CN109722283B

CN109722283B - Catalytic conversion process for reducing dry gas and coke yields

Info

Publication number: CN109722283B
Application number: CN201711052530.0A
Authority: CN
Inventors: 朱金泉; 高永灿; 崔琰; 杨轶男; 谢朝钢
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: 2017-10-30
Filing date: 2017-10-30
Publication date: 2021-06-11
Anticipated expiration: 2037-10-30
Also published as: CN109722283A

Abstract

The present disclosure relates to a catalytic conversion process for reducing dry gas and coke yields, the process comprising: carrying out a first catalytic reaction on a heavy hydrocarbon oil raw material in a first reactor to obtain a first carbon deposition catalyst and a first product; carrying out a second catalytic reaction on the light hydrocarbon in a second reactor to obtain a second carbon deposition catalyst and a second product; introducing the first and second carbon-deposited catalysts into a stripper for stripping, and introducing a first part of the stripped catalyst into an oxygen-containing fluidized zone of a regenerator for aerobic regeneration; introducing the second part of the stripped catalyst into an anaerobic fluidization area of a regenerator for anaerobic regeneration at 600-800 ℃. Compared with the prior art, the method can produce more low-carbon olefins in the catalytic conversion of heavy hydrocarbon oil, and simultaneously has lower dry gas and coke yield.

Description

Catalytic conversion process for reducing dry gas and coke yields

Technical Field

The present disclosure relates to a catalytic conversion process for reducing dry gas and coke yields.

Background

Modern petroleum processing technology focuses more on reducing unit feedstock processing energy consumption and reducing carbon emissions while pursuing high yields of high value products (e.g., ethylene, propylene, C8 aromatics). Hydrogen, methane and coke which are byproducts produced by petroleum hydrocarbon processing in modern refineries are usually removed by combustion and are main sources of carbon emission of refineries, so that the economical efficiency of atomic processing is greatly reduced.

Patent CN102071054A discloses a catalytic cracking method, which comprises the steps of contacting a heavy raw material with a catalyst containing zeolite with an average pore size of less than 0.7nm in a first riser reactor to perform a cracking reaction, and contacting a light raw material and cracked heavy oil with a catalyst containing zeolite with an average pore size of less than 0.7nm in a second riser reactor and a fluidized bed reactor to perform a cracking reaction. The method is used for heavy oil catalytic cracking, the heavy oil conversion rate and the propylene yield are high, and the dry gas and coke yield is low.

Patent CN102690681A discloses a catalytic cracking method for producing propylene, which comprises the steps of contacting a heavy raw material with a catalyst containing shape-selective zeolite with an average pore diameter of less than 0.7nm in a first riser reactor for a cracking reaction, and contacting a light hydrocarbon with a catalyst containing shape-selective zeolite with an average pore diameter of less than 0.7nm at a temperature of 550-690 ℃ in a second riser reactor and a fluidized bed reactor for a cracking reaction. The method is used for heavy oil catalytic cracking, the heavy oil conversion rate and the propylene yield are high, and the dry gas and coke yield is low.

Patent US6106697 discloses a method for selectively producing C2-C4 olefins by using wax oil or residual oil as raw material and adopting a two-stage reactor to carry out catalytic cracking reaction. Contacting a wax oil or residual oil raw material with a large-pore zeolite catalyst in a first-stage reactor under conventional catalytic cracking conditions to perform catalytic cracking reaction to generate different boiling range products including gasoline fractions; and (3) feeding the gasoline fraction generated by the first-stage reactor into a second-stage reactor, wherein the reaction temperature is 500-650 ℃, and the solvent-oil ratio is 4-10: 1. the hydrocarbon partial pressure is 70-280 kPa, and the catalyst is contacted with a medium-pore zeolite catalyst for further reaction to generate C2-C4 olefin.

From the technology of increasing the yield of low-carbon olefins by catalytic cracking, the catalyst mainly has two functions in the catalytic cracking reaction process, one is that the temperature of the catalyst is increased after coke is burnt out in a regenerator, the energy carried by the high-temperature catalyst provides reaction heat for the reaction of hydrocarbon oil raw materials, and the other is that shape-selective zeolite or mesoporous zeolite in the catalyst provides a proper catalytic environment for the reaction of light raw materials. However, the catalyst contains, in addition to the shape-selective zeolite or the mesoporous zeolite, other active components which tend to coke and which, when brought into contact with the light feedstock, cause it to form a part of coke and dry gas.

In the prior art, after the carbon-deposited catalyst is completely introduced into a regenerator and coke is completely burned off, the high-temperature regenerated catalyst returns to the reactor for recycling, but when the coke-burned off catalyst is in contact reaction with a light hydrocarbon raw material, coke is generated on the catalyst and dry gas is generated. Therefore, the prior art has the problem that during the recycling process of the catalyst between the regenerator and the reactor, a large amount of coke and dry gas are repeatedly generated when light hydrocarbons are in contact reaction with the catalyst.

Disclosure of Invention

It is an object of the present disclosure to provide a catalytic conversion process that reduces the yield of dry gas and coke.

To achieve the above objects, the present disclosure provides a catalytic conversion method for reducing yields of dry gas and coke, the method comprising:

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

enabling light hydrocarbons to be in contact with a second catalytic cracking catalyst in a second reactor to carry out a second catalytic reaction, introducing a reaction mixture containing the catalyst and obtained by the second reactor into a third reactor to continue a third catalytic reaction, and carrying out gas-agent separation on the reaction mixture obtained by the third reactor in a settler to obtain a second carbon deposition catalyst and a second product, wherein the distillation range of the light hydrocarbons is 8-253 ℃;

introducing the first and second coked catalysts into a stripper disposed below the settler for stripping, introducing a first portion of the stripped catalyst into an oxygen-containing fluidized zone of a regenerator for aerobic regeneration to obtain a first regenerated catalyst, and feeding the first regenerated catalyst into a first reactor as the first catalytic cracking catalyst;

introducing a second part of the stripped catalyst into an anaerobic fluidization area of a regenerator for anaerobic regeneration at 600-800 ℃ to obtain a second regenerated catalyst, and feeding the second regenerated catalyst serving as the second catalytic cracking catalyst into a second reactor;

and introducing the first product and the second product into a product separation system for product separation to obtain separated products including a low-carbon olefin product, a gasoline product, a diesel oil product and a heavy oil product.

Optionally, the method further comprises: introducing a fluidizing gas having an oxygen content of less than 3% by volume into the oxygen-free fluidizing zone, the fluidizing gas containing not less than 97% by volume of an inert gas, the inert gas being at least one selected from the group consisting of nitrogen, helium, neon, argon, krypton, and xenon.

