has any one or more, preferably all, of the following beneficial effects as compared to conventional methods for catalytically converting heavy hydrocarbon oils to produce higher olefins:

1. in the prior art, the C4 fraction and the light gasoline are directly recycled to the catalytic cracking or catalytic cracking device for reaction, but the C4 fraction and the light gasoline are short in molecular chain and difficult to crack, and the cracking needs very high reaction severity, namely high reaction temperature, catalyst-oil ratio, residence time and the like, so that the higher dry gas and coke yield and the poor propylene selectivity are caused. Therefore, in the prior art, although the propylene can be increased by circularly cracking the C4 fraction and the light gasoline fraction through a catalytic cracking device, the single pass conversion rate of the reaction is lower, a larger recycle ratio is needed, the propylene selectivity is lower, the dry gas selectivity is higher, the olefin in the product gasoline can not be completely converted, and the olefin content of the product gasoline is higher.

The present disclosure first produces C8 olefins (mainly trimethylpentene) from isobutene in the C4 fraction by selective polymerization, and simultaneously separates from other C4 olefins (such as n-butene, trans-2-butene, cis-2-butene), and is not suitable for direct catalytic cracking to produce propylene in high yield because of low propylene yield produced by catalytic cracking of trimethylpentene. The synthesized C8 olefin has the advantages that most of C8 olefin can undergo hydrogen transfer reaction and aromatization reaction at lower reaction temperature and lower weight hourly space velocity, the mass fraction of the olefin is reduced, trimethylpentene is converted into trimethylpentane or dimethylbenzene, a high-octane gasoline component can be obtained, and a new hydrogenation saturation device is not needed. And partial C8 olefins are cracked to generate a small amount of isobutene, the isobutene can be subjected to isomerisation reaction under the catalytic cracking condition to partially convert other C4 olefins (such as n-butene, trans-2-butene and cis-2-butene), and the partial C4 olefins (such as n-butene, trans-2-butene and cis-2-butene) can be recovered by a separation system to synthesize a superposition product which is easy to crack into propylene. Thus, the problem of low yield of propylene by catalytic cracking of isobutene in C4 olefins is solved, meanwhile, the superposition reaction of isobutene and other C4 olefins is avoided, the catalytic cracking reaction of other C4 olefins is interfered, and the gasoline component with high octane number and low olefin can be obtained.

2. The separated other C4 olefins (such as n-butene, trans-2-butene and cis-2-butene) or/and olefins in the light gasoline are further synthesized into hydrocarbons such as C9, C10, C12 and C13 which are easy to crack into propylene, so that the severity of the catalytic cracking reaction of other C4 olefins is greatly reduced, and the selectivity of cracking recycle streams into propylene is greatly improved.

While the C4 fraction and the light gasoline contain alkane and alkene, alkane cracking needs higher reaction severity than alkene cracking, and the prior art circularly cracks the C4 fraction and the light gasoline containing alkane and alkene at the same time, so that the requirements of the alkane and the alkene on reaction conditions are difficult to be considered: the reaction severity is low, and alkane is difficult to react; the reaction severity is high, and more thermal cracking reaction of olefin can occur to generate dry gas.

According to the method, the C4 olefin and the light gasoline olefin are firstly converted and separated through different superposition reactions, the residual C4 hydrocarbon and light naphtha which are enriched in alkane are returned to the catalytic cracking reactor in the form of olefin-lean circulation flow, and firstly, high-severity dehydrogenation and cracking reactions occur, so that the alkane cracking effect is improved. The highly cracked olefin products (C9, C10, C12, C13 olefin, etc.) which are easier to crack enter a regenerated catalytic cracking catalyst with reduced temperature to react, thereby greatly reducing the reaction severity and improving the selectivity of propylene.

The method realizes the classified conversion of C4 molecules with different molecular structures into high-octane gasoline components or the high-selectivity cracking into propylene.

Drawings

The accompanying drawings are included to provide a further understanding of the disclosure, and are incorporated in and constitute a part of this specification, illustrate the disclosure and together with the description serve to explain, but do not limit the disclosure. In the drawings:

fig. 1 is a schematic flow chart of a method for high selectivity catalytic cracking of high propylene yield and high gasoline yield provided by the present disclosure.

Description of the reference numerals

1. Heavy oil cracking reactor

2. Light hydrocarbon cracking reactor

3. Enhanced catalytic conversion reactor

7. Stripping zone

6. Sedimentation device

9. Regenerator device

61. Catalytic product separation system

71. Multi-branched-chain olefin synthesis reactor

72. High-cracking olefin synthesis reactor

62. First synthetic product separation system

63. Second synthetic product separation system

11. First catalytic cracking catalyst delivery line

12. Second catalytic cracking catalyst delivery line

17. Spent catalyst delivery line

21. Pipeline (injection heavy raw material)

20. Pipeline (conveying catalytic cracking reaction oil gas)

24. Pipeline (conveying dry gas)

25. Pipeline (conveying C3)

26. Pipeline (conveying C4 liquefied gas)

27. Pipeline (conveying light petrol)

30. Pipeline (conveying heavy petrol)

31. Pipeline (Diesel delivery)

32. Pipeline (conveying heavy oil)

41. Pipeline (injection atomizing steam)

43. Pipeline (injection atomizing steam)

44. Pipeline (injection atomizing steam)

45. Pipeline (injection atomizing steam)

47. Pipeline (injection steam)

51. Pipeline (injection pre-lifting medium)

52. Pipeline (injection pre-lifting medium)

90. Pipeline (injection air)

33. Pipeline (first synthetic product oil gas)

36. Pipeline (second synthetic product oil gas)

34. Pipeline (lean isobutene C4 stream)

38. Pipeline (lean olefin recycle stream)

35. Pipeline (conveying multi-branched olefin products)

37. Pipeline (delivering high cracking olefin products)

90. Pipeline (conveying air)

91. Pipeline (conveying smoke)

100. External heater (taking the excess heat of the regenerator, reducing the regeneration temperature)

Detailed Description

The technical scheme of the invention is further described below according to specific embodiments. The scope of the invention is not limited to the following examples, which are given for illustrative purposes only and do not limit the invention in any way.

The present disclosure provides a method for high selectivity catalytic cracking of high propylene yield and high gasoline yield, the method being performed in a catalytic cracking device comprising at least three reactors: a heavy oil cracking reactor, a light hydrocarbon cracking reactor and an enhanced catalytic conversion reactor;

the method comprises the following steps:

simultaneously, the multi-branched olefin product is injected into a heavy oil cracking reactor or an enhanced catalytic conversion reactor to carry out catalytic conversion reaction to obtain a third reaction mixture;

When the multi-branched olefin product is injected into a heavy oil cracking reactor, the injection point should be located after the heavy feedstock feed injection point; when the multi-branched olefin product is injected into the enhanced catalytic conversion reactor, the injection point is positioned at the lower part of the enhanced catalytic conversion reactor;

In one embodiment, the multi-branched olefin product has a distillation range of from 90 to 150 ℃, preferably from 100 to 130 ℃, and is predominantly trimethylpentene, wherein the amount of trimethylpentene is from 65 to 99% based on the total weight of the multi-branched olefin product. In one embodiment, the high-cracking olefin product has a distillation range of 130 to 330 ℃, preferably 140 to 290 ℃, and is predominantly C9-C12 olefins (C9 olefins having a boiling point of 147 ℃, C12 olefins having a boiling point of 213 ℃, and C13 olefins having a boiling point of 233 ℃), wherein the amount of C9-C12 olefins is 60 to 90% based on the total weight of the high-cracking olefin product.

The method provided by the present disclosure is further described below with reference to fig. 1, but the present disclosure is not limited thereto.

In fig. 1, a heavy oil cracking reactor 1 is a riser reactor, a light hydrocarbon cracking reactor 2 is a riser reactor, and an enhanced catalytic conversion reactor 3 is a fluidized bed reactor. The fluidized bed 3 is located above the stripping zone 7 in the settler 6.

