CN113736510B

CN113736510B - Method and system for high-selectivity catalytic cracking and high-yield propylene production

Info

Publication number: CN113736510B
Application number: CN202010473050.7A
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: 2023-01-24
Anticipated expiration: 2040-05-29
Also published as: CN113736510A

Abstract

The present disclosure relates to a method and a system for high selectivity catalytic cracking and high propylene yield. The method comprises the following steps: carrying out a first catalytic reaction on a heavy hydrocarbon oil raw material in a first catalytic reactor, then carrying out a catalytic reaction in a third catalytic reactor, and separating to obtain a first product; separating the first product from the second product to obtain a liquefied gas fraction after C3 removal and a light gasoline fraction; introducing the liquefied gas fraction subjected to C3 removal into a first superposition reactor for reaction, and separating to obtain unreacted C4 and a first superposition product; introducing the first superposed product into a second superposed reactor for reaction to obtain a second superposed product; introducing the unreacted C4 and the light gasoline fraction into a third superposed reactor for reaction, and separating to obtain an unreacted recycle stream and a third superposed product; and carrying out a second catalytic reaction on the unreacted circulating material flow, the third superposed product and the second superposed product in a second catalytic reactor, then carrying out a catalytic reaction in a third catalytic reactor, and separating to obtain the second product. The method can greatly reduce the reaction severity and greatly improve the selectivity of cracking the recycle stream into propylene.

Description

Method and system for high-selectivity catalytic cracking and high-yield propylene production

Technical Field

The present disclosure relates to a method and a system for high selectivity catalytic cracking and high propylene yield.

Background

In the national vehicle gasoline standard GB17930-2016, the technical requirements on vehicle gasoline are 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 development and reform committee and the like, the vehicle ethanol gasoline is popularized and used in 2020 nationwide, and the vehicle gasoline market is expected to be saturated according to the current popularization of E10 ethanol gasoline (containing 10% ethanol). At the same time, MTBE is limited to the motor gasoline and a new outlet needs to be found for C4 olefins as a feedstock for 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 catalytic cracking of C4 olefin can produce ethylene and propylene in high yield. For example, in journal of petroleum refining and chemical engineering, volume 36, 2, 2005, research on the preparation of propylene and ethylene by catalytic cracking of butene, which reports that ethylene and propylene can be produced by catalytic cracking of C4 olefins, is reported.

There are also patents disclosing the reaction process and catalyst for producing ethylene and propylene by using C4 or liquefied gas as raw material through polymerization and catalytic cracking reaction.

Patent CN102531824B discloses a process for preparing propylene and ethylene from liquefied gas containing butylene: (1) The liquefied gas exchanges heat with the reacted oil-gas mixture and/or is directly heated to reach 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 are subjected to a polymerization reaction to generate macromolecular hydrocarbon under the action of a polymerization catalyst, and the temperature of oil gas is increased due to the reaction heat release; (3) The oil-gas mixture generated by the superposition reaction enters the lower section of the reactor, the cracking reaction is carried out under the action of a cracking catalyst to generate a hydrocarbon mixture containing target products of propylene and ethylene, the hydrocarbon mixture flows out from the bottom of the reactor and enters a subsequent separation system, and the propylene, the ethylene and the aromatic oil generated by the reaction are recovered. The invention adopts two or more catalyst bed layers in the process technology for preparing propylene and ethylene by liquefied gas, so that the raw material liquefied gas is contacted with two or more catalysts in sequence, and the superposition reaction and the cracking reaction occur in sequence, thereby greatly improving the conversion rate of the raw material liquefied gas and the selectivity of propylene and ethylene. However, the upper section and the lower section of the reactor arranged in the technology are arranged in series, the reaction conditions of the upper section and the lower section can interfere with each other, the superposition reaction requires high pressure and low temperature, and the cracking reaction requires low pressure and high temperature. Meanwhile, when the liquefied gas containing C4 passes through the upper section of the reactor once, the full conversion is difficult to realize during the superposition reaction.

The patent CN100537721C discloses a catalytic conversion method for increasing the yield of propylene, the method injects preheated raw oil into a main riser of a double-riser reaction regeneration system, the raw oil is contacted with a hot catalyst to carry out catalytic cracking reaction, reaction products are separated, and a spent catalyst for coke formation is recycled after steam stripping and regeneration; injecting the liquefied gas product from which propylene is separated by the gas separation system into an auxiliary riser, contacting with a hot catalyst, sequentially carrying out superposition reaction, catalytic cracking and alkane dehydrogenation reaction in two reaction zones in the auxiliary riser, separating reaction products, and recycling the regenerated catalyst. The method provided by the invention is adopted to further convert the liquefied gas product after the propylene removal into propylene, and the propylene yield is obviously improved on the premise of not increasing the yield of the liquefied gas. The auxiliary riser is sequentially provided with a first reaction zone, a second reaction zone, an outlet zone and a horizontal pipe which are coaxial with each other 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 were: 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) in the auxiliary riser reactor is 0.1-0.4 MPa.

Similarly, patent CN100448954C discloses a catalytic conversion method for increasing propylene yield, which comprises injecting preheated raw oil into a main riser of a double-riser reaction regeneration system, contacting with a hot catalyst to perform catalytic cracking reaction, separating reaction products, and recycling the spent catalyst after regeneration; injecting the liquefied gas product with the separated propylene into an auxiliary lifting pipe, contacting with a hot catalyst, sequentially carrying out olefin polymerization, polymerization product cracking and alkane dehydrogenation reaction, separating reaction products, and recycling the regenerated catalyst; the catalyst is a mixture of two catalysts: a first cracking catalyst containing a Y-type molecular sieve and a second cracking catalyst containing a ZSM-5 molecular sieve, a transition metal additive and a phosphorus additive, wherein the dry basis weight ratio of the first cracking catalyst to the second cracking catalyst is 10-70: 30-90. The method provided by the invention is adopted to further convert the liquefied gas product after the propylene removal into propylene, and the propylene yield is obviously improved on the premise of not increasing the yield of the liquefied gas. The technology is characterized in that olefin polymerization, polymerization product cracking and alkane dehydrogenation reaction are simultaneously completed in the auxiliary riser. The auxiliary riser is also respectively provided with a first reaction zone (150-450 ℃) for promoting the polymerization reaction of light olefin and a second reaction zone (450-650 ℃) for promoting the further cracking of the polymerization product and the dehydrogenation of propane to generate propylene.

These documents mainly focus on how to carry out a polymerization reaction of liquefied gas or butene containing olefins to produce gasoline fractions or diesel oil fractions, and report less on how to further catalytically crack the polymerization products of liquefied gas or C4 olefins to produce ethylene and propylene. Moreover, the C4 olefin polymerization product is mainly C8 olefin, and the propylene yield and propylene selectivity of direct catalytic cracking are not high, because C8 olefin synthesized by C4 polymerization is easier to be further catalytically cracked into 2C 4 olefins, and the synthesized C8 olefin is also easier to be subjected to aromatization reaction during the catalytic cracking process to form coke, and the coke yield is higher.

Therefore, it is necessary to develop a combined process of low-carbon olefin polymerization and selective catalytic cracking multi-production chemical materials.

Disclosure of Invention

The present disclosure provides a method for high selectivity catalytic cracking to produce more propylene, which is characterized in that the method comprises:

contacting a heavy hydrocarbon oil raw material with a first catalytic cracking catalyst in a first catalytic reactor to perform a first catalytic reaction to obtain a first reaction mixture containing the first catalytic cracking catalyst; carrying out catalytic reaction on the first reaction mixture in a third catalytic reactor, and carrying out gas-agent separation on the obtained reaction mixture in a settler to obtain a first carbon-deposited catalyst and a first product;

introducing the first product and the second product into a first product separation system for separation to obtain a dry gas fraction, a C3 liquefied gas fraction, a liquefied gas fraction after C3 removal, a light gasoline fraction, a heavy gasoline fraction, a diesel oil fraction and a heavy oil fraction;

introducing the liquefied gas fraction subjected to C3 removal into a first superposition reactor for reaction to obtain first superposition reaction oil gas, further introducing the first superposition reaction oil gas into a first superposition product separation system for separation to obtain unreacted C4 material flow and a first superposition product; introducing at least a part of the first superposed product into a second superposed reactor for reaction to obtain a second superposed product;

introducing the unreacted C4 material flow and the light gasoline fraction into a third superposed reactor for reaction to obtain third superposed reaction oil gas, further introducing the third superposed reaction oil gas into a third superposed product separation system for separation, and separating to obtain an unreacted circulating material flow and a third superposed product;

contacting the unreacted recycle stream, the third stacked product and the second stacked product with a second catalytic cracking catalyst in a second catalytic reactor to perform a second catalytic reaction to obtain a second reaction mixture containing the second catalytic cracking catalyst; and carrying out catalytic reaction on the second reaction mixture in the third catalytic reactor, and carrying out gas-agent separation on the obtained reaction mixture in the settler to obtain a second carbon deposition catalyst and the second product.