Optionally, the oxygen-free fluidizing zone of the regenerator comprises: the system comprises a low-temperature catalyst conveying pipe, a gas-solid separator, a catalyst distributor, a heat taking exhaust pipe, a degassing pipe and a high-temperature catalyst conveying pipe;

one end of the low-temperature catalyst conveying pipe is communicated with the stripper, the other end of the low-temperature catalyst conveying pipe is connected with a gas-solid separator, a catalyst distributor and a heat taking exhaust pipe communicated with the catalyst distributor are arranged below the gas-solid separator, the heat taking exhaust pipe is a plurality of groups of pipelines which are arranged in the vertical direction and are communicated with the bottom ends of the pipelines, the degassing pipe is connected to the lower end of the heat taking exhaust pipe and is provided with a fluidized gas inlet at the joint, the degassing pipe is a vertical pipe, the lower end of the degassing pipe is connected with the high-temperature catalyst conveying pipe, and the high-temperature catalyst conveying pipe is communicated with.

Optionally, 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.

Optionally, the temperature of the second regenerated catalyst is 560-800 ℃, and the carbon deposit content of the second regenerated catalyst is 0.5-1.9 wt%, preferably 0.9-1.3 wt%, and more preferably 0.91-0.99 wt% based on the dry weight of the second regenerated catalyst.

Optionally, the hydrogen content in the carbon deposit of the second regenerated catalyst is 0.1 to 0.65 wt%, preferably 0.1 to 0.5 wt%, based on the weight of the carbon deposit in the second regenerated catalyst.

Optionally, 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.

Optionally, the heavy hydrocarbon oil feedstock has an average relative molecular mass of no less than 200.

Optionally, the light hydrocarbon is 9-160 ℃, and further preferably 9-60 ℃;

the light hydrocarbon has an olefin content of 30 to 90 wt%, preferably 45 to 90 wt%, for example 45 to 80 wt%, or 45 to 65 wt%, or 55 to 75 wt%, based on the total weight of the light hydrocarbon.

Optionally, the light hydrocarbons are at least partially derived from the separation products of the product separation system. In this case, the separation system further separates a light hydrocarbon separation product, and the weight ratio of the light hydrocarbon to the heavy hydrocarbon oil raw material is (0.01-0.6): 1, preferably (0.05-0.3): 1.

optionally, the first and second catalytic cracking catalysts 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 the MFI structure, ferrierites, chabazites, dachiardites, erionites, a-type zeolites, epistillomites, and nephelozeolites.

Optionally, the operating conditions of the first catalytic reaction include: the reaction temperature is 480 to 600 ℃, for example 500 to 550 ℃ or 510 to 540 ℃; the reaction time is 0.5 to 10 seconds, for example, 1 to 5 seconds or 2 to 4 seconds; the weight ratio of the agent oil is (5-15): 1, for example, 6 to 10: 1; the weight ratio of water to oil is (0.05-1): 1 is, for example, (0.08-0.5): 1 or (0.1-0.3): 1.

optionally, the operating conditions of the second catalytic reaction include: the reaction temperature is 520-750 ℃, for example, 520-600 ℃ or 520-550 ℃; the reaction time is 0.1 to 3 seconds, for example, 1 to 3 seconds or 1.4 to 3 seconds; the weight ratio of the agent to the oil is (6-40): 1 is, for example, (8-25): 1 or (10-20): 1 or (10-17): 1; the weight ratio of water to oil is (0.1-1): 1 is, for example, (0.1 to 0.5): 1 or (0.1-0.25): 1 or (0.1-0.3): 1.

optionally, the first reactor and the second 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.

Optionally, the third reactor is a fluidized bed reactor, and the operating conditions of the third reactor are: the reaction temperature is 450 to 750 ℃, for example, 500 to 600 ℃ or 510 to 550 ℃, preferably 510 to 560 ℃; the weight hourly space velocity is 1-30 h^-1For example, 3 to 20 or 5 to 15 hours^-1(ii) a The absolute pressure of the settler is 0.15-0.40 MPa.

The catalytic cracking method provided by the disclosure overcomes the difficult problems that two processes of catalyst heating and catalyst coking are strictly bound and cannot be separated in the existing process route and regenerator structure design, and the high-temperature regenerated catalyst circulation-heavy hydrocarbon oil reaction line and the high-temperature carbon deposition catalyst circulation-light hydrocarbon oil reaction line are respectively configured, so that the yield of dry gas and coke can be reduced. By establishing a special regeneration reaction process of the carbon-deposited catalyst, the key problems of high coke yield and low light hydrocarbon conversion rate caused by the fact that olefin generates coke when the carbon-deposited catalyst directly reacts with light hydrocarbon oil in the prior art are solved. Compared with the prior art, the method can produce more low-carbon olefins in the catalytic conversion of heavy hydrocarbon oil, and simultaneously has lower dry gas and coke yield.

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 in any one or more, preferably all of the following effects:

1. the regenerator of the prior art has two functions: firstly, the carbon deposit on the catalyst is burnt out in an oxygen-containing environment, and secondly, the temperature of the catalyst is raised to carry the heat required by the reaction, and the two functions are bound, so that the regeneration process can only provide the high-temperature catalyst for completely burning out the carbon deposit. The oxygen-containing fluidization area and the oxygen-free fluidization area are arranged in the regenerator, the oxygen-containing fluidization area still has two functions of burning off carbon deposit on the catalyst and heating the catalyst, the high-temperature catalyst with the carbon deposit completely burned off is provided, the oxygen-free fluidization area maintains the function of heating the catalyst, a new oxygen-free regeneration reaction process is established, the carbon in the carbon deposit is not burned off, the mass fraction of hydrogen in the carbon deposit is only greatly reduced, the property of the carbon deposit on the catalyst is changed, and the carbon deposit catalyst with the high-temperature specific reaction performance is provided. The regenerator in the disclosure realizes the relative independence of two functions of catalyst temperature rise and catalyst carbon deposit burning-off, thereby being capable of flexibly providing various catalysts for the reaction process.

2. In the prior patent, a to-be-generated carbon deposition catalyst is introduced to participate in a petroleum hydrocarbon reaction, so that the effect of reducing the yield of dry gas and coke is achieved, but the effect of reducing the yield of dry gas and coke is general, the yield of high-added-value products such as propylene is reduced, the mass fraction of hydrogen of carbon deposition on the untreated carbon deposition catalyst is high and is between 0.55 and 0.85 percent, the carbon deposition has chemical reaction activity, and is easy to generate hydrogen transfer activity with light hydrocarbons such as olefin, so that the cracking reaction of the olefin is hindered, even the olefin is generated into coke, a negative effect is achieved, the yield of the coke is high, and the conversion rate is low. The method thoroughly changes the properties of carbon deposit on the catalyst through the reaction regeneration process of the anaerobic fluidized zone, the hydrogen mass fraction of the carbon deposit is obviously reduced, the treated carbon deposit catalyst returns to the reactor through a second catalytic cracking catalyst and reacts with light hydrocarbons including olefin, the inhibition effect on olefin cracking is also greatly reduced, and the selectivity of converting the olefin into the low-carbon olefin can be improved. The light hydrocarbon containing olefin will have more cracking reaction to produce propylene and other low carbon olefin and reduce coke and dry gas.