The first catalytic cracking catalyst (hot regenerated catalyst) enters the riser bottom of the heavy oil cracking reactor 1 from the regenerator 9 via a first catalytic cracking catalyst delivery line 11 and is accelerated to flow upward by the pre-lift medium injected via line 51. The preheated heavy raw material is mixed with atomized steam from a pipeline 41 through a pipeline 21 and then injected into a riser of the heavy oil cracking reactor 1, and the weight ratio of the water steam to the hydrocarbon oil raw material is (0.05-1): 1, the outlet temperature of the riser reactor 1 is 480-600 ℃, the reaction time in the riser reactor 1 is 0.5-10 seconds, the weight ratio of the catalyst to the hydrocarbon oil raw material is 5-15, and the absolute pressure in the settler 6 is 0.1-0.40 MPa.

The multi-branched olefin product 35 is optionally injected into the heavy oil cracking reactor 1, and when the multi-branched olefin product 35 is injected into the heavy oil cracking reactor 1, the injection point should be located after the injection point of the heavy raw material feed, and if the total length from the injection point of the heavy raw material feed 21 to the outlet of the heavy oil cracking reactor 1 is 100% and the injection point of the heavy raw material feed 21 is 0%, the multi-branched olefin product 35 is optionally injected into the heavy oil cracking reactor 1 at a position of 50% -100%. The reaction conditions for the multi-branched olefin product 35 are: the reaction temperature is 450-600 ℃, preferably 460-550 ℃; weight hourly space velocity of 1-30 h ^-1 。

The multi-branched olefin product 35 is mainly trimethylpentene, after the first catalytic cracking catalyst (hot regenerated catalyst) is contacted and reacted with the heavy raw material feed 21, the temperature of the first catalytic cracking catalyst is greatly reduced, the multi-branched olefin product 35 is contacted and reacted with the first catalytic cracking catalyst with the greatly reduced temperature, so that the cracking reaction of the multi-branched olefin product 35 can be reduced, the lower reaction temperature is favorable for more hydrogen transfer reaction and aromatization reaction, and meanwhile, the condensation reaction can occur when the heavy raw material feed 21 is contacted and cracked, a large amount of active hydrogen required by the hydrogen transfer reaction can be provided, and the conversion of the trimethylpentene into trimethylpentane can be promoted.

The mixture of the reaction oil gas and the catalyst in the riser 1 can be further introduced into the enhanced catalytic conversion reactor 3 for further reaction through an outlet, the reacted first carbon deposition catalyst is introduced into the stripping zone 7, the separated reaction oil gas (first product) is sent into a subsequent catalytic product separation system 61 through a settler 6 and a pipeline 20 at the top of the settler for product separation, and products such as dry gas, C3 fractions (propylene and propane), C4 liquefied gas fractions (C4 fractions), light gasoline fractions, heavy gasoline fractions, diesel oil fractions, heavy oil fractions and the like are obtained after separation (respectively led out through pipelines 24, 25, 26, 27, 30, 31, 32).

The second catalytic cracking catalyst (hot regenerated catalyst) enters the riser bottom of the light hydrocarbon cracking reactor 2 from the regenerator 9 via a second catalytic cracking catalyst transfer line 12 and is accelerated to flow upward by the pre-lift medium injected via line 52. The lean olefin recycle stream 38 from the second synthesis product separation system 63 is mixed with the atomized steam from line 45 and then first injected into the light hydrocarbon cracking reactor 2, optionally, a portion of the light gasoline fraction 27 is mixed with the atomized steam and then injected into the light hydrocarbon cracking reactor 2, the high cracking olefin product 37 from the second synthesis product separation system 63 is mixed with the atomized steam from line 44 and then finally injected into the light hydrocarbon cracking reactor 2, and the weight ratio of water vapor to hydrocarbon oil feedstock (i.e., lean olefin recycle stream 38 injected into the light hydrocarbon cracking reactor, light gasoline fraction 27, high cracking olefin product 37) in the riser reactor 2 is (0.1-1): 1, the outlet temperature of the riser reactor 2 is 520-750 ℃, the reaction time in the riser reactor 2 is 0.1-3 seconds, and the weight ratio of the catalyst to the hydrocarbon oil raw material (namely the lean olefin recycle stream 38 injected into the light hydrocarbon cracking reactor, the light gasoline fraction 27 and the high cracking olefin product 37) is 6-40. The mixture of the reaction oil gas and the catalyst of the riser reactor 2 is further introduced into a fluidized bed of the enhanced catalytic conversion reactor 3 through a riser outlet for continuous reaction, the reaction temperature of the fluidized bed 3 is 450-750 ℃, and the weight hourly space velocity is 1-30 h ^-1 The reacted oil gas and a part of catalyst to be regenerated of carbon deposit enter a settler 6 to be separated through a fluidized bed reactor 3, the separated second carbon deposit catalyst enters a stripping zone 7, and the reacted oil gas (second product) is sent into a subsequent catalyst through a pipeline 20 at the top of the settler 6The chemical product separation system 61 performs product separation.

The C4 liquefied gas (C4 fraction) 26 is introduced into a multi-branched olefin synthesis reactor 71 to undergo a first synthesis reaction to obtain a first synthesis product oil gas 33, and the first synthesis product oil gas 33 is further introduced into a first synthesis product separation system 62 to be separated into an isobutene-lean C4 stream fraction 34 and a multi-branched olefin product 35.

The isobutene-lean C4 stream fraction 34 and/or the light gasoline fraction 27 are introduced into a high-cracking olefin synthesis reactor 72 to carry out a second synthesis reaction to obtain a second synthesis product oil gas 36, and the second synthesis product oil gas 36 is further introduced into a second synthesis product separation system 63 to be separated into an olefin-lean recycle stream 38 and a high-cracking olefin product 37.

Stripping steam is injected into the stripping zone 7 through a pipeline 47 and is in countercurrent contact with the spent catalyst of the carbon deposit, so that the reaction oil gas carried by the spent catalyst is stripped as clean as possible. Air is injected into the regenerator 9 through a pipeline 90, and the stripped first carbon deposition catalyst and the stripped second carbon deposition catalyst are sent into the regenerator 9 through a to-be-regenerated catalyst conveying pipeline 17, are contacted with the heated air and are regenerated at 600-800 ℃ to obtain a first regenerated catalyst and a second regenerated catalyst, so that the first regenerated catalyst and the second regenerated catalyst are recycled as the first catalytic cracking catalyst and the second catalytic cracking catalyst. Regeneration flue gas is led out through line 91. In fig. 1, 100 is an external heat collector, which is used to remove heat from the regenerator by heat exchange when necessary, and to reduce the regeneration temperature.

Contacting a heavy feedstock in a heavy oil cracking reactor with a first catalytic cracking catalyst to perform a first catalytic cracking reaction to obtain a first cracking reaction mixture comprising the first catalytic cracking catalyst, as described above; and carrying out catalytic reaction on the first cracking reaction mixture in an enhanced catalytic conversion reactor, and carrying out gas-agent separation on the obtained reaction mixture in a settler to obtain a first carbon deposit catalyst and a first product.

According to the present disclosure, a heavy feedstock is contacted with a first catalytic cracking catalyst in a fluidized state in a heavy oil cracking reactor to perform a first catalytic cracking reaction. The operating conditions of the first catalytic cracking reaction may include: the reaction temperature is 480 to 600 ℃, such as 500 to 560 ℃ or 510 to 550 ℃ or 510 to 530 ℃; the reaction time is 0.5-10 seconds; the weight ratio of the catalyst to the oil (namely the weight ratio of the first catalytic cracking catalyst to the heavy raw material) is (5-15): 1, or (6-12): 1, or (8-10): 1, a step of; the weight ratio of water to oil (namely the weight ratio of water vapor to heavy raw materials) is (0.05-1): 1, for example, (0.08 to 0.5): 1, or (0.1 to 0.3): 1. the reaction time refers to the residence time of the oil and gas in the heavy oil cracking reactor.