In one embodiment, the first polymerization product has a distillation range of 90 to 150 ℃, preferably 100 to 130 ℃, and is predominantly trimethylpentene; the distillation range of the second superimposed product is 240-330 ℃, preferably 260-300 ℃, and the second superimposed product is mainly C16 olefin; the distillation range of the third superposed product is 130-330 ℃, preferably 140-290 ℃, and the third superposed product 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 ℃, and more preferably 9 to 60 ℃;

the olefin content in the light gasoline fraction is 30 to 90 wt%, preferably 45 to 90 wt%, based on the total weight of the light gasoline fraction.

In one embodiment, the reaction conditions of the first polymerization reaction are: the reaction temperature is 300-450 ℃, the pressure is 0.5-2.0 MPa, and the mass space velocity is 1-5 h ^-1 (ii) a The reaction conditions of the second polymerization reactor are as follows: the reaction temperature is 180-280 ℃, the pressure is 1.0-2.0 MPa, and the mass space velocity is 0.7-2.0 h ^-1 (ii) a The reaction conditions of the third stacking reaction are as follows: the reaction temperature is 180-280 ℃, the pressure is 0.8-2.0 MPa, and the mass space velocity is 0.9-4.0 h ^-1 。

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

In one embodiment, the weight ratio of the second polymerization product entering the second catalytic reactor to the heavy hydrocarbon oil feedstock entering the first catalytic reactor is in the range of 0.01 to 0.3:1, preferably 0.05 to 0.15:1; the weight ratio of the third superposed product entering the second catalytic reactor to the heavy hydrocarbon oil raw material entering the first catalytic reactor is 0.01-0.4: 1, preferably 0.05 to 0.2:1.

in one embodiment, the first catalytic cracking catalyst and the second catalytic cracking catalyst each contain a shape selective zeolite having an average pore diameter of less than 0.7nm, the shape selective zeolite being at least one selected from the group consisting of zeolites having an MFI structure, ferrierites, chabazites, cyclospar, erionites, a-type zeolites, epistillomites, and laumontites.

In one embodiment, the operating conditions of the first catalytic 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; the weight ratio of water to oil is 0.05-1: 1.

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

In one embodiment, the first catalytic reactor and the second catalytic reactor are each one selected from the group consisting of a riser reactor, a downer reactor, a fluidized bed reactor, a riser and downer combined reactor, a riser and fluidized bed combined reactor, and a downer and fluidized bed combined reactor.

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

In one embodiment, the first and second carbon-deposited catalysts are stripped in a stripping zone of the settler; regenerating the stripped first and second carbon-deposited catalysts in a regenerator to obtain first and second regenerated catalysts, respectively; feeding said first regenerated catalyst to said first catalytic reactor as said first catalytic cracking catalyst and feeding said second regenerated catalyst to said second catalytic reactor as said second catalytic cracking catalyst.

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

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

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

In one embodiment, the heavy hydrocarbon oil feedstock is at least one selected from the group consisting of petroleum hydrocarbon oils, synthetic oils, coal liquefied oils, oil sand oils, and shale oils, preferably petroleum hydrocarbon oils, and the petroleum hydrocarbon oils are at least one selected from the group consisting of atmospheric gas oils, vacuum gas oils, coker gas oils, deasphalted oils, hydrotreated tail oils, atmospheric residues, vacuum residues, and crude oils.

In one embodiment, the method further comprises:

separating the unreacted recycle stream into a C4 hydrocarbon fraction and a paraffin-rich light gasoline fraction,

and combining the light gasoline fraction rich in alkane and the heavy gasoline fraction to obtain a gasoline product with low olefin content.

The present disclosure also provides a system for high selectivity catalytic cracking for high yield of propylene, comprising:

a reactor system, the reactor system comprising:

a first reactor having a first reactor feed inlet, a first reactor catalyst inlet, and a first reactor product outlet;

a second reactor having a plurality of second reactor feed inlets, second reactor catalyst inlets, and second reactor product outlets; and

a third reactor having a third reactor feed inlet and a third reactor product outlet; wherein the feed inlet of the third reactor is in communication with the first reactor product outlet and the second reactor product outlet such that streams from the first reactor and the second reactor enter the third reactor;

a first product separation system provided with a feedstock inlet and a plurality of fraction outlets; the raw material inlet of the first product separation system is communicated with the product outlet of the third reactor; the plurality of fraction outlets of the first product separation system comprise a dry gas fraction outlet, a C3 liquefied gas fraction outlet, a liquefied gas fraction outlet after C3 removal, a light gasoline fraction outlet, a heavy gasoline fraction outlet, a diesel oil fraction outlet and a heavy oil fraction outlet;

the first superposition reactor is provided with a first superposition reactor raw material inlet and a first superposition reactor product outlet, and the first superposition reactor raw material inlet is communicated with a liquefied gas fraction outlet of the first product separation system after C3 removal;

a first superimposed product separation system provided with a first superimposed product separation system raw material inlet and a plurality of first superimposed product separation system separation product outlets; the raw material inlet of the first superposed product separation system is communicated with the product outlet of the first superposed reactor; a plurality of said first superimposed product separation system separation product outlets including an unreacted C4 stream outlet and a first superimposed product outlet;

the second superposition reactor is provided with a second superposition reactor raw material inlet and a second superposition reactor product outlet; the second superposed reactor raw material inlet is communicated with the first superposed product outlet, and the second superposed reactor product outlet is communicated with one second reactor raw material inlet;

the third superposed reactor is provided with a third superposed reactor raw material inlet and a third superposed reactor product outlet, and the third superposed reactor raw material inlet is communicated with the unreacted C4 material flow outlet;

a third stacked product separation system, the third stacked product separation system provided with a third stacked product separation system raw material inlet and a plurality of third stacked product separation system separation product outlets; the third superposed product separation system raw material inlet is communicated with the third superposed reactor product outlet; the plurality of separation product outlets of the third superposed product separation system comprise an unreacted recycle stream outlet and a third superposed product outlet, and the unreacted recycle stream outlet and the third superposed product outlet are respectively communicated with one raw material inlet of the second reactor.

The inventor of the present disclosure finds that the performance of the catalytic cracking of the superimposed product after the superimposition of different C4 olefins (n-butene, isobutene, trans-2-butene, cis-2 butene) has a large difference in the yield of ethylene and propylene; the products obtained by different-depth superposition reaction of the same C4 olefin have larger difference in the performance of producing more ethylene and propylene through catalytic cracking; different olefins in the liquefied gas are superposed to obtain products, and the products are subjected to catalytic cracking to produce ethylene and propylene in different properties. Meanwhile, the different catalytic cracking reaction processes also have a decisive influence on the performance of the low-carbon olefin polymerization product for cracking to produce more ethylene and propylene.