3. The method provided by the disclosure reacts the light hydrocarbon with the carbon-deposited catalyst with high temperature and specific reaction performance, so that the heat required by the light hydrocarbon reaction is ensured, and the yield of dry gas and coke during the light hydrocarbon reaction is greatly reduced.

4. The design of the regenerator and the catalyst circulation circuit in the disclosure can flexibly adjust the initial temperature of the carbon-deposited catalyst entering the light hydrocarbon reaction zone, so that the adjustment of the catalyst-oil ratio in the reactor is not limited by the reaction temperature requirement of the reactor outlet, and the problem of thermal cracking reaction caused by direct contact of light hydrocarbons and a high-temperature regenerated catalyst can also be reduced.

5. The anaerobic fluidization area in the regenerator takes away part of heat in the regenerator by heating the carbon-deposited catalyst, reduces the temperature of the regenerator, reduces the occurrence of secondary combustion (also called tail combustion) in a dilute phase space of the regenerator, has partial functions of a regenerator heat-taking device, even can cancel the regenerator heat-taking device, and saves investment.

6. The process provided by the present disclosure has low coke selectivity, low dry gas selectivity, and higher propylene yield than existing processes.

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

Drawings

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

FIG. 1 is a schematic flow diagram of a catalytic cracking process for reducing dry gas and coke yields provided by the present disclosure.

FIG. 2 is a schematic diagram of the structure of an oxygen-free fluidizing zone in a regenerator.

Description of the reference numerals

1 first reactor

2 second reactor

3 third reactor

7 stripper

8 settler

9 regenerator

11 first catalytic cracking catalyst transfer line

12 second catalytic cracking catalyst transfer line

17 spent catalyst transfer line

18 spent catalyst conveying pipeline

19 spent catalyst transfer line

21 line (injecting heavy hydrocarbon oil feedstock)

22 line (injecting light hydrocarbons)

28 pipeline (transportation reaction oil gas)

30 pipeline (conveying dry gas)

31 pipeline (conveying liquefied gas)

32 pipeline (transportation gasoline)

33 pipeline (transportation diesel)

34 pipeline (for heavy oil)

41 pipeline (injecting atomized steam)

42 pipeline (injecting atomized steam)

47 line (steam stripping injection)

51 line (injection pre-lifting medium)

52 line (injection pre-lifting medium)

53 pipeline (injection lifting medium)

58 line (air injection)

59 pipeline (conveying regeneration smoke)

70 product separation system

91 regenerator oxygen containing fluidizing zone

92 regenerator oxygen-free fluidized zone

921 low temperature catalyst transfer pipe

922 gas-solid separator

923 catalyst distributor

924 heat-taking calandria

925 degassing pipe

926 high-temperature catalyst conveying pipe

927 fluidizing gas inlet

Detailed Description

The following detailed description of specific embodiments of the present disclosure is provided in connection with the accompanying drawings. It should be understood that the detailed description and specific examples, while indicating the present disclosure, are given by way of illustration and explanation only, not limitation.

The present disclosure provides a catalytic conversion process for reducing dry gas and coke yields, the process comprising:

introducing the first carbon-deposited catalyst and the second carbon-deposited catalyst into a stripper below a settler for stripping, introducing a first part of the stripped catalyst into an oxygen-containing fluidized zone of a regenerator for aerobic regeneration to obtain a first regenerated catalyst, and feeding the first regenerated catalyst into a first reactor as the first catalytic cracking catalyst;

and introducing the first product and the second product into a product separation system for product separation to obtain separated products comprising a low-carbon olefin product, a gasoline product, a heavy oil product and an optional light hydrocarbon product.

In the existing catalytic cracking device, after a catalytic cracking catalyst is regenerated by a regenerator, the temperature of the catalyst is raised to 600-700 ℃, and then the catalyst reacts with heavy hydrocarbon oil, the temperature of a carbon deposition catalyst is lowered to 480-600 ℃, at the moment, the heat carried by the carbon deposition catalyst is insufficient to provide sufficient reaction heat for the light hydrocarbon to react and crack to produce more low-carbon olefins, so that the carbon deposition catalyst must return to the regenerator to be regenerated to raise the temperature, and carbon deposition on the upper surface of the catalyst is burnt in an oxygen-containing environment in the regenerator; the regenerated catalyst burnt with clean coke reacts with light hydrocarbons, and more carbon deposits and more dry gas are formed on the catalyst, so that the selectivity of the product is reduced. In the prior art and the device, the carbon deposition catalyst with activity is unnecessarily burnt due to the requirement of process reaction heat, and the regenerated catalyst has to be selected to directly react with light hydrocarbon, so that the yield of coke and dry gas in the final product is high.

On the other hand, the inventors of the present disclosure have found that, if the carbon deposition catalyst reacted with the heavy hydrocarbon oil is directly reacted with the light hydrocarbon, even if the carbon deposition catalyst is stripped, the carbon deposition composition still contains a large amount of polycyclic aromatic hydrocarbons with alkyl side chains, the mass fraction of hydrogen in the carbon deposition is high, and the cracking reaction of the light hydrocarbon is easily hindered by the hydrogen transfer reaction during the reaction of the carbon deposition catalyst with the light hydrocarbon, and new coke is easily formed.

The inventor of the present disclosure finds, through laboratory studies, that after a catalytic cracking catalyst containing a shape-selective molecular sieve reacts with heavy hydrocarbon oil under appropriate conditions and is subjected to steam stripping, a carbon deposit catalyst is formed, and the carbon deposit catalyst is heated and subjected to anaerobic regeneration treatment to obtain an anaerobic regenerated catalyst. Compared with the reaction of light hydrocarbon on a regenerated catalyst of which coke is completely burned out, the method has the advantages that the light hydrocarbon continuously and repeatedly enters oil to react on the carbon-deposited catalyst after the treatment, the yield of the low-carbon olefin in the obtained product is not reduced, and the yields of dry gas and coke are greatly reduced.

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 dry weight (weight calcined at 800 ℃ for 1 hour) of the catalytic cracking catalyst, 15 to 50 wt% of clay on a dry basis, 15 to 50 wt% of molecular sieve on a dry basis, and 10 to 35 wt% of binder on a dry basis, wherein the molecular sieve is a zeolite of MFI structure or consists of 25 to 100 wt% of zeolite of MFI structure and 0 to 75 wt% of other zeolites except for zeolite of MFI structure; 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. Such as one or more of Y-type zeolite, beta zeolite. 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 reactor and the second reactor may be catalytic conversion reactors well known to those skilled in the art, for example, the first reactor and the second 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, a heavy hydrocarbon oil feedstock is contacted with a first catalytic cracking catalyst in a fluidized state in a first reactor to carry out 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 550 ℃ or 510 to 540 ℃; the reaction time is 0.5 to 10 seconds, for example, 1 to 5 seconds or 2 to 4 seconds; the weight ratio of the agent oil is (5-15): 1, for example, 6 to 10: 1; the weight ratio of water to oil is (0.05-1): 1 is, for example, (0.08-0.5): 1 or (0.1-0.3): 1.

according to the present disclosure, the heavy hydrocarbon oil feedstock may be at least one selected from the group consisting of petroleum hydrocarbon oils, synthetic oils, coal liquefaction oils, oil sand oils, and shale oils. The synthetic oil can be distillate oil obtained by Fischer-Tropsch (F-T) synthesis of coal and natural gas. Preferably, the heavy hydrocarbon oil feedstock is a petroleum hydrocarbon oil, such as 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.