Introducing the first product, the second product and the third product into a catalytic product separation system for separation, wherein separated products comprising a dry gas fraction, a liquefied gas product, a gasoline product, a diesel oil fraction and a heavy oil fraction can be obtained according to different distillation ranges (boiling point ranges); and the liquefied gas product can be further separated into a C3 liquefied gas fraction (propylene, propane) and a C4 liquefied gas fraction (C4 fraction), and the gasoline product can be further separated into a light gasoline fraction and a heavy gasoline fraction.

Methods for separating a first product from a second product in a catalytic product separation system are known, and for example, fractionation may be performed by means of a fractionating column, a rectifying column, or the like, to obtain various fractions according to a set distillation range: separation products including dry gas fractions, liquefied gas products, gasoline products, diesel fractions, and heavy oil fractions; and the liquefied gas product can be further separated into a C3 liquefied gas fraction (propylene, propane) and a C4 liquefied gas fraction (C4 fraction), and the gasoline product can be further separated into a light gasoline fraction and a heavy gasoline fraction. The catalytic product separation system may include one or more fractionation columns or rectifying columns.

In one embodiment, the light gasoline has a distillation range of 9 to 150 ℃, more preferably 9 to 100 ℃, still more preferably 9 to 60 ℃. In one embodiment, the light gasoline has an olefin content of 30 to 90 wt%, preferably 45 to 90 wt%, based on the total weight of the light gasoline.

In one embodiment, the dry gas fraction is mainly hydrogen, methane, ethylene and ethane, the C3 liquefied gas fraction is propylene and propane, the C4 liquefied gas fraction is C4 fraction, the distillation range of the heavy gasoline fraction is 100-220 ℃, the distillation range of the diesel oil fraction is 200-360 ℃, and the distillation range of the heavy oil fraction is 330-800 ℃.

According to the present disclosure, the multi-branched olefin synthesis reactor, the high-cracking olefin synthesis reactor may be selected from one or a combination of several of a fixed bed, a fixed fluidized bed reactor, and a circulating fluidized bed reactor. The multi-branched olefin synthesis reactor and the high-cracking olefin synthesis reactor can be arranged independently or together according to the requirement so as to save the space of a factory building.

According to the present disclosure, the reaction conditions of the multi-branched olefin synthesis reactor are: the reaction temperature is 300-450 ℃, the pressure is 0.5-2.0 MPa, and the mass airspeed is 1-5 h ^-1 . The reaction conditions of the high-cracking olefin synthesis reactor are as follows: the reaction temperature is 180-280 ℃, the pressure is 0.8-2.0 MPa, and the mass airspeed is 0.9-4.0 h ^-1 。

The first synthesis reaction catalyst includes 1 to 12 mass% of NiO, 45 to 82 mass% of amorphous aluminum silicate, and 10 to 50 mass% of aluminum oxide. The second synthesis reaction catalyst includes 1 to 20 mass% of NiO, 40 to 80 mass% of HZSM-5 zeolite, and 10 to 50 mass% of alumina.

The C4 liquefied gas fraction is mainly C4 alkane and C4 alkene, the lower reaction temperature and the higher reaction pressure in the multi-branched alkene synthesis reactor selectively lead 2 isobutene to carry out the superposition reaction under the action of an acid catalyst to generate a multi-branched alkene product (trimethylpentene), and the obtained multi-branched alkene product trimethylpentene and unreacted C4 are separated. The unreacted C4 (n-butene, trans-2-butene and cis-2-butene) or/and the olefin from the light gasoline are subjected to superposition reaction in a high-cracking olefin synthesis reactor to generate C9, C10, C12, C13 and other hydrocarbons, such as 1-pentene in the light gasoline can be subjected to superposition reaction with the n-butene to generate C9 olefin, and the generated C9 olefin has high selectivity of catalytic cracking into propylene because the generated C9 olefin has only one methyl group or does not contain methyl groups; the C9 olefin is further subjected to a superposition reaction with C4 olefin (n-butene, trans-2-butene and cis-2-butene) to generate C13 olefin with fewer methyl branched chains, so that the severity of catalytic cracking reaction of other C4 olefins is greatly reduced, the selectivity of cracking recycle streams into propylene is greatly improved, and the alkane of C4 alkane and light gasoline is reserved because the superposition reaction is difficult to occur. The process realizes the selective superposition of C4 molecules with different structures and different hydrocarbons of the light gasoline into specific superposition products which are easy to crack into propylene and specific superposition products which are difficult to crack into propylene, and the superposition products which are difficult to crack into propylene can be used for producing high-octane gasoline components.

And introducing the second synthetic product oil gas into a second synthetic product separation system for separation to obtain an olefin-lean circulation stream. The second synthetic product separation system performs oil-gas separation according to the molecular type and the distillation range, and the obtained olefin-lean recycle stream is mainly C4 alkane or alkane-rich light gasoline, and the distillation range is-12-120 ℃, and more preferably-7-60 ℃.

The method of the present disclosure further comprises: contacting the olefin lean recycle stream, the light gasoline fraction, and the high cracking olefin product in a light hydrocarbon cracking reactor with a second catalytic cracking catalyst to perform a second catalytic cracking reaction to obtain a second reaction mixture comprising the second catalytic cracking catalyst; and carrying out catalytic reaction on the second reaction mixture in an enhanced catalytic conversion reactor, and carrying out gas-agent separation on the obtained reaction mixture in the settler to obtain a second carbon deposition catalyst and a second product. The second product may be separated with the first product by entering a catalytic product separation system.

According to the present disclosure, an olefin lean recycle stream, a light gasoline fraction, and a high cracking olefin product are sequentially injected into a light hydrocarbon cracking reactor to contact a second catalytic cracking catalyst in a fluidized state to perform a second catalytic cracking reaction. The operating conditions of the second catalytic cracking reaction may include: the reaction temperature is 520 to 750 ℃, such as 520 to 600 ℃ or 520 to 560 ℃; the reaction time is 0.1 to 3 seconds, for example 0.5 to 3 seconds or 1 to 3 seconds or 1.3 to 3 seconds; the weight ratio of the catalyst to oil (i.e., the weight ratio of the second catalytic cracking catalyst to the total amount of injected olefin-lean recycle stream, light gasoline fraction, high cracking olefin product) is (6-40): 1 is, for example, (7 to 30): 1 or (8-25): 1 or (10-20); the water to oil weight ratio (i.e., the weight ratio of water vapor to the total amount of injected olefin lean recycle stream, light gasoline fraction, high cracking olefin product) is (0.1 to 1): 1 is, for example, (0.08 to 0.5): 1 or (0.1 to 0.3): 1. here, the reaction time refers to the residence time of the oil gas in the light hydrocarbon cracking reactor.

The lean olefin recycle stream (mainly C4 alkane or rich alkane light gasoline) returns to the light hydrocarbon cracking reactor, contacts with the high-temperature regenerated catalyst first, and generates high-severity dehydrogenation and cracking reaction at the reaction temperature of 640-750 ℃, thereby improving the alkane cracking effect. Light gasoline fraction (containing C5, C6 and C7 olefins), wherein the C6 and C7 olefins are relatively easy to crack, each C6 olefin is easy to directly crack from the middle of the molecule to generate 2 propylene in a light hydrocarbon cracking reactor at a relatively mild reaction temperature of 520-600 ℃, each C7 olefin molecule is easy to directly crack from the middle of the molecule to generate 1 propylene and 1 butene molecule, a majority of C5 olefins with smaller molecular number are reserved in the light gasoline, and finally the light gasoline is introduced into a high-cracking olefin synthesis reactor. The higher cracking olefin products (C9, C10, C12 olefin, etc.), the longer the molecular chain and the less the number of branched methyl groups, the greatly reduced reaction severity is needed for the reaction of cracking hydrocarbon into propylene, so that the reaction is carried out in the light hydrocarbon cracking reactor at last, the reaction is carried out at a more moderate reaction temperature of 520-560 ℃, for example, the C12 olefin molecules are easy to directly break into 2C 6 olefin molecules, each C6 olefin molecule is further broken into 2 propylene in the middle, and the propylene selectivity is greatly improved. The whole reaction process is set to ensure that specific molecules (n-butene, trans-2-butene and cis-2-butene) in the C4 hydrocarbon and the light gasoline design the reaction process according to the difference of molecular structures, and the conversion into propylene is high in selectivity and high in conversion rate.