The method for producing more propylene by high-selectivity catalytic cracking overcomes the problems of high dry gas yield, poor selectivity of the product propylene and high olefin content in the product gasoline in the heavy oil catalytic cracking reaction in the existing process route. The inventor of the present disclosure finds that the catalytic cracked C4 fraction or/and light gasoline are firstly subjected to a polymerization reaction to form a polymerization product, and propylene selectivity of different polymerization products subjected to a catalytic cracking reaction has a large difference, for example, C8 olefin as a product formed by polymerizing C4 olefin is more likely to be cracked into butene rather than propylene if the number of branches (methyl number) of C8 olefin is larger, such as n-octene, methylheptene, dimethylhexene, and trimethylpentene, the methyl number of n-octene, methylheptene, dimethylhexene, and trimethylpentene is gradually increased, and the yield of propylene produced by catalytic cracking is gradually reduced. Although all are C8 olefins, the yield of propylene produced by catalytic cracking of n-octene is more than twice the yield of propylene produced by catalytic cracking of trimethylpentene. The production of larger olefin molecules (e.g., C16 olefins) is required for the polymethyl C8 olefins (e.g., trimethylpentene), and the catalytic cracking can greatly increase the propylene yield. The de-C3 liquefied gas contains different C4 olefins (n-butene, isobutene, trans-2-butene and cis-2-butene), and a certain C4 olefin needs to be subjected to polymerization to generate a specific type of polymerization product, so that the propylene yield of catalytic cracking of the polymerization product is maximized. It is therefore desirable to obtain specific superimposed products for catalytic cracking propylene-rich processes.

Specifically, compared with the traditional method for producing low-carbon olefins by catalytic conversion of heavy hydrocarbon oil, the method disclosed by the invention has the following beneficial effects of any one or more, preferably all the following effects:

1. in the prior art, C4 fraction and light gasoline are directly recycled to a catalytic cracking or catalytic cracking device for reaction, but molecular chains of the C4 fraction and the light gasoline are short and difficult to crack, and cracking needs high reaction severity, namely high reaction temperature, high catalyst-oil ratio, high retention time and the like, so that high dry gas and coke yield is caused, and propylene selectivity is poor. Therefore, although the yield of propylene can be increased by cyclically cracking the C4 fraction and the light gasoline fraction by a catalytic cracking unit in the prior art, the conversion per pass of the reaction is low, a large recycle ratio is required, the selectivity of propylene is low, the selectivity of dry gas is high, the olefin in the product gasoline cannot be completely converted, and the olefin content of the product gasoline is high.

According to the method, firstly, isobutene in the C4 fraction is subjected to selective polymerization reaction to generate trimethylpentene, and simultaneously, the trimethylpentene is separated from other C4 olefins (such as n-butene, trans-2-butene and cis-2-butene), because the yield of propylene generated by the catalytic cracking of the trimethylpentene is very low, the propylene is not directly returned to a catalytic cracking device for reaction, but is further synthesized into C16 olefins, C20 olefins and the like, the synthesized C16 olefins can be returned to the catalytic cracking device for reaction, and the yield of the propylene is greatly improved. Therefore, the problem of low yield of propylene in the C4 olefin by catalytic cracking of isobutene is solved, and the phenomenon that isobutene and other C4 olefins generate a polymerization reaction to interfere the catalytic cracking reaction of other C4 olefins is avoided.

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

The C4 fraction and the light gasoline contain alkane and olefin, the cracking of the alkane needs higher reaction severity than the cracking of the olefin, while the C4 fraction and the light gasoline which simultaneously contain the alkane and the olefin are circularly cracked in the prior art, and the requirements of the two on the reaction conditions are hardly 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.

The method firstly converts and separates C4 olefin and light gasoline olefin through different superposition reactions, and returns the rest C4 hydrocarbon rich in alkane and light naphtha to a catalytic cracking reactor in the form of unreacted circulating material flow, so that high-rigor dehydrogenation and cracking reactions are firstly carried out, and the alkane cracking effect is improved. The second superposed product (C16 olefin, C20 olefin, etc.) and the third superposed product (C9, C10, C12, C13 olefin, etc.) which are easy to crack enter the regenerated catalytic cracking catalyst with reduced temperature to react, so that the reaction severity is greatly reduced, and the selectivity of propylene is improved.

The method disclosed by the invention realizes the high-selectivity reconversion of different hydrocarbons in C4 molecules with different molecular structures and light gasoline into propylene.

Drawings

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

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

Description of the reference numerals

1. First catalytic reactor

2-1 second catalytic reactor

2-2 third catalytic reactor

7. Stripping zone

6. Settling vessel

9. Regenerator

61. First product separation system

70-1 first polymerization reactor

70-2 second polymerization reactor

70-3 third stacked reactor

62. First superimposed product separation system

63. Third superimposed product separation system

11. First catalytic cracking catalyst transfer line

12. Second catalytic cracking catalyst transfer line

17. Spent catalyst transfer line

21. Pipe line (injecting heavy hydrocarbon oil feedstock)

20. Pipeline (transportation catalytic cracking reaction oil gas)

24. Pipeline (conveying dry gas)

25. Pipeline (transport C3)

26. Pipeline (transportation de-C3 back liquefied gas)

27. Pipeline (transportation light petrol)

30. Pipeline (transportation heavy petrol)

31. Pipeline (transportation diesel)

32. Pipeline (transporting heavy oil)

41. Pipe line (injecting atomized steam)

43. Pipe line (injecting atomized steam)

44. Pipe line (injecting atomized steam)

45. Pipe line (injecting atomizing steam)

47. Pipe (steam stripping injection)

51. Pipeline (injection pre-lifting medium)

52. Pipeline (injection pre-lifting medium)

90. Pipeline (air injection)

33. Pipeline (first superposition reaction oil gas)

35. Pipeline (third superposition reaction oil gas)

34-1 line (unreacted C4)

34-3 line (unreacted recycle stream)

37-1 line (for the first superimposed product)

37-2 line (for the second superimposed product)

37-3 pipeline (for third superposed product)

90. Pipeline (air conveying)

91. Pipeline (flue gas)

100. External heat collector (take away the excess heat of regenerator, reduce regeneration temperature)

Detailed Description

The technical solution of the present invention is further explained below according to specific embodiments. The scope of protection of the invention is not limited to the following examples, which are set forth for illustrative purposes only and are not intended to limit the invention in any way.

The present disclosure provides a method for high selectivity catalytic cracking for producing propylene in high yield, which comprises:

contacting the heavy hydrocarbon oil raw material with a first catalytic cracking catalyst in a first catalytic reactor to perform a first catalytic reaction to obtain a first reaction mixture containing the first catalytic cracking catalyst; carrying out catalytic reaction on the first reaction mixture in a third catalytic reactor, and carrying out gas-agent separation on the obtained reaction mixture in a settler to obtain a first carbon-deposited catalyst and a first product;

In one embodiment, the first polymerization product has a distillation range of 90 to 150 ℃, preferably 100 to 130 ℃, and is predominantly trimethylpentene, wherein the amount of trimethylpentene is 65 to 99% by weight based on the total weight of the first polymerization product. In one embodiment, the second polymerization product has a distillation range of 240 to 330 ℃, preferably 260 to 300 ℃, and is predominantly a C16 olefin (C16 olefin having a boiling point of 284 ℃), wherein the amount of C16 olefin is 60 to 90% based on the total weight of the second polymerization product; the distillation range of the third superimposed product is 130-330 ℃, preferably 140-290 ℃, and the third superimposed product is mainly C9-C12 olefin (the boiling point of C9 olefin is 147 ℃, the boiling point of C12 olefin is 213 ℃, and the boiling point of C13 olefin is 233 ℃), wherein the amount of C9-C12 olefin is 60-90% of the total weight of the third superimposed 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 first catalytic reactor 1 is a riser reactor, a second catalytic reactor 2-1 is a riser reactor, and a third catalytic reactor 2-2 is a fluidized bed reactor. The fluidized bed 2-2 in the settler 6 is located above the stripping zone 7.

The first catalytic cracking catalyst (hot regenerated catalyst) enters the bottom of the riser of the first catalytic reactor 1 from the regenerator 9 via the first catalytic cracking catalyst transfer line 11 and is accelerated to flow upward by the pre-lift medium injected via line 51. The preheated heavy hydrocarbon oil raw material is mixed with atomized steam from a pipeline 41 through a pipeline 21 and then injected into a riser of the first catalytic 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 mixture of the reaction oil gas and the catalyst in the riser 1 can be further introduced into a third catalytic reactor 2-2 for further reaction through an outlet, the first carbon-deposited catalyst after the reaction is introduced into a stripping zone 7, the separated reaction oil gas (first product) is sent into a subsequent first product separation system 61 through a settler 6 and a pipeline 20 at the top of the settler 6 for product separation, and products such as dry gas, C3 fractions (propylene and propane), liquefied gas fractions (C4 fractions) after C3 removal, light gasoline fractions, heavy gasoline fractions, diesel oil fractions, heavy oil fractions and the like are obtained after the separation (respectively introduced through

pipelines

24, 25, 26, 27, 30, 31 and 32).