In accordance with the present disclosure, the average relative molecular mass of the heavy hydrocarbon oil feedstock is preferably not less than 200 in order to achieve optimal carbon deposition on the catalyst surface. The heavy hydrocarbon oil raw material with the relative molecular mass can ensure enough carbon deposit content, and meanwhile, the macromolecular heavy hydrocarbon oil raw material cannot enter the micropores of the catalyst, so that the carbon deposit is deposited on the surface and the large and medium pores of the catalyst, and the secondary reaction effect of the carbon deposit catalyst is not influenced. The average relative molecular weight refers to the average relative molecular weight of the hydrocarbons, namely the sum of the relative atomic masses of all atoms in the chemical formula of the petroleum hydrocarbons, and is a basic physical parameter of the petroleum hydrocarbons, and the industry determination method comprises the following steps: hydrocarbon relative molecular weight determination SH/T0583.

According to the present disclosure, the light hydrocarbons are contacted with a second catalytic cracking catalyst in a second reactor under fluidized conditions for a second catalytic reaction. The operating conditions of the second catalytic reaction may include: the reaction temperature is 520-750 ℃, for example, 520-600 ℃ or 520-550 ℃; the reaction time is 0.1 to 3 seconds, for example, 1 to 3 seconds or 1.4 to 3 seconds; the weight ratio of the agent to the oil is (6-40): 1 is, for example, (8-25): 1 or (10-20): 1 or (10-17): 1; the weight ratio of water to oil is (0.1-1): 1 is, for example, (0.1 to 0.5): 1 or (0.1-0.25): 1 or (0.1-0.3): 1.

according to the present disclosure, the specific type of the light hydrocarbon is not particularly limited as long as the distillation range is 8-253 ℃, and the light hydrocarbon may include, for example, a C4 hydrocarbon fraction, straight-run gasoline, catalytically cracked naphtha, catalytically cracked stabilized gasoline, coker gasoline, fischer-tropsch (F-T) synthetic gasoline, various diesel light fractions, and the like, and the fractions may be obtained from the apparatuses described in the present disclosure, or may be obtained from other apparatuses. In order to further improve the yield of high-quality gasoline, the distillation range of the light hydrocarbon is preferably 9-160 ℃, and more preferably 9-60 ℃. The light hydrocarbon may have an olefin content of 30 to 90 wt%, preferably 45 to 90 wt%, for example, 45 to 80 wt%, or 45 to 65 wt%, or 55 to 75 wt%, based on the total weight of the light hydrocarbon.

According to the present disclosure, the light hydrocarbons may be at least partially derived from the separation product of the product separation system, in this case, the separation system further separates a light hydrocarbon separation product, and the weight ratio of the light hydrocarbons to the heavy hydrocarbon oil feedstock may be (0.01-0.6): 1, preferably (0.05-0.3): 1.

in accordance with the present disclosure, the third 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. The operating conditions of the third reactor may be: the reaction temperature is 450 to 750 ℃, for example, 500 to 600 ℃ or 510 to 550 ℃, preferably 510 to 560 ℃; the weight hourly space velocity is 1-30 h^-1For example, 3 to 20 or 5 to 15 hours^-1。

According to the disclosure, the mixture obtained from the reaction in the first reactor is subjected to gas-agent separation in a settler, and the obtained first carbon-deposited catalyst is introduced into a stripper below the settler for stripping; and the reaction mixture obtained in the third reactor is subjected to gas-agent separation in the settler, and the obtained second carbon-deposited catalyst is also introduced into a stripper for steam stripping; the absolute pressure in the settler may be 0.1-0.40 MPa.

In accordance with the present disclosure, a first portion of the stripped catalyst is introduced into an oxygen-containing fluidized zone of a regenerator for aerobic regeneration to yield a first regenerated catalyst. The first regenerated catalyst may have a temperature of 560 to 800 ℃, for example 620 to 720 ℃, and the first regenerated catalyst may have a carbon deposit content of 0.01 to 0.1 wt% based on the dry weight of the first regenerated catalyst.

According to the present disclosure, the second portion of the stripped catalyst is introduced into an anaerobic fluidized zone of a regenerator for anaerobic regeneration, so that the carbon deposit on the carbon deposit catalyst of the second portion is accelerated to perform anaerobic regeneration, and then a second regenerated catalyst is obtained. The temperature of the second regenerated catalyst may be 560 to 800 ℃, for example 600 to 780 ℃, or 620 to 720 ℃, or 650 to 710 ℃, or 630 to 700 ℃, and the carbon deposit content of the second regenerated catalyst may be 0.5 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. Further, the hydrogen content in the carbon deposit of the second regenerated catalyst may be 0.1 to 0.65% by weight, preferably 0.1 to 0.5% by weight, based on the weight of the carbon deposit in the second regenerated catalyst.

According to the present disclosure, the method may further comprise: a fluidizing gas having an oxygen content of less than 3% by volume (under standard conditions) is introduced into the oxygen-free fluidizing zone, which comprises an inert gas, which may be, for example, nitrogen, as well as the noble gases helium, neon, argon, krypton, xenon, etc. For example, the fluidizing gas contains 0-3 vol% or 0.5-2.5 vol% of oxygen gas, 97-100 vol% or 97.5-99.5 vol% of inert gas under standard conditions. The introduction of the fluidizing gas containing a small amount of oxygen can burn hydrogen in the catalyst carbon deposit, further reduce the mass fraction of the hydrogen in the carbon deposit, accelerate the graphitization process of the coke by converting polycyclic aromatic hydrocarbon in the carbon deposit into graphite, and quickly increase the temperature of the catalyst.