Thus, in one embodiment, in the light hydrocarbon cracking reactor, the injection point of the olefin-lean recycle stream is the forefront and the injection point of the high-cracking olefin product is the last, with the injection point of the light gasoline fraction in between, in terms of the stream direction.

The multi-branched olefin products and high-cracking olefin products may be self-produced from the processes described in the present disclosure, as well as from other units, in accordance with the present disclosure. The olefin content of the multi-branched olefin product may be from 60 to 100 wt%, preferably from 80 to 100 wt%, for example from 80 to 90 wt% or from 85 to 95 wt% or from 90 to 100 wt%, based on the total weight of the multi-branched olefin product. The olefin content of the high cracking olefin product may be from 60 to 100 wt%, preferably from 80 to 100 wt%, for example from 80 to 90 wt% or from 85 to 95 wt% or from 90 to 100 wt%, based on the total weight of the high cracking olefin product.

According to the present disclosure, the weight ratio of the multi-branched olefin product to the heavy feedstock is (0.01 to 0.3): 1, preferably (0.05 to 0.15): 1. the weight ratio of the high-cracking olefin product to the heavy feedstock is (0.01 to 0.4): 1, preferably (0.05 to 0.2): 1. the weight ratio of the olefin-lean recycle stream to the heavy feedstock is (0.01 to 0.3): 1, preferably (0.05 to 0.2): 1. the weight ratio of the light gasoline fraction to the heavy raw material is (0.01-0.3): 1, preferably (0.05 to 0.2): 1.

According to the present disclosure, the method further comprises: separating the olefin-lean recycle stream into a C4 hydrocarbon fraction and an alkane-rich light gasoline fraction, and combining the alkane-rich light gasoline fraction with the heavy gasoline fraction, a portion of the multi-branched olefin product, a gasoline product having a low olefin content may be obtained. In the low olefin gasoline product, the olefin content may be 6 to 16%. According to the present disclosure, the C4 paraffins in the C4 hydrocarbon fraction are 95% or more and the C5 and higher paraffins in the paraffin-rich light gasoline fraction are 85% or more. The separation of the olefin-lean recycle stream into a C4 hydrocarbon fraction and an alkane-rich light gasoline fraction may be carried out using conventional fractionation or rectification methods.

According to the present disclosure, the heavy feedstock is at least one selected from the group consisting of petroleum hydrocarbon oil, synthetic oil, coal liquefied oil, oil sand oil and shale oil, preferably petroleum hydrocarbon oil, which is at least one selected from the group consisting of atmospheric wax oil, vacuum wax oil, coker wax oil, deasphalted oil, hydrogenated tail oil, atmospheric residual oil, vacuum residual oil and crude oil, hydrogenated wax oil, hydrogenated residual oil.

According to the present disclosure, the first catalytic cracking catalyst and the second catalytic cracking catalyst may be the same catalytic cracking catalyst, or may be different catalytic cracking catalysts, preferably the same catalytic cracking catalyst, which are conventionally used in the art for performing catalytic cracking reactions.

The specific types of the first catalytic cracking catalyst and the second catalytic cracking catalyst are not particularly limited in the present disclosure. Preferably, the first and second catalytic cracking catalysts each contain a shape selective zeolite having an average pore diameter of less than 0.7 nm, and the shape selective zeolite may be at least one selected from the group consisting of zeolite having an MFI structure, ferrierite, chabazite, cyclospar, erionite, a-type zeolite, zeolite-column, and turbid zeolite. Wherein, the zeolite with MFI structure can be one or more of ZSM-5 and ZRP series zeolite, and can also be one or more of ZSM-5 and ZRP series zeolite modified by at least one element of RE, P, fe, co, ni, cu, zn, mo, mn, ga and Sn. In an alternative embodiment of the present disclosure, the catalytic cracking catalyst comprises 15 to 50 wt% clay on a dry basis (weight calcined at 800 ℃ for 1 hour), 15 to 50 wt% molecular sieve on a dry basis and 10 to 35 wt% binder on a dry basis, the molecular sieve being a zeolite of MFI structure or being composed of 25 to 100 wt% MFI structure zeolite and other zeolite than 0 to 75 wt% MFI structure zeolite; the MFI structure zeolite is preferably a ZSM-5 molecular sieve and/or HZSM-5 molecular sieve modified with phosphorus and at least one element selected from RE, P, fe, co, ni, cu, zn, mo, mn, ga and Sn. The clay is preferably, for example, one or more selected from the group consisting of kaolin, halloysite, montmorillonite, diatomaceous earth, halloysite, quasi-halloysite, saponite, rectorite, sepiolite, attapulgite, hydrotalcite and bentonite. The binder is, for example, one or more of acidified pseudo-boehmite, alumina sol, silica sol, magnesia alumina sol, zirconia sol, titanium sol, preferably acidified pseudo-boehmite, alumina sol and the like.

In accordance with the present disclosure, the heavy oil cracking reactor and the light hydrocarbon cracking reactor may be catalytic conversion reactors well known to those skilled in the art, for example, the heavy oil cracking reactor and the light hydrocarbon cracking reactor are each one selected from the group consisting of a riser reactor, a downpipe reactor, a fluidized bed reactor, a riser-downpipe composite reactor, a riser-fluidized bed composite reactor, and a downpipe-fluidized bed composite reactor. The fluidized bed reactor may be one selected from the group consisting of a fixed fluidized bed reactor, a bulk fluidized bed reactor, a bubbling bed reactor, a turbulent bed reactor, a fast bed reactor, a transport bed reactor, and a dense-phase fluidized bed. The riser reactor, the downer reactor and the fluidized bed reactor can be equal-diameter riser reactors, downer reactors and fluidized bed reactors, and can also be various variable-diameter riser reactors, downer reactors and fluidized bed reactors.

According to the method, the products of the first catalytic cracking reaction and the second catalytic cracking reaction can continuously perform catalytic reaction in the enhanced catalytic conversion reactor, so that the reaction residence time can be further prolonged, the reaction conversion rate is improved, and propylene is produced more.

The enhanced catalytic conversion reactor is preferably a fluidized bed reactor, which may be one selected from the group consisting of a fixed fluidized bed reactor, a bulk fluidized bed reactor, a bubbling bed reactor, a turbulent bed reactor, a fast bed reactor, a entrained bed reactor, and a dense-phase fluidized bed, according to the present disclosure. The fluidized bed reactor can be of a fluidized bed structure with equal diameter or a variable-diameter fluidized bed structure. The operating conditions of the enhanced catalytic conversion reactor may be: the reaction temperature is 450-750deg.C, such as 480-to-ultra600 ℃, or 500-580 ℃, or 510-560 ℃, or 520-550 ℃, preferably 510-560 ℃; weight hourly space velocity of 1-30 h ^-1 For example 3 to 28 hours ^-1 Or 5 to 25 hours ^-1 Or 6 to 20 hours ^-1 。

According to the present disclosure, the method further comprises:

stripping the first carbon deposit catalyst and the second carbon deposit catalyst in a stripping zone of the settler,

regenerating the stripped first carbon deposition catalyst and the stripped second carbon deposition catalyst in a regenerator to obtain a first regenerated catalyst and a second regenerated catalyst respectively;

and (c) feeding the first regenerated catalyst as the first catalytic cracking catalyst into the heavy oil cracking reactor, and feeding the second regenerated catalyst as the second catalytic cracking catalyst into the light hydrocarbon cracking reactor.