The second catalytic cracking catalyst (hot regenerated catalyst) enters the bottom of the riser of the second catalytic reactor 2-1 from the regenerator 9 through a second catalytic cracking catalyst delivery line 12 and is accelerated to flow upward by the pre-lift medium injected through line 52. The unreacted recycle stream 34-3 from the second product separation system 62 is mixed with the atomized steam from the pipeline 45 and then injected into the second catalytic reactor 2-1, optionally, part of the light gasoline fraction 27 is mixed with the atomized steam and then injected into the second catalytic reactor 2-1, the second superimposed product 37-3 from the third product separation system 63 is mixed with the atomized steam from the pipeline 43 and then injected into the second catalytic reactor 2-1, the second superimposed product 37-2 from the second product separation system 62 is mixed with the atomized steam from the pipeline 44 and then injected into the second catalytic reactor 2-1, and the weight ratio of the water vapor and the hydrocarbon oil raw material (i.e., the unreacted recycle stream 34, the light gasoline fraction 27, the third superimposed product 37-3, and the second superimposed product 37-2 injected into the second catalytic reactor 2-1) in the riser reactor 2-1 is (0.1-1): 1, the outlet temperature of the riser reactor 2-1 is 520-750 ℃, the reaction time in the riser reactor 2-1 is 0.1-3 seconds, and the weight ratio of the catalyst to the hydrocarbon oil raw material (namely, the unreacted recycle stream 34 injected into the second catalytic reactor, the light gasoline fraction 27, the third superimposed product 37-3 and the second superimposed product 37-2) is 6-40. Mixing of reaction oil gas and catalyst of riser reactor 2-1The material is further introduced into a fluidized bed of a third catalytic reactor 2-2 through a riser outlet to continue reacting, the reaction temperature of the fluidized bed 2-2 is 450-750 ℃, and the weight hourly space velocity is 1-30 h ^-1 After the reaction, the oil gas and a part of the carbon-deposited spent catalyst enter a settler 6 for separation through a fluidized bed reactor 2-2, the separated second carbon-deposited catalyst enters a stripping zone 7, and the reaction oil gas (second product) is sent to a subsequent first product separation system 61 for product separation through the settler 6 and a pipeline 20 at the top of the settler.

Introducing the liquefied gas (C4 fraction) 26 subjected to C3 removal into a first superposition reactor 70-1 to react to obtain a first superposition reaction oil gas 33, further introducing the first superposition reaction oil gas 33 into a second product separation system 62 to separate the first superposition reaction oil gas into an unreacted C4 fraction 34-1 and a first superposition product 37-1, and further introducing the first superposition product 37-1 into a second superposition reactor 70-2 to react to generate a second superposition product 37-2; and introducing the unreacted C4 fraction 34-1 or/and the light gasoline fraction 27 into a third superposed reactor 70-3 to react to obtain a third superposed reaction oil gas 35, and further introducing the third superposed reaction oil gas 35 into a third product separation system 63 to separate into an unreacted recycle stream 34-3 and a third superposed product 37-3.

The steam stripping steam is injected into the steam stripping zone 7 through a pipeline 47 and contacts with the carbon-deposited spent catalyst in a countercurrent way, so that the reaction oil gas carried by the spent catalyst is stripped as completely as possible. Air is injected into the regenerator 9 through a pipeline 90, the stripped first carbon-deposited catalyst and the stripped second carbon-deposited catalyst are sent into the regenerator 9 through a spent agent conveying pipeline 17, and are contacted with the heated air and regenerated at the temperature of 600-800 ℃, so that a first regenerated catalyst and a second regenerated catalyst are obtained and are used as a first catalytic cracking catalyst and a second catalytic cracking catalyst for recycling. The regeneration flue gas is led out through a line 91. In fig. 1, 100 is an external heat remover, which is used for removing heat of the regenerator through heat exchange when necessary and reducing the regeneration temperature.

As mentioned above, the heavy hydrocarbon oil feedstock is contacted with a first catalytic cracking catalyst in a first catalytic reactor to perform a first catalytic reaction to obtain a first reaction mixture containing the first catalytic cracking catalyst; and carrying out catalytic reaction on the first reaction mixture in a third catalytic reactor, and carrying out gas-agent separation on the obtained reaction mixture in a settler to obtain a first carbon-deposited catalyst and a first product.

According to the present disclosure, a heavy hydrocarbon oil feedstock is contacted with a first catalytic cracking catalyst in a fluidized state in a first catalytic reactor to perform a first catalytic reaction. The operating conditions of the first catalytic reaction may include: the reaction temperature is 480 to 600 ℃, for example 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 hydrocarbon oil raw material) is (5-15): 1, or (6 to 12): 1, or (8 to 10): 1; the weight ratio of water to oil (namely the weight ratio of the water vapor to the heavy hydrocarbon oil raw material) is (0.05-1): 1, 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 first catalytic reactor.

Introducing the first product into a first 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); the liquefied gas product can be further separated into a C3 liquefied gas fraction (propylene and propane) and a liquefied gas fraction (C4 fraction) after C3 removal, and the gasoline product can be further separated into a light gasoline fraction and a heavy gasoline fraction.

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

In one embodiment, the light gasoline has a boiling range of 9 to 150 ℃, more preferably 9 to 100 ℃, and even 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 primarily hydrogen, methane, ethylene and ethane, the C3 liquefied gas fraction is propylene and propane, the liquefied gas fraction after C3 removal 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 first, second and third superimposed reactors can be selected from one or a combination of fixed bed, fixed fluidized bed reactor and circulating fluidized bed reactor. The first superposition reactor, the second superposition reactor and the third superposition reactor can be arranged independently or together according to requirements, so that the space of a plant is saved.

According to the present disclosure, the reaction conditions of the first polymerization reactor are: the reaction temperature is 300-450 ℃, the pressure is 0.5-2.0 MPa, and the mass space velocity is 1-5 h ^-1 . The reaction conditions of the second polymerization reactor are as follows: the reaction temperature is 180-280 ℃, the pressure is 1.0-2.0 MPa, and the mass space velocity is 0.7-2.0 h ^-1 . The reaction conditions of the third superposed reactor are as follows: the reaction temperature is 180-280 ℃, the pressure is 0.8-2.0 MPa, and the mass space velocity is 0.9-4.0 h ^-1 。

The first polymerization catalyst comprises 1-12 mass% of NiO, 45-82 mass% of amorphous aluminum silicate and 10-50 mass% of alumina. The second polymerization catalyst comprises NiO 1-12 wt%, amorphous aluminum silicate 45-82 wt% and alumina 10-50 wt%. The third catalyst for the superposition reaction comprises 1-20 mass percent of NiO, 40-80 mass percent of HZSM-5 zeolite and 10-50 mass percent of alumina.

The liquefied gas fraction after C3 removal mainly comprises C4 alkane and C4 olefin, under the action of an acid catalyst, 2 isobutene are subjected to a polymerization reaction selectively to generate a first polymerization product (trimethylpentene) in a first polymerization reactor at a lower reaction temperature and a higher reaction pressure, the obtained first polymerization product trimethylpentene and unreacted C4 are separated, the first polymerization product trimethylpentene is further subjected to a deeper polymerization reaction in a second polymerization reactor to generate a second polymerization product (C16 olefin, C20 olefin and the like), and therefore the isobutene which is difficult to crack and the trimethylpentene which is poor in cracking propylene selectivity are converted into macromolecular olefins such as C16 olefin and C20 olefin which are easy to crack and are high in propylene selectivity. Unreacted C4 (n-butene, trans-2-butene, cis-2-butene) or/and olefin from light gasoline are subjected to a polymerization reaction in a third polymerization reactor to generate hydrocarbons such as C9, C10, C12, C13 and the like, for example, 1-pentene in the light gasoline can be subjected to a polymerization reaction with n-butene to generate C9 olefin, and the generated C9 olefin has only one methyl group or does not contain the methyl group, so that the selectivity of catalytic cracking to propylene is high; the C9 olefin further generates a polymerization reaction with C4 olefin (n-butene, trans-2-butene, cis-2-butene) to generate C13 olefin with less methyl branch, thereby greatly reducing the severity of other C4 olefin catalytic cracking reaction, greatly improving the selectivity of cracking recycle stream into propylene, and reserving the C4 alkane and the alkane of light gasoline because the polymerization reaction is difficult to generate. The process simultaneously realizes the selective combination of C4 molecules with different structures and different hydrocarbons of light gasoline into a specific superposed product which is easy to crack into propylene.