In accordance with the present disclosure, the structure of the oxygen-free fluidizing zone can be of any form capable of accomplishing the objects of the present disclosure. In an alternative embodiment of the present disclosure, as shown in fig. 2, the oxygen-free fluidizing zone of the regenerator may comprise: a low temperature catalyst delivery pipe 921, a gas-solid separator 922, a catalyst distributor 923, a heat extraction calandria 924, a degassing pipe 925, and a high temperature catalyst delivery pipe 926. One end of the low-temperature catalyst delivery pipe 921 is communicated with the stripper 7, and the other end is connected with a gas-solid separator 922. The gas-solid separator 922 is provided with a catalyst distributor 923 and a heat-taking calandria 924 communicated with the catalyst distributor 923. Get hot calandria 924 and be the pipeline that the multiunit set up and bottom intercommunication along the vertical direction, catalyst distributor 923 have with get a plurality of distribution openings that hot calandria 924's multiunit pipeline is linked together, the catalyst gets into receive behind the catalyst distributor 923 the action of gravity by these a plurality of distribution openings evenly distributed get in the multiunit pipeline of hot calandria 924. The degassing pipe 925 is connected to the lower end of the heat taking calandria 924 and is provided with a fluidizing gas inlet 927 at the connection position, the degassing pipe 925 is a vertical pipe with the lower end connected with the high-temperature catalyst delivery pipe 926, and the high-temperature catalyst delivery pipe 926 is communicated with the second reactor 2. The temperature (600-800 ℃) of the anaerobic regeneration in the present disclosure refers to the temperature in the degassing tube 925.

According to the disclosure, the first product and the second product are introduced into a product separation system for product separation, so as to obtain separated products including a dry gas product, a liquefied gas product, a gasoline product, a diesel oil product, a heavy oil product and the like, wherein the liquefied gas is further separated to obtain target products such as ethylene, propylene and the like.

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 reactor 1 is a riser reactor, a second reactor 2 is a riser reactor, and a third reactor 3 is a fluidized bed reactorAnd (4) applying the device. The first catalytic cracking catalyst (hot regenerated catalyst) enters the bottom of the riser 1 of the first reactor from the regenerator oxygen-containing fluidized zone 91 via 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 the riser 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 8 is 0.1-0.40 MPa. The mixture of reaction oil gas and catalyst in the riser 1 is separated by the fast separating device gas agent at the outlet, the first carbon-deposited catalyst is introduced into the stripper 7, the separated reaction oil gas is sent into the subsequent product separating system 70 for product separation through the settler 8 and the pipeline 28 at the top thereof, and products such as dry gas, liquefied gas, gasoline, diesel oil, heavy oil and the like are obtained after separation (respectively led out through the

pipelines

30, 31, 32, 33 and 34), wherein the liquefied gas is further separated to obtain target products such as ethylene, propylene and the like, and simultaneously, light hydrocarbons for recycling can be further separated and obtained (led out through the pipeline 22). The light hydrocarbon is mixed with the atomized steam from the pipeline 42 through the pipeline 22 and then injected into the second reactor 2, and the weight ratio of the water steam to the hydrocarbon oil raw material is (0.1-1): 1, the outlet temperature of the riser reactor 2 is 520-750 ℃, the reaction time in the riser 2 is 0.1-3 seconds, and the weight ratio of the catalyst to the hydrocarbon oil raw material is 6-40. The mixture of the reaction oil gas and the catalyst in the riser reactor 2 is further introduced into a fluidized bed in a third reactor 3 through a riser outlet to continue reacting, the reaction temperature of the fluidized bed 3 is 450-750 ℃, and the weight hourly space velocity is 1-30 h^-1After the reaction, the oil gas and a part of the carbon-deposited spent catalyst enter a settler 8 through the fluidized bed reactor 3 for separation, and the second carbon-deposited catalyst is separated and enters a stripper 7. The stripping steam is injected into the stripper 7 through a pipeline 47 and contacts with the coked spent catalyst in a countercurrent manner, and the reaction oil gas carried by the spent catalyst is stripped as clean as possible. The regenerator 9 comprises an oxygen-containing fluidizing zone 91 and an oxygen-free fluidizing zone 92. The stripped carbon-deposited catalyst is introduced into a spent catalyst conveying pipeline17, part of the coked catalyst is further transported to an oxygen-free fluidizing zone 92 in the regenerator 9 via spent catalyst transport line 19 under the action of a lift medium injected via line 53, firstly, a small amount of gas carried by a gas-solid separator at the top of the anaerobic fluidization area is separated from the upper part, the gas is introduced into a plurality of heat taking discharge pipes of the anaerobic fluidization area through a catalyst distributor under the action of gravity, and a part of the fluidization gas with the oxygen content of less than 3 percent by volume is introduced into the anaerobic fluidization area, reacting and heating in a regenerator under high temperature environment to accelerate the carbon deposition on the carbon deposition catalyst to generate a coke graphitization process, the hydrogen mass fraction of the carbon deposit is obviously reduced, the carbon deposit catalyst from the heat-taking calandria finally enters the degassing pipe to remove a small amount of entrained gas, the reaction process is completed, and a second regenerated catalyst is obtained and is returned to the second reactor 2 for recycling as a second catalytic cracking catalyst through a second catalytic cracking catalyst conveying pipeline 12. The rest of the carbon deposited catalyst is sent to the oxygen-containing fluidization area 91 of the regenerator 9 through the spent catalyst conveying line 18, air is injected into the oxygen-containing fluidization area 91 through the line 58, the catalyst is contacted with heated air in the regenerator and is regenerated at the temperature of 600 ℃ -800 ℃ (degassing pipe temperature), so that a first regenerated catalyst is obtained, and the first regenerated catalyst is used as a first catalytic cracking catalyst and is returned to the first reactor for recycling through the first catalytic cracking catalyst conveying line 11. Regeneration flue gas is withdrawn via line 59.

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

The catalyst used in the examples is a cracking catalyst produced by the Chinese petrochemical catalyst, Qilu division, having a trade mark of MMC-2, and the specific properties are shown in Table 1, and the catalyst contains shape selective zeolite with 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 reducing dry gas and coke yields during catalytic conversion of hydrocarbon oils.