According to the present disclosure, the first carbon deposit catalyst and the second carbon deposit catalyst may be stripped in a stripping zone of the settler. The reaction oil gas carried by the spent catalyst can be stripped as clean as possible by steam stripping. The absolute pressure in the settler may be 0.1 to 0.40MPa.

According to the disclosure, the stripped first carbon deposition catalyst and the stripped second carbon deposition catalyst are introduced into a regenerator to be regenerated, so as to obtain a first regenerated catalyst and a second regenerated catalyst, the first regenerated catalyst is used as the first catalytic cracking catalyst to be sent into a heavy oil cracking reactor, and the second regenerated catalyst is used as the second catalytic cracking catalyst to be sent into a light hydrocarbon cracking reactor. Thereby, the first catalytic cracking catalyst and the second catalytic cracking catalyst can be recycled. In the regenerator, the stripped first and second carbon deposit catalysts are contacted with heated air and regenerated at 600-800 ℃.

According to the present disclosure, the temperature of the first regenerated catalyst is 560 to 800 ℃. In another embodiment, according to the present disclosure, the temperature of the second regenerated catalyst is 560 to 800 ℃.

As shown in fig. 1, the present disclosure also provides a system for high selectivity catalytic cracking of high propylene yield and high gasoline yield comprising:

A reactor system, the reactor system comprising:

a heavy oil cracking reactor 1 provided with a plurality of heavy oil cracking reactor raw material inlets, a heavy oil cracking reactor catalyst inlet and a heavy oil cracking reactor product outlet; the heavy oil cracking reactor feed inlet comprises a heavy feed inlet, and optionally a multi-branched olefin product inlet;

the light hydrocarbon cracking reactor 2 is provided with a plurality of light hydrocarbon cracking reactor raw material inlets, a light hydrocarbon cracking reactor catalyst inlet and a light hydrocarbon cracking reactor product outlet; and

the enhanced catalytic conversion reactor 3 is provided with an enhanced catalytic conversion reactor raw material inlet and an enhanced catalytic conversion reactor product outlet; the material inlet of the enhanced catalytic conversion reactor is communicated with the product outlet of the heavy oil cracking reactor, and the product outlet of the light hydrocarbon cracking reactor is communicated with the middle lower part of the enhanced catalytic conversion reactor, so that material flows from the heavy oil cracking reactor and the light hydrocarbon cracking reactor enter the enhanced catalytic conversion reactor;

a catalytic product separation system 61 provided with a feedstock inlet and a plurality of fraction outlets; the raw material inlet of the catalytic product separation system is communicated with the product outlet of the enhanced catalytic conversion reactor; the plurality of fraction outlets of the catalytic product separation system comprise a dry gas fraction outlet, a C3 liquefied gas fraction outlet, a C4 liquefied gas fraction outlet, a light gasoline fraction outlet, a heavy gasoline fraction outlet, a diesel fraction outlet and a heavy oil fraction outlet;

A multi-branched olefin synthesis reactor 71 provided with a multi-branched olefin synthesis reactor feed inlet and a multi-branched olefin synthesis reactor product outlet, the multi-branched olefin synthesis reactor feed inlet being in communication with the C4 liquefied gas fraction outlet of the catalytic product separation system;

a first synthesis product separation system 62, the first synthesis product separation system having a first synthesis product separation system feed inlet and a plurality of first synthesis product separation system separation product outlets; the raw material inlet of the first synthesis product separation system is communicated with the product outlet of the multi-branched olefin synthesis reactor; the plurality of first synthesis product separation system separation product outlets comprises an isobutylene-lean C4 stream outlet and a multi-branched olefin product outlet;

a high-split olefin synthesis reactor 72 provided with a high-split olefin synthesis reactor feed inlet and a high-split olefin synthesis reactor product outlet; the high-cracking olefin synthesis reactor raw material inlet is communicated with the isobutene-lean C4 stream outlet;

a second synthesis product separation system 63, the second synthesis product separation system having a second synthesis product separation system feed inlet and a plurality of second synthesis product separation system separation product outlets; the raw material inlet of the second synthesis product separation system is communicated with the product outlet of the high-cracking olefin synthesis reactor; the separation product outlets of the second synthetic product separation systems comprise an olefin-lean circulation flow outlet and a high-cracking olefin product outlet, and the olefin-lean circulation flow outlet and the high-cracking olefin product outlet are respectively communicated with the raw material inlet of one light hydrocarbon cracking reactor;

In one embodiment, the total length from the heavy feedstock inlet of the heavy oil cracking reactor to the heavy oil cracking reactor outlet is 100%, the heavy feedstock inlet position is 0%, and the multi-branched olefin product inlet is located at 50% -100% of the heavy oil cracking reactor.

In one embodiment, the heavy oil cracking reactor product outlet and the light hydrocarbon cracking reactor product outlet are located inside the enhanced catalytic conversion reactor such that the streams from the heavy oil cracking reactor and the light hydrocarbon cracking reactor (the reaction mixture comprising product and catalyst) may directly enter the enhanced catalytic conversion reactor for continued reaction. In this disclosure, the term "in communication" means that two in communication may be connected by a pipeline.

Embodiments described in the method of the present disclosure are also applicable to the system of the present disclosure, and are not described here again.

The methods provided by the present disclosure are further illustrated by the following examples, but the present disclosure is not so limited.

The first catalytic cracking catalyst and the second catalytic cracking catalyst used in the following examples and comparative examples were cracking catalysts sold under the trade name OMT-plus, manufactured by zilutong, a petrochemical catalyst, and specific properties thereof are shown in table 1-1, and the catalysts contain shape selective zeolite having an average pore diameter of less than 0.7 nm.

Example 1

Example l illustrates the effect of propylene and gasoline yield increase during the high selectivity catalytic cracking process of the present disclosure for propylene and gasoline yield increase.

Experiments were performed using a medium-sized apparatus with a continuous reaction-regeneration operation of three reactors, in which the heavy oil cracking reactor was a riser with an inner diameter of 16 mm and a height of 3800 mm. The light hydrocarbon cracking reactor is a riser reactor with an inner diameter of 16 mm and a height of 3200 mm. The first riser reactor and the second riser reactor are introduced into the fluidized bed reactor as intensified catalytic conversion reactor, the inner diameter of the fluidized bed reactor is 64 mm, and the height of the fluidized bed reactor is 300 mm.