And introducing the third superposed reaction oil gas into a third product separation system for separation to obtain an unreacted recycle stream. The third product separation system carries out oil-gas separation according to the molecular type and the distillation range, and the obtained unreacted recycle stream is mainly C4 alkane or light gasoline rich in alkane, the distillation range is-12 to 120 ℃, and the optimal selection is-7 to 60 ℃; the distillation range of the first superimposed product is 90-150 ℃, preferably 100-130 ℃, and the first superimposed product is mainly trimethylpentene; the distillation range of the second superimposed product is 240-330 ℃, preferably 260-300 ℃, and the second superimposed product is mainly C16 alkene; the distillation range of the third superposed product is 130-330 ℃, preferably 140-290 ℃, and the third superposed product mainly contains C9-C12 alkene and the like.

The method of the present disclosure further comprises: contacting the unreacted circulating material flow, the light gasoline fraction, the third superposed product and the second superposed product in a second catalytic reactor with a second catalytic cracking catalyst to perform a second catalytic reaction to obtain a second reaction mixture containing the second catalytic cracking catalyst; and carrying out catalytic reaction on the second reaction mixture in a third catalytic reactor, and carrying out gas-agent separation on the obtained reaction mixture in the settler to obtain a second carbon-deposited catalyst and a second product. The second product can be separated from the first product by passing the second product into a first product separation system.

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

The unreacted recycle stream (mainly C4 alkane or light gasoline rich in alkane) returns to the second catalytic reactor, firstly contacts with the high-temperature regenerated catalyst, and carries out high-rigor dehydrogenation and cracking reaction at the reaction temperature of 640-750 ℃, thereby improving the alkane cracking effect. A light gasoline fraction (containing C5, C6 and C7 olefins) in which the C6 and C7 olefins are relatively susceptible to cracking reactions, each C6 olefin being susceptible to cleavage directly from the molecular middle to produce 2 propylene in the second catalytic reactor at a relatively mild reaction temperature of 520 to 600 ℃, each C7 olefin molecule being susceptible to cleavage directly from the molecular middle to produce 1 propylene and 1 butene molecule, the lower molecular number C5 olefins being largely retained in the light gasoline and finally introduced into the third cascade reactor. The reaction severity required by hydrocarbon cracking reaction of the third superposed product (C9, C10, C12 olefin and the like) and the second superposed product (C16, C20 olefin and the like) is greatly reduced as the molecular chain is longer, so that the reaction is finally introduced into a second catalytic reactor and reacts at a more moderate reaction temperature of 520-560 ℃, for example, the C12 olefin molecule is easy to directly generate 2C 6 olefin molecules by intermediate chain scission, each C6 olefin molecule is further generated by intermediate scission to generate 2 propylene, and the propylene selectivity is greatly improved. The whole reaction process is set, so that the C4 hydrocarbon and the light gasoline are converted into propylene with high selectivity and high conversion rate according to the difference of molecular structures to design a reaction process.

According to the present disclosure, the second polymerization product may be produced by the process of the present disclosure, or may be produced by other devices. The olefin content of the second superimposed product may be 60 to 100 wt%, preferably 80 to 100 wt%, for example 80 to 90 wt% or 85 to 95 wt% or 90 to 100 wt%, based on the total weight of the second superimposed product. The olefin content of the second superimposed product may be 60 to 100 wt. -%, preferably 80 to 100 wt. -%, for example 80 to 90 wt. -% or 85 to 95 wt. -% or 90 to 100 wt. -%, based on the total weight of the third superimposed product.

According to the present disclosure, the weight ratio of the second superimposed product to the heavy hydrocarbon oil feedstock is (0.01-0.3): 1, preferably (0.05 to 0.15): 1. the weight ratio of the third superposed product to the heavy hydrocarbon oil raw material is (0.01-0.4): 1, preferably (0.05 to 0.2): 1. the weight ratio of the unreacted recycle material flow to the heavy hydrocarbon oil raw material is (0.01-0.3): 1, preferably (0.05 to 0.2): 1. the weight ratio of the light gasoline fraction to the heavy hydrocarbon oil 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 unreacted recycle stream into a C4 hydrocarbon fraction and an alkane-rich light gasoline fraction, combining the alkane-rich light gasoline fraction and the heavy gasoline fraction, may yield a gasoline product having a low olefin content. The olefin content in the low olefin gasoline product may be 6-16%. According to the present disclosure, the C4 paraffins in the C4 hydrocarbon fraction account for more than 95%, and the paraffin content of the C5 and above in the paraffin-rich light gasoline fraction accounts for more than 85%. The separation of the unreacted recycle stream into a C4 hydrocarbon fraction and a light paraffin-rich gasoline fraction may be carried out using conventional fractionation or rectification methods.

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

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, and are preferably the same catalytic cracking catalyst, which are conventionally used in the field of catalytic cracking reactions.

The present disclosure is not particularly limited with respect to the specific kinds of the first catalytic cracking catalyst and the second catalytic cracking catalyst. Preferably, the first catalytic cracking catalyst and the second catalytic cracking catalyst each contain a shape-selective zeolite having an average pore diameter of less than 0.7nm, and the shape-selective zeolite may be at least one selected from the group consisting of zeolite having an MFI structure, ferrierite, chabazite, dachiardite, erionite, a-type zeolite, epistilbite, and turbid zeolite. Wherein the MFI structure zeolite may be one or more of ZSM-5 and ZRP series zeolites, and may be one or more of ZSM-5 and ZRP series zeolites modified with 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, based on the weight of the catalytic cracking catalyst on a dry basis (weight calcined at 800 ℃ for 1 hour), 15 to 50 wt% of clay on a dry basis, 15 to 50 wt% of a molecular sieve which is a zeolite of MFI structure or consists of 25 to 100 wt% of a zeolite of MFI structure and 0 to 75 wt% of another zeolite other than the zeolite of MFI structure, and 10 to 35 wt% of a binder on a dry basis; the MFI structure zeolite is preferably a ZSM-5 molecular sieve and/or an 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 kaolin, halloysite, montmorillonite, diatomaceous earth, halloysite, pseudohalloysite, saponite, rectorite, sepiolite, attapulgite, hydrotalcite and bentonite. The binder is one or more of acidified pseudo-boehmite, aluminum sol, silica sol, magnesium aluminum sol, zirconium sol and titanium sol, preferably acidified pseudo-boehmite, aluminum sol and the like.

In accordance with the present disclosure, the first catalytic reactor and the second catalytic reactor may be catalytic conversion reactors well known to those skilled in the art, for example, the first catalytic reactor and the second catalytic reactor are each one selected from a riser reactor, a downer reactor, a fluidized bed reactor, a riser and downer combined reactor, a riser and fluidized bed combined reactor, and a downer and fluidized bed combined 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 variable-diameter riser reactors, downer reactors and fluidized bed reactors.

According to the present disclosure, the products of the first catalytic reaction and the second catalytic reaction (the first product and the second product) can both continue the catalytic reaction in the third catalytic reactor, which can further prolong the reaction residence time, improve the reaction conversion rate, and increase the yield of propylene.