The experiment was carried out using a medium-sized apparatus for continuous reaction-regeneration operation having three reactors, wherein the first reactor was a riser, and the riser reactor had an inner diameter of 16 mm and a height of 3800 mm. The second reactor was a riser reactor with an internal diameter of 16 mm and a height of 3200 mm. The outlet of the second riser reactor was introduced into a third reactor, which had an internal diameter of 64 mm and a height of 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 reactor through a first catalytic cracking catalyst inclined tube, and flows upwards under the action of pre-lifting steam. Heavy hydrocarbon oil feedstock (see table 2 for main properties) is heated to 350 ℃ in a preheating furnace, mixed with atomized water vapor, sprayed into a first reactor through a feed nozzle, and contacted with a hot first regenerated catalyst to perform catalytic conversion reaction. The reaction oil gas (first product) and the first carbon-deposited catalyst enter a settler from the outlet of a riser of the first reactor for rapid separation, and the first carbon-deposited catalyst enters a stripper for stripping. The second catalytic cracking catalyst is a treated graphitized carbon-deposited catalyst (second regenerated catalyst), the temperature of the second catalytic cracking catalyst is 680 ℃, the second catalytic cracking catalyst enters the bottom of the riser reactor of the second reactor through the second catalytic cracking catalyst inclined tube and flows upwards under the action of pre-lifting steam. And (3) allowing the light hydrocarbon raw material (9-57 ℃) to enter a second reactor to contact with a hot second regenerated catalyst for catalytic reaction. The weight ratio of the light hydrocarbon to the heavy hydrocarbon oil raw material is 0.2: 1. introducing the reaction oil mixture from the second reactor into a fluidized bed of a third 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 stripper for stripping under the action of gravity. The reaction oil gas from the first reactor and the third reactor is led out from the settler and introduced into the product separating system for product separation to obtain gas product and various liquid products, and partial reaction oil gas is separated to obtain light hydrocarbon. After the carbon deposition catalyst of the stripper is stripped, one part of the carbon deposition catalyst enters an anaerobic fluidization region of a regenerator through a spent agent conveying pipe, and one part of fluidizing gas (high-purity nitrogen is selected in a laboratory, the nitrogen content of the high-purity nitrogen is more than or equal to 99.99 volume percent and the oxygen content of the high-purity nitrogen is less than or equal to 0.01 volume percent under the standard condition) is introduced into the anaerobic fluidization region to carry out anaerobic regeneration at 695 ℃, the carbon deposition accelerates the occurrence of a coke graphitization process, meanwhile, the mass fraction of hydrogen of the carbon deposition is obviously reduced, and the obtained second regenerated catalyst is returned to a second reactor for recycling; the rest part enters into an oxygen-containing fluidization area of the regenerator through a spent agent conveying pipe to be contacted with heated air and is subjected to aerobic regeneration at 700 ℃, and the first regenerated catalyst which obtains heat is returned to the first reactor for recycling. The carbon deposit content of the first catalytic cracking catalyst (first regenerated catalyst), and the carbon deposit content and the hydrogen mass fraction of the carbon deposit of the second catalytic cracking catalyst (second regenerated catalyst) were detected during continuous operation. The main operating conditions and results are listed in table 3.

Comparative example l

Comparative example l illustrates the effect of the hydrocarbon oil catalytic conversion process when the regenerator has only an oxygen-containing fluidized zone and the second catalytic cracking catalyst is a 680 c regenerated catalyst.

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 reactor through a first catalytic cracking catalyst inclined tube, and flows upwards under the action of pre-lifting steam. Heavy hydrocarbon oil feedstock (see table 2 for main properties) is heated to 350 ℃ in a preheating furnace, mixed with atomized water vapor, sprayed into a first reactor through a feed nozzle, and contacted with a hot first regenerated catalyst to perform catalytic conversion reaction. The reaction oil gas (first product) and the first carbon-deposited catalyst enter a settler from the outlet of a riser of the first reactor for rapid separation, and the first carbon-deposited catalyst enters a stripper for stripping. The second catalytic cracking catalyst is a second regenerated catalyst with the temperature of 680 ℃, enters the bottom of the riser reactor of the second reactor through a second catalytic cracking catalyst inclined tube, and flows upwards under the action of pre-lifting steam. And (3) allowing the light hydrocarbon raw material (9-57 ℃) to enter a second reactor to contact with a hot second regenerated catalyst for catalytic conversion reaction. The weight ratio of the light hydrocarbon to the heavy hydrocarbon oil raw material is 0.2: 1. introducing the reaction oil mixture from the second reactor into a fluidized bed of a third reactor at the outlet of a riser for further reaction, separating oil gas (a second product) and a second carbon-deposited catalyst in a settler after the reaction, and allowing the separated second carbon-deposited catalyst to enter a stripper for stripping under the action of gravity. The reaction oil gas from the first reactor and the third reactor is led out from the settler and introduced into the product separating system for product separation to obtain gas product and various liquid products, and partial reaction oil gas is separated to obtain light hydrocarbon. The carbon deposited catalyst in the stripper is stripped and then is totally returned to the oxygen-containing fluidization area of the regenerator through a spent agent conveying pipe, is contacted with heated air and is subjected to aerobic regeneration at 700 ℃, and the obtained hot regenerated catalyst is returned to the first reactor and the second reactor for recycling. The regenerator is additionally provided with an external heat taking device to take away surplus heat, and the reaction heat balance and the regeneration temperature of the regenerator are controlled. The carbon deposit content of the first catalytic cracking catalyst (first regenerated catalyst), and the carbon deposit content and the hydrogen mass fraction of the carbon deposit of the second catalytic cracking catalyst (second regenerated catalyst) were detected during continuous operation. The main operating conditions and results are listed in table 3. The main operating conditions and results are listed in table 3.

Comparative example 2

Comparative example 2 illustrates the effect of the catalytic conversion process of hydrocarbon oils with a regenerator having only an oxygen-containing fluidized zone, the second catalytic cracking catalyst being a coked catalyst directly from the stripper, which was rapidly heated to 680 ℃ by electrical heating.

The first catalytic cracking catalyst is regenerated catalyst with temperature of 680 ℃, enters the bottom of the riser reactor of the first reactor through the first catalytic cracking catalyst inclined tube, and flows upwards under the action of pre-lifting steam. Heavy hydrocarbon oil raw materials (main properties are shown in table 2) are heated to 350 ℃ by a preheating furnace, mixed with atomized water vapor, sprayed into a first reactor through a feeding nozzle, and contacted with a hot regenerated catalyst to carry out catalytic conversion reaction. The reaction oil gas (first product) and the first carbon-deposited catalyst enter a settler from the outlet of a riser of the first reactor for rapid separation, and the first carbon-deposited catalyst enters a stripper for stripping. The second catalytic cracking catalyst is a coked catalyst directly from the stripper. The bottom of the stripper is connected with the bottom of the riser reactor of the second reactor through a second spent agent conveying pipe, an electric heating facility is arranged on the second spent agent conveying pipe, the carbon-deposited catalyst at 500 ℃ in the stripper is rapidly heated to 680 ℃ and directly conveyed to the bottom of the riser reactor of the second reactor, and flows upwards under the action of pre-lifting steam. And (3) allowing a light hydrocarbon raw material (9-57 ℃) to enter a second reactor and contact with a carbon deposition catalyst from a stripper for catalytic conversion reaction. The weight ratio of the light hydrocarbon to the heavy hydrocarbon oil raw material is 0.2: 1. the reaction oil mixture from the second reactor is further introduced into a fluidized bed of a third reactor at the outlet of a riser for further reaction, oil gas (a second product) and a second carbon-deposited catalyst after the reaction enter a settler through the fluidized bed reactor for separation, and the separated second carbon-deposited catalyst enters a stripper for stripping under the action of gravity. The reaction oil gas from the first reactor and the third reactor is led out from the settler and introduced into the product separating system for product separation to obtain gas product and various liquid products, and partial reaction oil gas is separated to obtain light hydrocarbon. After the carbon-deposited catalyst from the first reactor and the second reactor is stripped in the stripper, one part of the carbon-deposited catalyst is directly returned to the second reactor for recycling, the rest is returned to the oxygen-containing fluidizing zone of the regenerator through the first spent agent conveying pipe, all the carbon-deposited catalyst in the regenerator is contacted with heated air and is subjected to aerobic regeneration at 700 ℃, and the obtained hot regenerated catalyst is returned to the first reactor for recycling. The regenerator is additionally provided with an external heat taking device to take away surplus heat, and the reaction heat balance and the regeneration temperature of the regenerator are controlled. The carbon deposit content of the first catalytic cracking catalyst (first regenerated catalyst), and the carbon deposit content and the hydrogen mass fraction of the carbon deposit of the second catalytic cracking catalyst (second regenerated catalyst) were detected during continuous operation. The main operating conditions and results are listed in table 3. 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 reducing dry gas and coke yields during catalytic conversion of hydrocarbon oils.