The first catalytic cracking catalyst is a first regenerated catalyst with the temperature of 680 ℃, enters the bottom of a riser reactor of the heavy oil cracking reactor through a first catalytic cracking catalyst inclined tube, and flows upwards under the action of pre-lifting steam. The heavy raw material is tail oil obtained by a residual oil hydrogenation device, namely hydrogenated residual oil (the main properties are shown in table 2), after being heated to 350 ℃ by a preheating furnace, the hydrogenated residual oil is mixed with atomized water vapor, and then is sprayed into a heavy oil cracking reactor through a feeding nozzle to contact with a hot first regenerated catalyst for catalytic conversion reaction, and high-cracking olefin products (100-130 ℃) are injected from the position (0 percent in the residual oil nozzle) to the middle 50 percent (100 percent in the riser reactor outlet position) of the heavy oil cracking reactor and contact with the first regenerated catalyst with reduced temperature for catalytic conversion reaction. After the reaction, the oil gas (first product and third product) and the first carbon deposit catalyst enter a fluidized bed of the enhanced catalytic conversion reactor from a riser outlet of the heavy oil cracking reactor to continue the reaction, then enter a settler to perform rapid separation, and the first carbon deposit catalyst enters a stripping zone to perform stripping. The second catalytic cracking catalyst is a second regenerated catalyst with the temperature of 680 ℃, enters the bottom of a riser reactor of the light hydrocarbon cracking reactor through a second catalytic cracking catalyst inclined tube, and flows upwards under the action of pre-lifting steam. The olefin-lean circulation material flow (-12-120 ℃), the light gasoline (9-120 ℃) and the high-cracking olefin product (140-250 ℃) sequentially enter a light hydrocarbon cracking reactor to contact with a hot second regenerated catalyst for catalytic reaction. The weight ratio of olefin lean recycle stream to heavy feedstock was 0.1:1, the weight ratio of the light gasoline to the heavy raw materials is 0.03:1, the weight ratio of the multi-branched olefin product to the heavy feedstock is 0.10:1, the weight ratio of highly cracking olefin product to heavy feedstock is 0.25:1. the reaction oil mixture from the light hydrocarbon cracking reactor is introduced into the fluidized bed of the enhanced catalytic conversion reactor at the outlet of the riser for further reaction, and the reacted oil gas (second product) and the second carbon deposit catalyst enter a settler for separation through the fluidized bed reactor, and the separated second carbon deposit catalyst enters a stripping zone for stripping under the action of gravity. The first carbon deposition catalyst and the second carbon deposition catalyst after steam stripping enter a regenerator through a to-be-regenerated agent conveying pipe, contact with heated air and are regenerated at 700 ℃, and the obtained hot first regenerated catalyst and the obtained hot second regenerated catalyst are respectively returned to a heavy oil cracking reactor and a light hydrocarbon cracking reactor for recycling. The reacted oil gas (first product and second product) is led out of the settler together, and is led into a catalytic product separation system for product separation to obtain gas products and various liquid products, and meanwhile, the gas products and various liquid products are partially separated to obtain C4 fraction, light gasoline and heavy gasoline.

Introducing the C4 fraction into a multi-branched olefin synthesis reactor to perform superposition reaction to obtain oil gas of a first synthesis product, and further introducing the oil gas into a first synthesis product separation system to separate, thereby obtaining an isobutene-poor C4 material flow and a multi-branched olefin product (100-130 ℃); introducing the isobutene-lean C4 material flow and the light gasoline fraction into a high-cracking olefin synthesis reactor for reaction to obtain second synthesis product oil gas, and further introducing the second synthesis product oil gas into a second synthesis product separation system for separation to obtain olefin-lean recycle material flow (-12-120 ℃) and high-cracking olefin products (140-250 ℃).

The first synthetic reaction catalyst is LXC-10 catalyst produced by petrochemical industry scientific institute of China petrochemical industry, and the second synthetic reaction catalyst is SXC-6 catalyst produced by petrochemical industry scientific institute of China petrochemical industry. The superposition reactor is a fixed bed reactor; the first synthesis reaction conditions were: the reaction temperature is 350 ℃, the reaction pressure is 1.1Mpa, and the weight hourly space velocity is 3.0h ^-1 The method comprises the steps of carrying out a first treatment on the surface of the The second synthesis reaction conditions were: the reaction temperature is 270 ℃, the reaction pressure is 1.5Mpa, and the weight hourly space velocity is 2.0h ^-1 . The preparation of the SXC-6 catalyst and the LXC-10 catalyst is shown later.

The olefin-lean recycle stream may further separate the C4 hydrocarbons and the paraffinic light gasoline, some of which may be blended into a low olefin gasoline product.

The main operating conditions and results are shown in Table 3.

The preparation process of the SXC-6 catalyst is as follows:

mixing 13.4g of HZSM-5 zeolite with 8.5g of alumina powder, uniformly mixing 15g of deionized water, adding a proper amount of dilute nitric acid solution, kneading and extruding into strips with the diameter of 1.5 mm, airing at room temperature, drying at 120 ℃ for 4 hours, roasting at 540 ℃ for 3 hours, crushing and sieving to obtain particles with the diameter of 0.6-0.9 mm, thereby obtaining HZSM-5-Al ₂ O ₃ And (3) a composite carrier. SiO of HZSM-5 zeolite used ₂ /Al ₂ O ₃ The molar ratio was 200.

The obtained 11g of composite carrier was treated with 5.7g of Ni (NO) ₃ ) ₂ ﹒6H ₂ Soaking the solution prepared by O for 6h according to conventional method, filtering, oven drying at 100deg.C, and adding N ₂ Activating at 450 ℃ for 6 hours in the atmosphere to obtain the catalyst of the invention, wherein the specific surface is 320m ² Per gram, pore volume was 0.26ml/g.

The LXC-10 catalyst was prepared as follows:

mixing 12.3g of amorphous aluminum silicate powder with 7.3g of aluminum oxide powder, uniformly mixing 15g of deionized water, adding a proper amount of dilute nitric acid solution, kneading and extruding into strips with the diameter of 1.5 mm, airing at room temperature, drying at 120 ℃ for 4 hours, roasting at 540 ℃ for 3 hours, crushing and sieving to obtain amorphous aluminum silicate-Al particles with the diameter of 0.6-0.9 mm ₂ O ₃ And (3) a composite carrier. SiO of the amorphous aluminum silicate powder used ₂ /Al ₂ O ₃ The molar ratio was 10.

10g of the obtained composite carrier was treated with 2.2g of Ni (NO ₃ ) ₂ ﹒6H ₂ Soaking the solution prepared by O for 6h according to conventional method, filtering, oven drying at 100deg.C, and adding N ₂ Activating at 450 ℃ for 6 hours in the atmosphere to obtain the catalyst of the invention, wherein the specific surface is 310m ² Per gram, pore volume was 0.31ml/g.

The specific chemical composition properties of the SXC-6 catalyst and the LXC-10 catalyst are shown in tables 1-2.

Comparative example l-1

Comparative examples 1-l demonstrate the effect of cycle cracking whole fraction light gasoline and C4 fraction to increase propylene and gasoline production during a process for high selectivity catalytic cracking to increase propylene and gasoline production.

The first catalytic cracking catalyst is a first regenerated catalyst with the temperature of 680 ℃, enters the bottom of a riser reactor of the heavy oil cracking reactor through a first catalytic cracking catalyst inclined tube, and flows upwards under the action of pre-lifting steam. The heavy raw material is tail oil (main properties are shown in table 2) obtained by a residual oil hydrogenation device, the tail oil is heated to 350 ℃ by a preheating furnace, then is mixed with atomized water vapor, is sprayed into a heavy oil cracking reactor through a feeding nozzle, and is contacted with a hot first regenerated catalyst to carry out catalytic conversion reaction. The reaction oil gas (first product) and the first carbon deposition catalyst enter a fluidized bed of an enhanced catalytic conversion reactor from a riser outlet of a heavy oil cracking reactor, then further enter a settler for rapid separation, and the first carbon deposition catalyst enters a stripping zone for stripping. The second catalytic cracking catalyst is a second regenerated catalyst with the temperature of 680 ℃, enters the bottom of a riser reactor of the light hydrocarbon cracking reactor through a second catalytic cracking catalyst inclined tube, and flows upwards under the action of pre-lifting steam. The C4 fraction and the light gasoline (9-120 ℃) are sequentially fed into a light hydrocarbon cracking reactor to be contacted with a hot second regenerated catalyst for catalytic reaction. The weight ratio of C4 fraction to heavy feedstock was 0.20:1, the weight ratio of the light gasoline to the heavy raw materials is 0.20:1. the reaction oil mixture from the light hydrocarbon cracking reactor is introduced into the fluidized bed of the enhanced catalytic conversion reactor at the outlet of the riser for further reaction, and the reacted oil gas (second product) and the second carbon deposit catalyst enter a settler for separation through the fluidized bed reactor, and the separated second carbon deposit catalyst enters a stripping zone for stripping under the action of gravity. The first carbon deposition catalyst and the second carbon deposition catalyst after steam stripping enter a regenerator through a to-be-regenerated agent conveying pipe, contact with heated air and are regenerated at 700 ℃, and the obtained hot first regenerated catalyst and the obtained hot second regenerated catalyst are respectively returned to a heavy oil cracking reactor and a light hydrocarbon cracking reactor for recycling. The reaction oil gas from the heavy oil cracking reactor and the enhanced catalytic conversion reactor is led out of the settler together, and is led into a product separation system for product separation to obtain gas products and various liquid products, and meanwhile, the gas products and various liquid products are partially separated to obtain C4 fraction, light gasoline and heavy gasoline.