In accordance with the present disclosure, the third catalytic 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 transport bed reactor, and a dense phase fluidized bed. The fluidized bed reactor can be in a constant-diameter fluidized bed structure or a variable-diameter fluidized bed structure. What is needed isThe operating conditions of the third catalytic reactor may be: the reaction temperature is 450 to 750 ℃, for example 480 to 600 ℃, or 500 to 580 ℃, or 510 to 560 ℃, or 520 to 550 ℃, preferably 510 to 560 ℃; the weight hourly space velocity is 1-30 h ^-1 For example, 3 to 28h ^-1 Or 5 to 25h ^-1 Or 6 to 20 hours ^-1 。

According to the present disclosure, the method further comprises:

stripping the first and second carbon deposited catalysts in a stripping zone of the settler,

regenerating the stripped first and second carbon deposited catalysts in a regenerator to obtain first and second regenerated catalysts, respectively;

feeding said first regenerated catalyst as said first catalytic cracking catalyst into said first catalytic reactor and feeding said second regenerated catalyst as said second catalytic cracking catalyst into said second catalytic reactor.

In accordance with the present disclosure, the first and second carbon-deposited catalysts may be stripped in a stripping zone of the settler. By stripping, the reaction oil gas carried by the spent catalyst can be stripped as clean as possible. The absolute pressure in the settler may be 0.1-0.40 MPa.

According to the present disclosure, the stripped first and second carbon deposited catalysts are introduced into a regenerator for regeneration to obtain first and second regenerated catalysts, the first regenerated catalyst is sent into a first catalytic reactor as the first catalytic cracking catalyst, and the second regenerated catalyst is sent into a second catalytic reactor as the second catalytic cracking catalyst. Thereby, the first catalytic cracking catalyst and the second catalytic cracking catalyst can be recycled and reused. In the regenerator, the stripped first and second carbon-deposited catalysts are contacted with heated air and regenerated at 600-800 ℃.

According to the present disclosure, the first regenerated catalyst has a temperature of 560 to 800 ℃ and a carbon deposit content of 0.01 to 0.1 wt.% based on the dry weight of the first regenerated catalyst. In another embodiment, the second regenerated catalyst according to the present disclosure has a temperature of 560 to 800 ℃ and a carbon deposit content of 0.01 to 1.9 wt.%, preferably 0.9 to 1.3 wt.%, and more preferably 0.91 to 0.99 wt.%, based on the dry weight of the second regenerated catalyst.

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

a reactor system, the reactor system comprising:

a first reactor 1 having a first reactor feed inlet, a first reactor catalyst inlet, and a first reactor product outlet;

a second reactor 2-1 having a plurality of second reactor feed inlets, second reactor catalyst inlets, and second reactor product outlets; and

a third reactor 2-2 having a third reactor feed inlet and a third reactor product outlet; wherein the feed inlet of the third reactor is in communication with the first reactor product outlet and the second reactor product outlet such that streams from the first reactor and the second reactor enter the third reactor;

a first product separation system 61 provided with a feed inlet and a plurality of fraction outlets; the raw material inlet of the first product separation system is communicated with the product outlet of the third reactor; the plurality of fraction outlets of the first product separation system comprise a dry gas fraction outlet, a C3 liquefied gas fraction outlet, a liquefied gas fraction outlet after C3 removal, a light gasoline fraction outlet, a heavy gasoline fraction outlet, a diesel oil fraction outlet and a heavy oil fraction outlet;

the first superposition reactor 70-1 is provided with a first superposition reactor raw material inlet and a first superposition reactor product outlet, and the first superposition reactor raw material inlet is communicated with a liquefied gas fraction outlet of the first product separation system after C3 removal;

a first laminated product separation system 62 having a first laminated product separation system feed inlet and a plurality of first laminated product separation system separation product outlets; the raw material inlet of the first superposed product separation system is communicated with the product outlet of the first superposed reactor; the plurality of first stacked product separation system separation product outlets comprises an unreacted C4 stream outlet and a first stacked product outlet;

a second superposition reactor 70-2, wherein the second superposition reactor is provided with a second superposition reactor raw material inlet and a second superposition reactor product outlet; the second superposed reactor raw material inlet is communicated with the first superposed product outlet, and the second superposed reactor product outlet is communicated with one second reactor raw material inlet;

a third overlap reactor 70-3, wherein the third overlap reactor is provided with a third overlap reactor raw material inlet and a third overlap reactor product outlet, and the third overlap reactor raw material inlet is communicated with the unreacted C4 material flow outlet;

a third stacked product separation system 63 having a third stacked product separation system feed inlet and a plurality of third stacked product separation system separation product outlets; the third superposed product separation system raw material inlet is communicated with the third superposed reactor product outlet; the plurality of separation product outlets of the third superposed product separation system comprise an unreacted recycle stream outlet and a third superposed product outlet, and the unreacted recycle stream outlet and the third superposed product outlet are respectively communicated with one raw material inlet of the second reactor.

In one embodiment, the first reactor product outlet and the second reactor product outlet are located inside the third reactor such that the streams from the first and second reactors (reaction mixture comprising product and catalyst) can pass directly into the third reactor to continue the reaction. In the present disclosure, the term "in communication" means that two in communication can be connected by a pipeline.

The embodiments described in the method of the present disclosure are also applicable to the system of the present disclosure, and are not described herein again.

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

The first and second catalytic cracking catalysts used in the following examples and comparative examples are cracking catalysts manufactured by the Chinese petrochemical catalyst, qilu division, having a trade mark of OMT, and having specific properties as shown in Table 1-1, and comprise shape selective zeolite having an average pore diameter of less than 0.7 nm.

Example 1

Example i illustrates the effect of the process provided by the present disclosure on high selectivity propylene production during a high selectivity catalytic cracking process for high propylene production.

The experiment was carried out using a medium-sized apparatus for continuous reaction-regeneration operation having three reactors, in which the first catalytic reactor was a riser, and the riser reactor had an inner diameter of 16 mm and a height of 3800 mm. The second catalytic reactor was a riser reactor with an internal diameter of 16 mm and a height of 3200 mm. The outlet of the first riser reactor and the outlet of the second riser reactor are introduced into a fluidized bed reactor which is a third catalytic 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 first catalytic reactor through a first catalytic cracking catalyst inclined tube, and flows upwards under the action of pre-lifting steam. The heavy hydrocarbon oil raw material is atmospheric residue (the main properties are shown in table 2) which is heated to 350 ℃ by a preheating furnace and then mixed with atomized water vapor, and the mixture is sprayed into a first catalytic reactor through a feeding nozzle and contacts with a hot first regenerated catalyst to carry out catalytic conversion reaction. After reaction, oil gas (a first product) and a first carbon deposit catalyst enter a fluidized bed of a third catalytic reactor from an outlet of a riser of the first catalytic reactor for continuous reaction, and then enter a settler for rapid separation, and the first carbon deposit 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 second catalytic reactor through a second catalytic cracking catalyst inclined tube, and flows upwards under the action of pre-lifting steam. Unreacted recycle stream (-12-120 ℃), light gasoline (9-120 ℃), a third superposed product (140-250 ℃), and a second superposed product (260-290 ℃) enter a second catalytic reactor in sequence to contact with a hot second regenerated catalyst for catalytic reaction. The weight ratio of the unreacted recycle stream to the heavy hydrocarbon oil feedstock is 0.1:1, the weight ratio of the light gasoline to the heavy hydrocarbon oil raw material is 0.05:1, the weight ratio of the second superimposed product to the heavy hydrocarbon oil raw material is 0.11:1, the weight ratio of the third superposed product to the heavy hydrocarbon oil raw material is 0.24:1. introducing the reaction oil mixture from the second catalytic reactor into a fluidized bed of a third catalytic reactor at the outlet of a riser for further reaction, separating oil gas (a second product) and a second carbon-deposited catalyst after the reaction in a settler through the fluidized bed reactor, and allowing the separated second carbon-deposited catalyst to enter a stripping zone for stripping under the action of gravity. And the first carbon-deposited catalyst and the second carbon-deposited catalyst after steam stripping enter a regenerator through a spent agent conveying pipe, are contacted with heated air and are regenerated at 700 ℃, and the obtained hot first regenerated catalyst and the second regenerated catalyst are respectively returned to the first catalytic reactor and the second catalytic reactor for recycling. And (3) leading the oil gas (the first product and the second product) after the reaction out of the settler, introducing the oil gas into a first product separation system for product separation to obtain a gas product and various liquid products, and simultaneously partially separating to obtain a C4 fraction, light gasoline and heavy gasoline.

Introducing the C4 fraction into a first superposition reactor to carry out superposition reaction to obtain first superposition reaction oil gas, further introducing the first superposition reaction oil gas into a first superposition product separation system to carry out separation to obtain unreacted C4 and a first superposition product (100-130 ℃); introducing the first superposed product into a second superposed reactor for reaction to obtain a second superposed product (at 260-290 ℃); introducing the unreacted C4 and the light gasoline fraction into a third superposed reactor for reaction to obtain third superposed reaction oil gas, further introducing the third superposed reaction oil gas into a third superposed product separation system for separation to obtain unreacted circulating material flow (-12-120 ℃) and a third superposed product (140-250 ℃);

the first polymerization catalyst and the second polymerization catalyst are LXC-10 catalysts generated by the petrochemical engineering scientific research institute of China petrochemical Co., ltd, and the third polymerization catalyst used is an SXC-6 catalyst produced by the petrochemical engineering scientific research institute of China petrochemical Co., ltd. The polymerization reactor is a fixed bed reactor; the first polymerization reaction conditions are as follows: the reaction temperature is 350 ℃, the reaction pressure is 1.1Mpa, and the weight hourly space velocity is 3.0h ^-1 (ii) a The second polymerization reaction conditions are as follows: the reaction temperature is 250 ℃, the reaction pressure is 1.8Mpa, and the weight hourly space velocity is 1.0h ^-1 (ii) a The third stacking reaction condition is as follows: 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 described below.

The unreacted recycle stream may further separate the C4 hydrocarbons and the light paraffin-rich gasoline, a portion of which may be blended into a low olefin gasoline product.

The main operating conditions and results are listed in table 3.

The SXC-6 catalyst is prepared by the following steps:

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 a strip 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 HZSM-5-Al particles with the diameter of 0.6-0.9 mm 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 composite carrier (11 g) was mixed with Ni (NO) (5.7 g) ₃ ) ₂ ﹒6H ₂ Soaking the solution prepared by O for 6h by conventional method, filtering, drying at 100 deg.C, and adding N ₂ Activating for 6 hours at 450 ℃ in atmosphere to obtain the catalyst of the invention, the specific surface of which is 320m ² Pore volume was 0.26ml/g.

The preparation process of the LXC-10 catalyst is as follows:

mixing 12.3g of amorphous aluminum silicate powder and 7.3g of alumina powder, uniformly mixing 15g of deionized water, adding a proper amount of dilute nitric acid solution, kneading and extruding into a strip 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 into particles with the diameter of 0.6-0.9 mm to obtain amorphous aluminum silicate-Al ₂ O ₃ And (3) a composite carrier. SiO of amorphous aluminum silicate powder ₂ /Al ₂ O ₃ The molar ratio was 10.

The obtained composite carrier (10 g) was mixed with Ni (NO) (2.2 g) ₃ ) ₂ ﹒6H ₂ Soaking the solution prepared by O for 6h by conventional method, filtering, drying at 100 deg.C, and adding N ₂ Activating for 6 hours at 450 ℃ in atmosphere to obtain the catalyst of the invention, the specific surface of which is 310m ² 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 1-1

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

The first catalytic cracking catalyst is a first regenerated catalyst with the temperature of 680 ℃, enters the bottom of a riser reactor of the first catalytic reactor through a first catalytic cracking catalyst inclined tube, and flows upwards under the action of pre-lifting steam. The heavy hydrocarbon oil raw material is atmospheric residue (the main properties are shown in table 2) which is heated to 350 ℃ by a preheating furnace and then mixed with atomized water vapor, and the mixture is sprayed into a first catalytic reactor through a feeding nozzle and contacts with a hot first regenerated catalyst to carry out catalytic conversion reaction. The reaction oil gas (first product) and the first carbon-deposited catalyst enter a third catalytic reactor fluidized bed from the outlet of a first catalytic reactor riser, and then further enter a settler for rapid separation, and the first carbon-deposited 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 second catalytic reactor through a second catalytic cracking catalyst inclined tube, and flows upwards under the action of pre-lifting steam. C4 fraction and light gasoline (9-120 ℃) enter a second catalytic reactor in sequence to contact with a hot second regenerated catalyst for catalytic reaction. The weight ratio of the C4 fraction to the heavy hydrocarbon oil feedstock is 0.25:1, the weight ratio of the light gasoline to the heavy hydrocarbon oil raw material is 0.25:1. introducing the reaction oil mixture from the second catalytic reactor into a fluidized bed of a third catalytic reactor at the outlet of a riser for further reaction, separating oil gas (a second product) and a second carbon-deposited catalyst after the reaction in a settler through the fluidized bed reactor, and allowing the separated second carbon-deposited catalyst to enter a stripping zone for stripping under the action of gravity. And the first carbon-deposited catalyst and the second carbon-deposited catalyst after steam stripping enter a regenerator through a spent agent conveying pipe, are contacted with heated air and are regenerated at 700 ℃, and the obtained hot first regenerated catalyst and the second regenerated catalyst are respectively returned to the first catalytic reactor and the second catalytic reactor for recycling. Reaction oil gas from the first catalytic reactor and the third catalytic reactor is led out from the settler together, introduced into a product separation system for product separation to obtain gas products and various liquid products, and meanwhile, part of the reaction oil gas is separated to obtain C4 fraction, light gasoline and heavy gasoline.

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

Of these, the reaction conditions of the second and third catalytic reactors were more severe than in example 1 (the temperatures of the second and third catalytic reactors were higher compared to example 1), and the main operating conditions and results are listed in table 3.

Comparative examples 1 to 2

Comparative examples 1-2 illustrate the effect of cyclically cracking light gasoline and C4 fractions to increase propylene yield during the process of high-selectivity catalytic cracking to increase propylene yield.

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

The main operating conditions and results are listed in table 3.

Example 2

Example 2 illustrates the effect of the process provided by the present disclosure on high selectivity propylene production during a high selectivity catalytic cracking process for high propylene production.

The reaction apparatus used was the same as in example i. The main experimental procedure is the same as in example i. The weight ratio of the unreacted recycle stream (-12-5 ℃) to the heavy hydrocarbon oil raw material is 0.1:1. the weight ratio of the light gasoline (9-120 ℃) to the heavy hydrocarbon oil raw material is 0.3:1. the weight ratio of the second superimposed product (260-290 ℃) to the heavy hydrocarbon oil raw material is 0.11:1. the weight ratio of the third superposed product (170-300 ℃) to the heavy hydrocarbon oil raw material is 0.09:1.

in the embodiment 2, only unreacted C4 fraction is introduced into a third superposed reactor to carry out superposed reaction to obtain third superposed reaction oil gas, and the third superposed reaction oil gas is further introduced into a third product separation system to be separated into unreacted recycle stream (the distillation range is-12-5 ℃) and a third superposed product (the distillation range is 170-300 ℃). The catalysts and reaction conditions of the first polymerization reactor, the second polymerization reactor and the third polymerization reactor are the same as those of example 1.

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

The main operating conditions and results are listed in table 3.

Example 3

Example 3 illustrates the effect of the process provided by the present disclosure on high selectivity propylene production during a high selectivity catalytic cracking process for high propylene production.

The reaction apparatus used was the same as in example i. The main experimental steps are the same as example l, except that the heavy hydrocarbon oil raw material is hydrogenated wax oil. The main operating conditions and results are listed in table 3.

TABLE 1-1

Tables 1 to 2

Name of catalyst	SXC-6	LXC-10
			Chemical property, weight%
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

As can be seen from table 3, in example 1, the yield of propylene prepared by the method for producing propylene in high yield by high-selectivity catalytic cracking using heavy hydrocarbon oil as a feedstock according to the method provided by the present disclosure is 26.46 wt%, which is significantly higher than that of comparative examples 1-1 and 1-2, and the olefin content of the product gasoline is only 8.06 wt%; with the same reaction conditions, the yield of propylene prepared by the cyclic catalytic conversion method of the C4 fraction and the light gasoline using the heavy hydrocarbon oil feedstock as the feed in comparative examples 1-2 was 17.28 wt%.

In comparative example 1-1, the yield of propylene prepared by the cyclic catalytic conversion method of the C4 fraction and the light gasoline using the heavy hydrocarbon oil as the feed was 21.58 wt%, but the yield of dry gas and coke was greatly increased after the reaction severity was increased. Example 1 compared to comparative example 1-1, the dry gas and coke yields decreased by 3.84 wt% and 1.37 wt%, respectively, and the propylene yield increased by 4.88 wt%.

Example 2 using the method provided by the present disclosure, the yield of propylene prepared by the highly selective catalytic cracking method for producing propylene in high yield reaches 28.99 wt% using a heavy hydrocarbon oil feedstock as a feedstock.

Example 3 using the process provided by the present disclosure, the yield of propylene produced by the high selectivity catalytic cracking process for the production of propylene was 30.96 wt% with the heavy hydrocarbon oil feedstock as the feed, and the olefin content of the product gasoline was as low as 8.53 wt%.

It should be noted by those skilled in the art that the described embodiments of the present invention are merely exemplary and that various other substitutions, alterations, and modifications may be made within the scope of the present invention. Accordingly, the present invention is not limited to the above-described embodiments, but is only limited by the claims.

Claims

1. A method for producing more propylene by high-selectivity catalytic cracking is characterized by comprising the following steps:

introducing the liquefied gas fraction subjected to C3 removal into a first superposition reactor for reaction to obtain first superposition reaction oil gas, further introducing the first superposition reaction oil gas into a first superposition product separation system for separation, and separating to obtain an unreacted C4 material flow and a first superposition product; introducing at least a part of the first superposed product into a second superposed reactor for reaction to obtain a second superposed product;

2. The process according to claim 1, wherein the first polymerization product has a distillation range of 90 to 150 ℃, and is predominantly trimethylpentene; the distillation range of the second superimposed product is 240-330 ℃, and the second superimposed product is mainly C16 olefin; the distillation range of the third superposed product is 130-330 ℃, and the third superposed product is mainly C9-C12 olefin.

3. The process according to claim 1, wherein the first polymeric product has a boiling range of 100 to 130 ℃ and is predominantly trimethylpentene; the distillation range of the second superimposed product is 260-300 ℃, and the second superimposed product is mainly C16 olefin; the distillation range of the third superposed product is 140-290 ℃, and the third superposed product is mainly C9-C12 olefin.

4. The process according to claim 1, wherein the light gasoline fraction has a boiling range of 9 to 150 ℃;

and the olefin content in the light gasoline fraction is 30 to 90 wt% based on the total weight of the light gasoline fraction.

5. The process according to claim 4, wherein the light gasoline fraction has a boiling range of 9 to 100 ℃;

the olefin content in the light gasoline fraction is 45 to 90 wt% based on the total weight of the light gasoline fraction.

6. A process according to claim 5, wherein the light gasoline fraction has a boiling range of 9 to 60 ℃.

7. The method of claim 1, wherein,

the reaction conditions of the first polymerization reaction are as follows: inverse directionThe temperature is 300-450 ℃, the pressure is 0.5-2.0 MPa, and the mass space velocity is 1-5 h ^-1 ；

The reaction conditions of the second polymerization reaction are as follows: the reaction temperature is 180-280 ℃, the pressure is 1.0-2.0 MPa, and the mass space velocity is 0.7-2.0 h ^-1 ；

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

8. The method of claim 1, wherein,

the first polymerization catalyst comprises 1-12 mass% of NiO, 45-82 mass% of amorphous aluminum silicate and 10-50 mass% of alumina;

the second polymerization catalyst comprises 1-12 mass% of NiO, 45-82 mass% of amorphous aluminum silicate and 10-50 mass% of alumina;

the third catalyst for the superposition reaction comprises 1-20 mass percent of NiO, 40-80 mass percent of HZSM-5 zeolite and 10-50 mass percent of alumina.

9. The method of claim 1, wherein,

the weight ratio of the second superposed product entering the second catalytic reactor to the heavy hydrocarbon oil raw material entering the first catalytic reactor is 0.01-0.3: 1;

the weight ratio of the third superposed product entering the second catalytic reactor to the heavy hydrocarbon oil raw material entering the first catalytic reactor is 0.01-0.4: 1.

10. the method of claim 9, wherein,

the weight ratio of the second superposed product entering the second catalytic reactor to the heavy hydrocarbon oil raw material entering the first catalytic reactor is 0.05-0.15: 1;

the weight ratio of the third superposed product entering the second catalytic reactor to the heavy hydrocarbon oil raw material entering the first catalytic reactor is 0.05 to 0.2:1.

11. the process of any one of claims 1-9, wherein the first catalytic cracking catalyst and the second catalytic cracking catalyst each contain a shape selective zeolite having an average pore diameter of less than 0.7nm, the shape selective zeolite being at least one selected from the group consisting of zeolites having an MFI structure, ferrierites, chabazites, dachiardites, erionites, a-type zeolites, epistilites, and laumontites.

12. The method of any one of claims 1-9, wherein the operating conditions of the first catalytic reaction comprise: the reaction temperature is 480 to 600 ℃; the reaction time is 0.5 to 10 seconds; the weight ratio of the agent to the oil is 5 to 15:1; the weight ratio of water to oil is 0.05 to 1:1.

13. the method of any of claims 1-9, wherein the operating conditions of the second catalytic reaction comprise: the reaction temperature is 520 to 750 ℃; the reaction time is 0.1 to 3 seconds; the weight ratio of the agent oil to the agent oil is 6 to 40:1; the weight ratio of water to oil is 0.1 to 1:1.

14. the method of any of claims 1-9, wherein the first catalytic reactor and the second catalytic reactor are each one selected from a riser reactor, a downer reactor, a fluidized bed reactor, a riser and downer combined reactor, a riser and fluidized bed combined reactor, and a downer and fluidized bed combined reactor.

15. The process of any one of claims 1-9, wherein the third catalytic reactor is a fluidized bed reactor, the operating conditions of the third catalytic reactor being: the reaction temperature is 450 to 750 ℃; the weight hourly space velocity is 1 to 30h ^-1 。

16. The method of claim 15, wherein the third catalytic reactor is a fluidized bed reactor and the operating conditions of the third catalytic reactor are: the reaction temperature is 510 to 560 ℃.

17. The method according to any one of claims 1-9, further comprising:

stripping the first and second carbon-deposited catalysts in a stripping zone of the settler,

regenerating the stripped first and second carbon-deposited catalysts in a regenerator to obtain first and second regenerated catalysts, respectively;

feeding said first regenerated catalyst to said first catalytic reactor as said first catalytic cracking catalyst and feeding said second regenerated catalyst to said second catalytic reactor as said second catalytic cracking catalyst.

18. The process of claim 17, wherein the regeneration is carried out at a temperature of 600 to 800 ℃.

19. The process of claim 17, wherein the first regenerated catalyst has a temperature of 560 to 800 ℃; the temperature of the second regenerated catalyst is 560-800 ℃.

20. The process of claim 17, wherein the first coked catalyst is stripped in a first stripping zone of the settler and the second coked catalyst is stripped in a second stripping zone of the settler.

21. The process according to any one of claims 1-9, wherein the heavy hydrocarbon oil 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.

22. The process of claim 21, wherein the heavy hydrocarbon oil feedstock is a petroleum hydrocarbon oil, which is at least one selected from the group consisting of atmospheric gas oil, vacuum gas oil, coker gas oil, deasphalted oil, hydrogenated tail oil, atmospheric residue, vacuum residue, and crude oil.

23. The method according to any one of claims 1-9, further comprising:

separating the unreacted recycle stream into a C4 hydrocarbon fraction and a light paraffin-rich gasoline fraction,

combining the light gasoline fraction rich in alkanes and the heavy gasoline fraction to obtain a gasoline product with a low olefin content.

24. A system for high selectivity catalytic cracking and high propylene yield comprises:

a reactor system, the reactor system comprising:

a first reactor having a first reactor feedstock inlet, a first reactor catalyst inlet, and a first reactor product outlet;