The reaction apparatus used was the same as in example i. The raw materials and the main experimental steps are the same as those in the example I, except that the distillation range of the selected light hydrocarbon is 9-205 ℃. The weight ratio of light hydrocarbon to heavy hydrocarbon oil raw material is 0.25: 1. the main operating conditions and results are listed in table 3. In the oxygen-free fluidizing zone of the regenerator, a portion of the fluidizing gas (laboratory selected from normal nitrogen, which under standard conditions had a nitrogen content of > 99.5% by volume and an oxygen content of < 0.5% by volume) was introduced.

Comparative example 3

Comparative example 3 illustrates the effect of the hydrocarbon oil catalytic conversion process when the regenerator has only an oxygen-containing fluidized zone and the second catalytic cracking catalyst is 500 c coked catalyst from the stripper and 680 c regenerated catalyst from the regenerator are mixed in a weight ratio of 1: 1.

The experiment was carried out using a medium-sized apparatus for continuous reaction-regeneration operation having three reactors, wherein the first reactor was a riser, and the riser reactor had an inner diameter of 16 mm and a height of 3800 mm. The second reactor was a riser reactor with an internal diameter of 16 mm and a height of 3200 mm. The outlet of the second riser reactor was introduced into a fluidized bed reactor having an internal diameter of 64 mm and a height of 300 mm.

The first catalytic cracking catalyst is regenerated catalyst with temperature of 680 ℃, enters the bottom of the riser reactor of the first reactor through the first catalytic cracking catalyst inclined tube, and flows upwards under the action of pre-lifting steam. Heavy hydrocarbon oil raw materials (main properties are shown in table 2) are heated to 350 ℃ by a preheating furnace, mixed with atomized water vapor, sprayed into a first reactor through a feeding nozzle, and contacted with a hot regenerated catalyst to carry out catalytic conversion reaction. The reaction oil gas (first product) and the first carbon-deposited catalyst enter a settler from the outlet of a riser of the first reactor for rapid separation, and the first carbon-deposited catalyst enters a stripper. The second catalytic cracking catalyst is a carbon deposit catalyst from a stripper and a regenerated catalyst from a regenerator at 680 ℃ which are mixed according to the weight ratio of 1: 1. The bottom of the stripper is connected with the bottom of the riser reactor of the second reactor through a second spent catalyst conveying pipe, an electric heating facility is arranged on the second spent catalyst conveying pipe, the carbon-deposited catalyst at 500 ℃ in the stripper is rapidly heated to 520 ℃ and directly conveyed to the bottom of the riser reactor of the second reactor, and meanwhile, the regenerated catalyst is also conveyed to the bottom of the riser reactor of the second reactor through a second catalytic cracking catalyst inclined pipe and flows upwards under the action of pre-lifting steam. And (3) allowing the light hydrocarbon raw material (9-57 ℃) to enter a second reactor to contact with a second catalytic cracking catalyst for catalytic conversion reaction. The weight ratio of light hydrocarbon to heavy hydrocarbon oil raw material is 0.25: 1. the reaction oil mixture from the second reactor is further introduced into a fluidized bed of a third reactor at the outlet of a riser for further reaction, oil gas (a second product) and a second carbon-deposited catalyst after the reaction enter a settler through the fluidized bed reactor for separation, and the separated second carbon-deposited catalyst enters a stripper under the action of gravity. The reaction oil gas from the first reactor and the third reactor is led out from the settler and introduced into the product separating system for product separation to obtain gas product and various liquid products, and partial reaction oil gas is separated to obtain light hydrocarbon. After the carbon deposited catalysts from the first reactor and the second reactor are stripped in a stripper, one part of the carbon deposited catalysts are directly returned to the second reactor for recycling, the rest of the carbon deposited catalysts are returned to an oxygen-containing fluidization area of a regenerator through a first standby agent conveying pipe, all the carbon deposited catalysts in the regenerator are contacted with heated air and regenerated at 700 ℃ to obtain hot regenerated catalysts, one part of the regenerated catalysts are mixed with the first carbon deposited catalysts from the stripper and are sent to the second reactor for recycling, and the rest of the regenerated catalysts are returned to the first reactor for recycling. The regenerator is additionally provided with an external heat taking device to take away surplus heat, and the reaction heat balance and the regeneration temperature of the regenerator are controlled. The carbon deposit content of the first catalytic cracking catalyst (first regenerated catalyst), and the carbon deposit content and the hydrogen mass fraction of the carbon deposit of the second catalytic cracking catalyst (second regenerated catalyst) were detected during continuous operation. The main operating conditions and results are listed in table 3. The main operating conditions and results are listed in table 3.

Example 3

The examples illustrate the effectiveness of the process of the present invention in reducing the yield of dry gas and coke during the catalytic conversion of hydrocarbon oils.

The reaction apparatus used was the same as in example i. The raw materials and the main experimental steps are the same as those in the example I, except that the distillation range of the selected light hydrocarbon is 145-235 ℃. The weight ratio of the light hydrocarbon to the heavy hydrocarbon oil raw material is respectively 0.2: 1. the main operating conditions and results are listed in table 3. In the oxygen-free fluidizing zone of the regenerator, a portion of the fluidizing gas (laboratory selected from normal nitrogen, which under standard conditions had a nitrogen content of > 99.9% by volume and an oxygen content of < 0.1% by volume) was introduced.

TABLE 1

TABLE 2

Raw oil name	Atmospheric residuum
		Density (20 ℃), kg/m³	891.6
The composition of elements%
		C	86.20
H	13.06
		S	0.28
N	0.29
		Basic nitrogen	922
Group composition of%
		Saturated hydrocarbons	59.0
Aromatic hydrocarbons	22.3
		Glue	18.3
Asphaltenes	0.4
		Residual carbon value,%)	5.44
Kinematic viscosity, mm²/s
		80℃	32.65
100℃	18.77
		Freezing point, DEG C	>50
Refractive index, 70 deg.C	1.4848
		Total acid value of mgKOH/g	0.44
Average relative molecular mass	528
		Metal content, mg/kg
Fe	4.2
		Ni	17.9
Cu	<0.1
		V	0.2
Na	0.3
		Ca	0.7
Zn	0.9
		Reduced pressure volumetric distillation range, deg.C
IBP	258.0
		5％	365.9
10％	388.7
		30％	435.7
50％	489.0
		66.5％	569.4

TABLE 3

As can be seen from table 3, example 1, using the process provided by the present disclosure, the yield of propylene from catalytic cracking was 18.89 wt%, and the dry gas and coke yields were reduced by 3.17 wt% in total compared to comparative example 1.

The second catalytic cracking catalyst in comparative example 2 was a heated coke catalyst with a higher hydrogen mass fraction of coke and a propylene yield of only 15.92 wt% although the dry gas and coke yields were reduced. While the mass fraction of hydrogen in the carbon deposit on the second catalytic cracking catalyst in example 1 was significantly reduced, the dry gas and coke yield decreased by 1.60 wt% and the propylene yield increased by 2.97 wt% compared to comparative example 2.

In comparison with comparative example 3, in which the second catalytic cracking catalyst is a mixture of heated carbon-deposited catalyst and regenerated catalyst, in example 2, the yield of propylene produced by catalytic conversion is up to 18.72 wt% and the yield of dry gas and coke is reduced to 1.91 wt% in total by using the method provided by the present disclosure.

In example 3, using relatively mild reaction conditions, the yield of propylene reached 17.11 wt%, the yield of dry gas was only 4.22 wt%, and the yield of coke was only 8.04 wt%.

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

It should be noted that, in the foregoing embodiments, various features described in the above embodiments may be combined in any suitable manner, and in order to avoid unnecessary repetition, various combinations that are possible in the present disclosure are not described again.

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

Claims

1. A catalytic conversion process for reducing dry gas and coke yields, the process comprising:

introducing the first carbon-deposited catalyst and the second carbon-deposited catalyst into a stripper for stripping, introducing a first part of the stripped catalyst into an oxygen-containing fluidized zone of a regenerator for aerobic regeneration to obtain a first regenerated catalyst, and feeding the first regenerated catalyst into a first reactor as the first catalytic cracking catalyst;

2. The method of claim 1, wherein the method further comprises: introducing a fluidizing gas having an oxygen content of less than 3% by volume into the oxygen-free fluidizing zone, the fluidizing gas containing not less than 97% by volume of an inert gas, the inert gas being at least one selected from the group consisting of nitrogen, helium, neon, argon, krypton, and xenon.

3. The process of claim 1, wherein the oxygen-free fluidizing zone of the regenerator comprises: a low-temperature catalyst conveying pipe (921), a gas-solid separator (922), a catalyst distributor (923), a heat taking exhaust pipe (924), a degassing pipe (925) and a high-temperature catalyst conveying pipe (926);

the one end of low temperature catalyst conveyer pipe (921) with stripper (7) intercommunication, the other end is connected with gas-solid separator (922), the below of gas-solid separator (922) have catalyst distributor (923) and with catalyst distributor (923) intercommunication get hot calandria (924), get hot calandria (924) for the pipeline that the multiunit set up and bottom intercommunication along the vertical direction, degasification pipe (925) connect in get the lower extreme of hot calandria (924) and seted up fluidization gas entry (927) in the junction, degasification pipe (925) for the lower extreme with the riser that high temperature catalyst conveyer pipe (926) link to each other, high temperature catalyst conveyer pipe (926) with second reactor (2) intercommunication.

4. A process according to claim 1, wherein the first regenerated catalyst has a temperature of 560 to 800 ℃ and a carbon deposit content of 0.01 to 0.1% by weight, based on the dry weight of the first regenerated catalyst.

5. The process according to claim 1, wherein the temperature of the second regenerated catalyst is 560 to 800 ℃, and the carbon deposit content of the second regenerated catalyst is 0.5 to 1.9 wt% based on the dry weight of the second regenerated catalyst.

6. The method of claim 5, wherein the second regenerated catalyst has a carbon deposit content of 0.9 to 1.3 wt% based on the dry weight of the second regenerated catalyst.

7. The method of claim 5, wherein the second regenerated catalyst has a carbon deposit content of 0.91 to 0.99 wt% based on the dry weight of the second regenerated catalyst.

8. The method according to claim 5, wherein the hydrogen content in the carbon deposit of the second regenerated catalyst is 0.1 to 0.65% by weight based on the weight of the carbon deposit in the second regenerated catalyst.

9. The method according to claim 8, wherein the hydrogen content in the carbon deposit of the second regenerated catalyst is 0.1 to 0.5% by weight based on the weight of the carbon deposit in the second regenerated catalyst.

10. The process according to claim 1, wherein 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, 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.

11. The process according to claim 1, wherein the heavy hydrocarbon oil feedstock is a petroleum hydrocarbon oil.

12. The process according to claim 10, wherein the average relative molecular mass of the heavy hydrocarbon oil feedstock is not less than 200.

13. The method according to claim 1, wherein the light hydrocarbon has a distillation range of 9 to 160 ℃;

the olefin content of the light hydrocarbon is 30 to 90 wt% based on the total weight of the light hydrocarbon.

14. The process according to claim 13, wherein the light hydrocarbons have a distillation range of 9 to 60 ℃.

15. The method of claim 13, wherein the light hydrocarbon has an olefin content of 45 to 90 wt% based on the total weight of the light hydrocarbon.

16. The method of claim 1, wherein the light hydrocarbons are at least partially derived from the separation products of the product separation system.

17. The method of claim 1, wherein the weight ratio of the light hydrocarbons to the heavy hydrocarbon oil feedstock is (0.01-0.6): 1.

18. the process according to claim 17, wherein the weight ratio of the light hydrocarbons to the heavy hydrocarbon oil feedstock is (0.05-0.3): 1.

19. the process of claim 1 wherein the first and second catalytic cracking catalysts 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 the MFI structure, ferrierites, chabazites, cyclospar, erionites, a-zeolites, epistillomites, and laumontites.

20. The method of claim 1, wherein the operating conditions of the first catalytic reaction comprise: the reaction temperature is 480-600 ℃; the reaction time is 0.5-10 seconds; the weight ratio of the agent oil is (5-15): 1; the weight ratio of water to oil is (0.05-1): 1.

21. the method of claim 1, wherein the operating conditions of the second catalytic 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; the weight ratio of water to oil is (0.1-1): 1.

22. the process of claim i, wherein the first reactor and the second reactor are each one selected from the group consisting of a riser reactor, a downer reactor, a fluidized bed reactor, a riser and downer composite reactor, a riser and fluidized bed composite reactor, and a downer and fluidized bed composite reactor.

23. The process of claim 1, wherein the third reactor is a fluidized bed reactor and the operating conditions of the third reactor are: the reaction temperature is 450-750 ℃; the weight hourly space velocity is 1-30 h^-1(ii) a The absolute pressure of the settler is 0.15-0.40 MPa.

24. The process of claim 23, wherein the reaction temperature of the third reactor is 510-560 ℃.