Part of the light gasoline and the heavy gasoline are mixed into a low-olefin gasoline product.

Among them, the reaction conditions of the light hydrocarbon cracking reactor and the enhanced catalytic conversion reactor were more severe than those in example 1 (the temperatures of the light hydrocarbon cracking reactor and the enhanced catalytic conversion reactor were higher than those in example 1), and the main operation conditions and results are shown in Table 3.

Comparative examples 1 to 2

Comparative examples 1-2 demonstrate the effect of cycle pyrolysis of light gasoline and C4 fractions to yield propylene and gasoline during the process of high selectivity catalytic pyrolysis to yield more propylene and to yield gasoline.

The reaction apparatus used was the same as that of comparative example 1-1. The raw materials and main experimental steps are the same as those of comparative example 1-1, except that the reaction severity of the light hydrocarbon cracking reactor and the enhanced catalytic conversion reactor is greatly reduced, namely, the same reaction conditions as those of example 1 are adopted.

The main operating conditions and results are shown in Table 3.

Example 2

Example 2 illustrates the effect of propylene and gasoline yield increase during the high selectivity catalytic cracking process of the present disclosure for propylene and gasoline yield increase.

The reaction apparatus used was the same as in example l. The main experimental procedure is the same as in example l. The weight ratio of olefin-lean recycle stream (-12-5 ℃) to heavy feedstock was 0.1:1. the weight ratio of the light gasoline (9-120 ℃) to the heavy raw material is 0.3:1. the weight ratio of the multi-branched olefin product (100-130 ℃) to the heavy raw material is 0.11:1. the weight ratio of the high cracking olefin product (170-300 ℃) to the heavy raw material is 0.09:1.

In the example 2, only the fraction of the C4 material flow lean in isobutene is introduced into a high-cracking olefin synthesis reactor to carry out superposition reaction to obtain oil gas of a second synthesis product, and the oil gas is further introduced into a second synthesis product separation system to be separated into a circulating material flow lean in olefin (the distillation range is-12-5 ℃), and a high-cracking olefin product (the distillation range is 170-300 ℃). The multi-branched olefin synthesis reactor, the catalyst for the high-cracking olefin synthesis reactor and the reaction conditions were the same as in example 1.

And mixing part of the light gasoline and the heavy gasoline obtained by catalytic cracking into a gasoline product with low olefin content.

The main operating conditions and results are shown in Table 3.

Example 3

Example 3 illustrates the effect of propylene and gasoline yield increase during the high selectivity catalytic cracking process of the present disclosure for propylene and gasoline yield increase.

The reaction apparatus used was the same as in example l. The main experimental procedure is the same as in example l, except that the heavy feedstock selected is a mixed feedstock of hydrogenated wax oil blended with four-line heavy oil. The main operating conditions and results are shown in Table 3.

TABLE 1-1

Catalyst name	OMT-plus
		Chemical nature, weight percent
Al ₂ O ₃	53.5
		P ₂ O ₅	2.35
RE ₂ O ₃	0.81
		Physical Properties
Total pore volume, ml/g	0.20
		Micropore volume, ml/g	0.02
Specific surface, m ² /g	139
		Micropore area, m ² /g	109
Specific surface, m of matrix ² /g	38
		Bulk density, g/ml	0.74
Particle size distribution, wt%
		0～20μm	1.8
0～40μm	13.5
		0～80μm	59.2
0～110μm	78.2
		0～149μm	93.1
Cleavage Activity, wt%	65

TABLE 1-2

Catalyst name	SXC-6	LXC-10
			Chemical nature, weight percent
Al ₂ O ₃	30.4	36.2
			NiO	14.2	2.8
HZSM-5	55.4
			Amorphous aluminum silicate		61.0

TABLE 2

TABLE 3 Table 3

Examples	Example 1	Comparative examples 1 to 1	Comparative examples 1 to 2	Example 2	Example 3
						Intermediate product Properties
Olefin lean recycle stream, DEG C	-12～120			-12～5	-12～120
						Light gasoline distillation range, DEG C	9-120	9-120	9-120	9-120	9-120
The distillation range of multi-branched olefin products, DEG C	100-130			100-130	100-130
						High cracking olefin product distillation range, DEG C	140～250			170-300	140-250
Feed-back ratio
						Weight ratio of olefin lean recycle stream to heavy hydrocarbon oil	0.10			0.10	0.07
Weight ratio of light gasoline to heavy hydrocarbon oil	0.03	0.20	0.20	0.30	0.05
						Weight ratio of C4 to heavy hydrocarbon oil		0.20	0.20
Weight ratio of multi-branched olefin product to heavy hydrocarbon oil	0.10			0.11	0.11
						Weight ratio of highly cracking olefin product to heavy hydrocarbon oil	0.25			0.09	0.24
Reaction conditions of the heavy oil cracking reactor:
						raising the outlet temperature of the pipe at DEG C	530.0	540.0	540.0	500.0	530
Total reaction time of riser, sec	2.0	2.2	2.2	2.0	2.1
						Catalyst to raw material weight ratio	8.6	9.0	9.0	6.8	8.3
Atomized water vapor proportion, weight percent	15.0	15.0	15.0	15.0	25
						Temperature of the first catalytic cracking catalyst, DEG C	680	680	680	680	680
Reaction conditions of the light hydrocarbon cracking reactor:
						raising the outlet temperature of the pipe at DEG C	560.0	620.0	560.0	620.0	560
Total reaction time of riser, sec	1.48	1.46	1.46	1.46	1.41
						Catalyst to raw material weight ratio	10.0	10.0	10.0	10.0	10
Atomized water vapor proportion, weight percent	25.0	25.0	25.0	25.0	25
						Second catalytic cracking catalyst temperature, DEG C	680	680	680	680	680
Reaction conditions of the enhanced catalytic conversion reactor:
						reaction temperature, DEG C	520.0	615.0	555.0	480.0	520
Fluidized bed mass weight hourly space velocity, h ^-1	10.0	10.0	10.0	8.0	10
						Settler pressure, MPa (absolute pressure)	0.21	0.21	0.21	0.21	0.21
Balance of materials, weight percent
						Dry gas	4.99	8.09	5.51	4.83	5.92
Liquefied gas	31.15	34.36	34.71	32.55	35.66
						C5 gasoline (C5-221 ℃, TBP)	41.65	34.08	37.37	40.42	38.16
Diesel (221-330 ℃, TBP)	11.41	11.55	11.45	11.40	10.85
						Heavy oil>330℃,TBP)	3.47	3.46	3.44	3.44	2.68
Coke	7.33	8.46	7.52	7.36	6.73
						Totals to	100.00	100.00	100.00	100.00	100.00
Propylene yield, wt%	24.06	20.18	16.74	25.40	27.83
						C5 gasoline olefin content, wt%	8.79	18.12	26.35	14.94	6.05
Octane number RON of C5 gasoline	97.3	96.4	95.1	97.2	97.1

As can be seen from Table 3, example 1, using the method provided by the present disclosure, produced propylene in high yield by high selectivity catalytic cracking with heavy feedstock as feed and increased gasoline yield of 24.06 wt%, gasoline yield of 41.65 wt%, is significantly higher than comparative examples 1-1 and 1-2, and the product gasoline has an olefin content of only 8.79 wt%, and RON as high as 97.3; in comparative examples 1-2, the reaction severity was higher with a 16.74 wt% propylene yield by the C4 cut and light gasoline recycle catalytic conversion method, and the gasoline yield was 37.37 wt%.

In comparative example 1-1, the yield of propylene produced by the C4 cut and light gasoline cycle catalytic conversion method was 20.18 wt% with heavy raw materials as the feed, but after the reaction severity was improved, the dry gas and coke yields were greatly improved. Example 1 has 3.10 wt% and 1.13 wt% lower dry gas and coke yields, respectively, 3.88 wt% higher propylene yield and 7.57 wt% higher gasoline yield, as compared to comparative example 1-1.

Example 2 using the process provided by the present disclosure, propylene yields of 25.40 wt.% and gasoline yields of 40.42 wt.% were produced by high selectivity catalytic cracking of high propylene yields and increased gasoline, with the gasoline RON being as high as 97.2, fed with heavy feedstock.

Example 3 using the process provided by the present disclosure, propylene was produced in 27.83 wt.% yield, 38.16 wt.% gasoline yield, gasoline RON as high as 97.1, and product gasoline olefins as low as 6.05 wt.% by high selectivity catalytic cracking of the heavy feedstock to produce more propylene and increased gasoline.

It will be appreciated by persons skilled in the art that the embodiments described herein are merely exemplary and that various other alternatives, modifications and improvements may be made within the scope of the invention. Thus, the present invention is not limited to the above-described embodiments, but only by the claims.

Claims

1. A method for producing more propylene and increasing the yield of gasoline by high-selectivity catalytic pyrolysis is characterized in that,

The method comprises the following steps:

injecting the multi-branched olefin product into the upper part of a heavy oil cracking reactor or an enhanced catalytic conversion reactor, and performing catalytic conversion reaction to obtain a third reaction mixture;

performing gas-agent separation on the obtained first reaction mixture, second reaction mixture and third reaction mixture in the settler to obtain a first carbon deposit catalyst, a second carbon deposit catalyst, a first product, a second product and a third product;

2. The method of claim 1, wherein,

the reaction conditions of the first synthesis reaction are as follows: the reaction temperature is 300-450 ℃, the pressure is 0.5-2.0 MPa, and the mass airspeed is 1-5 h ^-1 ；

The reaction conditions of the second synthesis reaction are as follows: the reaction temperature is 180-280 ℃, the pressure is 0.8-2.0 MPa, and the mass airspeed is 0.9-4.0 h ^-1 。

3. The method of claim 1, wherein,

the first synthesis reaction catalyst comprises 1 to 12 mass% of NiO, 45 to 82 mass% of amorphous aluminum silicate and 10 to 50 mass% of aluminum oxide;

the second synthesis reaction catalyst includes 1 to 20 mass% of NiO, 40 to 80 mass% of HZSM-5 zeolite, and 10 to 50 mass% of alumina.

4. The process of claim 1 wherein the multi-branched olefin product has a distillation range of 90 to 150 ℃, primarily trimethylpentene; the high-cracking olefin product has a distillation range of 130-330 ℃ and is mainly C9-C12 olefin.

5. The process of claim 4, wherein the multi-branched olefin product has a distillation range of 100 to 130 ℃; the distillation range of the high-cracking olefin product is 140-290 ℃.

6. The method of claim 1, wherein the light gasoline fraction has a distillation range of 9-150 ℃;

And the total weight of the light gasoline fraction is taken as a reference, and the olefin content of the light gasoline fraction is 20-90 wt%.

7. The method of claim 6, wherein the light gasoline fraction has a distillation range of 9-100 ℃;

the total weight of the light gasoline fraction is taken as a reference, and the olefin content of the light gasoline fraction is 35-90 wt%.

8. The method according to claim 7, wherein the light gasoline fraction has a distillation range of 9-60 ℃.

9. The method of claim 1, wherein,

the weight ratio of the multi-branched olefin product to the heavy raw material is 0.01-0.3: 1, a step of;

the weight ratio of the high-cracking olefin product to the heavy raw material is 0.01-0.4: 1.

10. the method of claim 9, wherein,

the weight ratio of the multi-branched olefin product to the heavy raw material is 0.05-0.15: 1, a step of;

the weight ratio of the high-cracking olefin product to the heavy raw material is 0.05-0.2: 1.

11. the method of any of claims 1-10, wherein the heavy feedstock is at least one selected from the group consisting of petroleum hydrocarbon oil, synthetic oil, coal liquefaction oil, oil sand oil, and shale oil.

12. The method of claim 11, wherein the heavy feedstock is a petroleum hydrocarbon oil that is at least one selected from the group consisting of atmospheric wax oil, vacuum wax oil, coker wax oil, deasphalted oil, hydrogenated tail oil, atmospheric residuum, vacuum residuum, and crude oil, hydrogenated wax oil, hydrogenated residuum.

13. The method of any of claims 1-10, wherein the operating conditions of the first catalytic cracking reaction comprise: the reaction temperature is 480-600 ℃; the reaction time is 0.5-10 seconds; the weight ratio of the agent to the oil is 5-15: 1, a step of; the weight ratio of water to oil is 0.05-1: 1.

14. the method of any of claims 1-10, wherein the operating conditions of the second catalytic cracking reaction comprise: the reaction temperature is 520-750 ℃; the reaction time is 0.1-3 seconds; the weight ratio of the agent to the oil is 6-40: 1, a step of; the weight ratio of water to oil is 0.1-1: 1.

15. the process of any one of claims 1-10, wherein the enhanced catalytic conversion reactor is a fluidized bed reactor, the operating conditions of the enhanced catalytic conversion reactor being: the reaction temperature is 450-750 ℃; the weight hourly space velocity is 1-30 h ^-1 。

16. The method of claim 15, wherein the reaction temperature is 510-560 ℃.

17. The process as claimed in claim 1 or 9, wherein when the multi-branched olefin product is injected into the heavy oil cracking reactor, the injection point should be located after the injection point of the heavy feedstock feed, with the total length from the injection point of the heavy feedstock feed to the outlet of the heavy oil cracking reactor being 100%, the injection point of the heavy feedstock feed being 0%, and the multi-branched olefin product being injected into the heavy oil cracking reactor at a position of 50% -100%;

when the multi-branched olefin product is injected into a heavy oil cracking reactor or an enhanced catalytic conversion reactor, the operating conditions are: the reaction temperature is 450-600 ℃; the weight hourly space velocity is 1-30 h ^-1 。

18. The method of claim 17, wherein when the multi-branched olefin product is injected into a heavy oil cracking reactor or an enhanced catalytic conversion reactor, the operating conditions are: the reaction temperature is 460-550 ℃.

19. The process of any one of claims 1-10, wherein the first and second catalytic cracking catalysts each comprise a shape selective zeolite having an average pore size of less than 0.7 nanometers, the shape selective zeolite being at least one selected from the group consisting of zeolites having MFI structure, ferrierite, chabazite, cyclospar, erionite, a-type zeolite, column zeolite, and turbid zeolite.

20. The method of any of claims 1-10, wherein the heavy oil cracking reactor and the light hydrocarbon cracking reactor are each one selected from a riser reactor, a downer reactor, a fluidized bed reactor, a riser-downer combined reactor, a riser-fluidized bed combined reactor, a downer-fluidized bed combined reactor.

21. The method of any of claims 1-10, further comprising:

and sending the first regenerated catalyst into the heavy oil cracking reactor to serve as the first catalytic cracking catalyst, and sending the second regenerated catalyst into the light hydrocarbon cracking reactor to serve as the second catalytic cracking catalyst.

22. The method of any of claims 1-10, further comprising:

separating the olefin-lean recycle stream into a C4 hydrocarbon fraction and an alkane-rich light gasoline fraction,

combining the paraffinic light gasoline-rich fraction with the heavy gasoline fraction, a portion of the light gasoline fraction, and a portion of the high-cracking olefinic product may result in a gasoline product having a low olefinic content.

23. A system for high selectivity catalytic cracking of high yields of propylene and increased yields of gasoline comprising:

a reactor system, the reactor system comprising: