CN110724561B

CN110724561B - Catalytic cracking method and system for producing propylene and light aromatic hydrocarbon

Info

Publication number: CN110724561B
Application number: CN201810779830.7A
Authority: CN
Inventors: 魏晓丽; 龚剑洪; 张执刚
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: 2018-07-16
Filing date: 2018-07-16
Publication date: 2021-03-16
Anticipated expiration: 2038-07-16
Also published as: CN110724561A

Abstract

The invention relates to a catalytic cracking method and a catalytic cracking system for producing propylene and light aromatic hydrocarbon, wherein the method comprises the following steps: introducing the light raw oil into the dilute phase transport bed from the lower part of the dilute phase transport bed to contact with a first catalytic cracking catalyst and carrying out a first catalytic cracking reaction from bottom to top to obtain a first reaction product and a semi-spent catalyst; sending the obtained first reaction product and the semi-spent catalyst into a dense-phase fluidized bed for continuously carrying out a second catalytic cracking reaction to obtain a second reaction product and a first spent catalyst; and introducing the preheated inferior heavy oil into the rapid fluidized bed from the lower part of the rapid fluidized bed to contact with the second catalytic cracking catalyst and perform a third catalytic cracking reaction from bottom to top to obtain a third reaction product and a second spent catalyst. The method and the system have high yield of propylene and light aromatic hydrocarbon for catalytic cracking and low yield of coke and dry gas.

Description

Catalytic cracking method and system for producing propylene and light aromatic hydrocarbon

Technical Field

The invention relates to a catalytic cracking method and a catalytic cracking system for producing propylene and light aromatic hydrocarbons.

Background

Propylene is an important organic chemical raw material, the equivalent consumption of propylene in 2016 years is 3380 ten thousand tons, and the equivalent self-supporting rate is 75.2%. The equivalent consumption of propylene in China can reach 3900 ten thousand tons by 2020, and a certain space exists in the gap of the capacity. Benzene, toluene, xylene (BTX) are important basic chemical raw materials, among which paraxyleneBenzene (PX) accounts for about 45% of the total BTX consumed. With the development of polyester and other industries in China, the demand of BTX is expected to continue to increase at a high speed. About 90% of ethylene, about 70% of propylene, 90% of butadiene, and 30% of aromatics are all from steam cracking by-products. Although the steam cracking technology is developed for decades and the technology is continuously improved, the steam cracking technology still has the advantages of high energy consumption, high production cost and CO₂The discharge amount is large, the product structure is not easy to adjust, and other technical limitations are imposed, if the petrochemical industry adopts the traditional route of preparing ethylene and propylene by steam cracking, the petrochemical industry faces a plurality of restrictive factors such as shortage of light raw oil, insufficient production capacity, high cost and the like, and in addition, along with the lightening of the steam cracking raw material, the reduction of the yield of propylene and light aromatic hydrocarbon is more an aggravated supply-demand contradiction. The catalytic cracking technology can be used as a beneficial supplement to the production process for producing the low-carbon olefin and the light aromatic hydrocarbon, and has obvious social and economic benefits for oil refining and chemical engineering integrated enterprises by adopting a catalytic technical route to produce chemical raw materials.

Chinese patent CN98101765.7 discloses a method for simultaneously preparing low-carbon olefin and high-aromatic gasoline from heavy oil, which is to make heavy petroleum hydrocarbon and steam undergo catalytic cracking reaction in a composite reactor composed of a lift pipe and a dense-phase fluidized bed, so as to increase the yield of low-carbon olefin, especially propylene, and simultaneously increase the aromatic content in gasoline to about 80 wt%. CN01119807.9 discloses a method for increasing the yield of ethylene and propylene by catalytic conversion of heavy petroleum hydrocarbon, which is to make the hydrocarbon oil raw material contact and react with a catalyst containing pentasil zeolite in a riser or fluidized bed reactor. Chinese patent CN 200410068934.5 discloses a method for producing low-carbon olefins and aromatics by catalytic cracking in two reaction zones, wherein the two reaction zones adopt different weight hourly space velocities to achieve the purpose of producing low-carbon olefins such as propylene and ethylene from heavy raw materials to the maximum extent, wherein the yield of propylene exceeds 20 wt%, and simultaneously co-producing aromatics such as toluene and xylene.

US patents US2002003103 and US2002189973 employ a dual riser FCC unit for propylene production. Wherein gasoline (60-300 DEG F/15-150 ℃) generated by the cracking reaction is fed into the second riser, and the catalyst is a mixture of USY molecular sieve and ZSM-5 molecular sieve catalyst.

U.S. patent US2002195373 uses a downflow reactor operating at high temperature (1020-1200 ° f/550-650 ℃), short contact time (<0.5 seconds) and large solvent to oil ratio (15-25). The procatalyst (Y-type faujasite) has low hydrogen transfer activity and is formulated to maximize light olefin yield in conjunction with operating conditions. The high efficiency separator separates the product from the catalyst within 0.1 seconds, minimizing secondary reactions and coke formation. In addition, LCO is used to quench the separated gaseous product to about 930 DEG F/500 ℃ and prevent further cracking.

US patent US6153089 uses a circulating fluidized bed reactor/regenerator to convert olefin-rich hydrocarbons to hydrogen and C2-C4 olefins. It uses a catalyst containing dehydrogenation metal, shape-selective molecular sieve and macroporous acidic component. The reactor/regenerator is operated at 840-1380 DEG F/450-750 ℃ and WHSV of 0.1-60 h-1.

US patent nos. US6538169 and US2003121825 are also constructed of a reaction-regeneration system employing two reaction zones and a common regenerator. In the first reaction zone, the heavy feedstock is cracked to light olefins or intermediates that can be converted to light olefins using high temperature and high catalyst to oil ratios. The second reaction zone consists of a second riser where the operating conditions are more severe and more light components are produced from the gasoline product. The conversion of gasoline to light olefins is aided by the use of shape selective molecular sieves such as ZSM-5, suitable feedstocks include VGO, HVGO and hydrogenated gas oil.

Chinese patent CN200710120105 discloses a method for preparing ethylene, propylene and light aromatic hydrocarbon, which adopts a reducing riser reactor to divide raw materials into three types of easy cracking, difficult cracking and difficult cracking, wherein raw materials with different properties enter different reaction regions of the reducing riser for catalytic conversion. Chinese patent CN201010233651 discloses a catalytic conversion method for producing propylene and light aromatic hydrocarbons, which adopts a combined reactor, a heavy raw material adopts a riser connected in series with a dense-phase fluidized bed reactor, and light hydrocarbons adopt a riser reactor.

The structural contradiction of the oil refining chemical industry in China is increasingly serious, on one hand, the excess capacity of the traditional petrochemical products and the contradiction between the supply and the demand of the finished oil are prominent, on the other hand, the shortage of resource products and high-end petrochemical products is prominent, and the transformation of oil refining to the chemical industry is great tendency. At present, the proportion of atmospheric residue oil blended by a catalytic cracking device is getting larger and larger, and even the requirement of blending vacuum residue oil is raised, the most advanced catalytic cracking technology of the existing catalytic cracking technology which usually takes vacuum wax oil or paraffin-based atmospheric residue oil as a raw material adopts a reactor with double lifting pipes or lifting pipes connected in series with a dense-phase bed layer, and under the reaction condition of higher severity, the aim of producing more low-carbon olefin and/or light aromatic hydrocarbon is achieved, and the problem of high yield of dry gas and coke inevitably occurs when the reactor is used for processing slag-blended heavy oil. A decrease in coke yield can be achieved with a downflow reactor, but the reaction conversion is relatively low and requires a specialized catalyst. Along with the heavy-duty of raw materials, the requirements of blending residual oil in a catalytic cracking device are more and more, and in order to efficiently utilize inferior heavy oil resources and meet the increasing demands of chemical raw materials such as low-carbon olefins and aromatic hydrocarbons, it is necessary to develop a catalytic cracking method for converting the inferior heavy oil raw materials into high value-added products.

Disclosure of Invention

The invention aims to provide a catalytic cracking method and a catalytic cracking system for producing propylene and light aromatic hydrocarbons.

In order to achieve the above object, the present invention provides a catalytic cracking process for producing propylene and light aromatic hydrocarbons, the process comprising:

introducing the light raw oil into the dilute phase transport bed from the lower part of the dilute phase transport bed to contact with a first catalytic cracking catalyst and carrying out a first catalytic cracking reaction from bottom to top to obtain a first reaction product and a semi-spent catalyst; the light raw oil is one, two or three selected from C4 fraction, gasoline fraction and diesel oil fraction;

sending the obtained first reaction product and the semi-spent catalyst into a dense-phase fluidized bed for continuously carrying out a second catalytic cracking reaction to obtain a second reaction product and a first spent catalyst;

introducing preheated inferior heavy oil into the rapid fluidized bed from the lower part of the rapid fluidized bed to contact with a second catalytic cracking catalyst and perform a third catalytic cracking reaction from bottom to top to obtain a third reaction product and a second spent catalyst; the properties of the inferior heavy oil meet one, two, three or four of the following indexes: the density at 20 ℃ is 900-³2-10 wt% of carbon residue, 2-30ppm of total content of nickel and vanadium, and a characteristic factor K value less than 12.1; the catalyst in the fast fluidized bed is distributed in a full dense phase, and the distribution of the axial solid fraction epsilon in the fast fluidized bed meets the following requirements: epsilon is more than or equal to 0.1 and less than or equal to 0.2;

separating the second reaction product and the third reaction product to obtain dry gas, liquefied gas, pyrolysis gasoline, pyrolysis diesel oil and pyrolysis heavy oil;

and feeding the first spent catalyst and the second spent catalyst into a regenerator for coke burning regeneration, and returning the obtained regenerated catalyst serving as the first catalytic cracking catalyst and the second catalytic cracking catalyst to the bottoms of the dilute phase conveying bed and the fast fluidized bed.

Optionally, the properties of the inferior heavy oil meet one, two, three or four of the following criteria: the density at 20 ℃ is 910-³The carbon residue is 3-8 wt%, the total content of nickel and vanadium is 5-20ppm, and the characteristic factor K value is less than 12.0.

Optionally, the inferior heavy oil is heavy petroleum hydrocarbon and/or other mineral oil; the heavy petroleum hydrocarbon is one or more selected from vacuum residue, poor atmospheric residue, poor hydrogenated residue, coker gas oil, deasphalted oil, vacuum wax oil, high acid value crude oil and high metal crude oil, and the other mineral oil is one or more selected from coal liquefied oil, oil sand oil and shale oil.

Optionally, the first catalytic cracking catalyst comprises, on a dry basis and by weight on a dry basis of the first catalytic cracking catalyst, from 1 to 50 wt% of a zeolite, from 5 to 99 wt% of an inorganic oxide, and from 0 to 70 wt% of a clay;

on a dry basis and based on the weight of the second catalytic cracking catalyst on a dry basis, the second catalytic cracking catalyst comprising from 1 to 50 wt% of a zeolite, from 5 to 99 wt% of an inorganic oxide, and from 0 to 70 wt% of a clay;

the zeolites include medium pore zeolites which are ZSM series zeolites and/or ZRP zeolites and optionally large pore zeolites which are one or more selected from rare earth Y, rare earth hydrogen Y, ultrastable Y and high silica Y.

Optionally, the medium pore zeolite of the first catalytic cracking catalyst comprises from 0 to 50 wt% of the total weight of zeolite on a dry basis;

the medium pore zeolite of the second catalytic cracking catalyst comprises from 0 to 50 wt% of the total weight of zeolite on a dry basis.

Optionally, the medium pore zeolite of the first catalytic cracking catalyst comprises 0 to 20 wt% of the total weight of zeolite on a dry basis;

the medium pore zeolite of the second catalytic cracking catalyst comprises from 0 to 20 wt% of the total weight of zeolite on a dry basis.

Optionally, the conditions of the first catalytic cracking reaction include: the reaction temperature is 550-700 ℃, the reaction time is 0.1-5 seconds, and the weight ratio of the catalyst to the oil is (5-30): 1, the weight ratio of water to oil is (0.05-2): 1, the catalyst density is 20-100 kg/m³Gas linear speed of 4-18 m/s, reaction pressure of 0.2-1.2 MPa, and catalyst mass flow rate G_sIs 180-500 kg/(meter)²Seconds).

Optionally, the conditions of the second catalytic cracking reaction include: the reaction temperature is 510-650 ℃, and the weight hourly space velocity is 1-20 h^-1The weight ratio of the agent oil is (5-30): 1, catalyst density of 300-³The gas linear speed is 0.4-0.8 m/s, and the reaction pressure is 130-450 MPa.

Optionally, the conditions of the third catalytic cracking reaction include: the reaction temperature is 510-650 ℃, the reaction time is 1-20 seconds, and the weight ratio of the catalyst to the oil is (3-50): 1, the weight ratio of water to oil is (0.03-0.8): 1,the catalyst density is 120-290 kg/m³Gas linear speed of 0.8-2.5 m/s, reaction pressure of 130-450 kPa, and catalyst mass flow rate G_sIs 15-150 kg/(meter)²Seconds).

Optionally, the conditions of the third catalytic cracking reaction include: the reaction temperature is 550-620 ℃, the reaction time is 3-15 seconds, and the weight ratio of the catalyst to the oil is (10-30): 1, the weight ratio of water to oil is (0.05-0.5): 1, catalyst density of 150-³Gas linear velocity of 1-1.8 m/s, catalyst mass flow rate G_sIs 20-130 kg/(meter)²Seconds).

Optionally, the method further includes: replenishing the catalyst into the dilute phase conveying bed and/or the fast fluidized bed; wherein the carbon content of the supplemented catalyst is 0-1.0 wt.%.

Optionally, the catalyst supplemented to the dilute-phase conveying bed accounts for 0-50 wt% of the catalyst circulation amount of the dilute-phase conveying bed; the replenished catalyst of the fast fluidized bed accounts for 0-50 wt% of the circulating amount of the fast fluidized bed catalyst.

Optionally, the catalyst supplemented to the dilute-phase conveying bed accounts for 5-30 wt% of the catalyst circulation amount of the dilute-phase conveying bed; the replenished catalyst of the fast fluidized bed accounts for 5-30 wt% of the circulating amount of the fast fluidized bed catalyst.

Optionally, the distance between the catalyst replenishing position of the dilute phase conveying bed and the bottom of the dilute phase conveying bed accounts for 0-2/3 of the total height of the dilute phase conveying bed;

the distance between the catalyst replenishing position of the fast fluidized bed and the bottom of the fast fluidized bed accounts for 0 to 2/3 of the total height of the fast fluidized bed.

Optionally, the regenerated catalyst obtained by the regeneration of the regenerator through coking is cooled to 600-680 ℃ by a cooler and then returns to the bottom of the dilute phase conveying bed and/or the fast fluidized bed.

The invention also provides a catalytic cracking system, which comprises a dilute phase conveying bed, a dense phase fluidized bed, a fast fluidized bed, optional oil agent separation equipment, reaction product separation equipment and a regenerator, wherein a settler which is communicated with fluid is arranged above the dense phase fluidized bed, and a stripping section which is communicated with fluid is arranged below the dense phase fluidized bed;

the dilute phase conveying bed is provided with a catalyst inlet at the bottom, a light raw oil inlet at the lower part and an oil outlet at the top, the dense phase fluidized bed is provided with an oil inlet, the stripping section is provided with a catalyst outlet, the settler is provided with a reaction product outlet and an oil inlet, the fast fluidized bed is respectively and independently provided with a catalyst inlet at the bottom, an inferior heavy oil inlet at the lower part and an oil outlet at the top, the oil separation equipment is provided with an oil inlet, a catalyst outlet and a reaction product outlet, the reaction product separation equipment is provided with a reaction product inlet, a dry gas outlet, a liquefied gas outlet, a pyrolysis gasoline outlet, a pyrolysis diesel oil outlet and a pyrolysis heavy oil outlet, and the regenerator is provided with a catalyst inlet and a catalyst outlet;

catalyst inlets of the dilute phase conveying bed and the fast fluidized bed are both communicated with a catalyst outlet of the regenerator in a fluid mode, the dilute phase conveying bed extends into an oil agent inlet of the dense phase fluidized bed from bottom to top, an oil agent outlet of the dilute phase conveying bed is located in the dense phase fluidized bed, a reaction product outlet of the settler is communicated with an oil agent inlet of the oil agent separation equipment in a fluid mode, an oil agent outlet of the fast fluidized bed is communicated with the oil agent inlet of the settler in a fluid mode or communicated with the oil agent inlet of the oil agent separation equipment in a fluid mode, a reaction product outlet of the oil agent separation equipment is communicated with a reaction product inlet of the reaction product separation equipment in a fluid mode, and a catalyst outlet of the stripping section is communicated with a catalyst inlet of the regenerator in a fluid mode.

The poor quality heavy raw material and the light raw oil are respectively subjected to catalytic cracking reaction in different reactors. Different reactors may employ different reaction conditions for the nature of the feedstock, which helps to improve feedstock conversion and target product yield.

The invention adopts the fast fluidized bed to effectively improve the density of the reaction catalyst, thereby greatly improving the ratio of the instantaneous catalyst to the raw oil in the reactor, controlling the relatively long reaction time, leading the catalyst to be capable of fully reacting with the inferior heavy oil, not only improving the reaction conversion rate, but also improving the yield of the low-carbon olefin and the light aromatic hydrocarbon, simultaneously effectively reducing the generation of dry gas and coke, and improving the product distribution and the product quality.

The invention can enable petrochemical enterprises to produce high-added-value chemical raw materials from cheap inferior heavy oil to the maximum extent, is beneficial to promoting the refining and chemical integration process of oil refining enterprises in China, not only solves the problem of petrochemical raw material shortage, but also improves the economic benefit and social benefit of petrochemical industry.

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

Drawings

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

FIG. 1 includes a schematic flow diagram of one embodiment of the method of the present invention and also includes a schematic structural diagram of one embodiment of the system of the present invention.

Description of the reference numerals

I fast fluidized bed II dilute phase transport bed III dense phase fluidized bed

1 pipeline 2 pre-lifting section 3 outlet section

4 settler 5 stripping section 6 cyclone

7 gas collection chamber, 8 pipeline, 9 to-be-grown inclined tube

10 regenerator 11 regeneration inclined tube 12 pipeline

13 air distributor 14 line 15 regeneration chute

16 line 17 line

Detailed Description

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

In the invention:

(1) the axial solid fraction is the pressure difference between two points in the axial direction of the reactor measured by a pressure difference meter, the distance between two points in the axial direction divided by the density of the catalyst particles; the unit of the pressure difference is kilogram/meter²The distance between two axial points is measured in meters, and the density of catalyst particles is kg/m³And the two axial points are any two axial points of the reactor.

(2) The reaction time is equal to the volume of the reactor/the logarithmic mean volume flow of oil gas; volume unit of the reactor is meter³The unit of logarithmic mean volume flow of oil and gas is meter³A/second;

logarithmic mean volume flow rate of oil and gas (V)_out-V_in)/ln(V_out/V_in)，V_outAnd V_inThe volume flow of oil gas at the outlet and the inlet of the reactor respectively;

the volume flow of oil gas at the outlet of the reactor is m/rho₃The volume flow of oil gas at the inlet of the reactor is m/rho₄(ii) a m is the feeding amount of raw oil and atomized steam in unit time, and the unit is kilogram/second; rho₃The density of oil gas at the outlet of the reactor is measured in kg/m³；ρ₄The density of the oil gas at the inlet of the reactor is measured in kg/m³。

(3) The density of the catalyst in the reactor is the reaction time multiplied by the catalyst circulation volume divided by the volume of the reactor; the reaction time is in seconds, the catalyst circulation is in kilograms per second, and the reactor volume is in meters³。

(4) And the linear gas velocity is the logarithmic mean volume flow of oil gas/sectional area of the reactor.

(5) Catalyst mass flow rate G_sCatalyst circulation rate ÷ reactor cross-sectional area; the unit of the circulating amount of the catalyst is kilogram/second;

the catalyst circulation amount is divided by the coke generation speed (the carbon content of the spent catalyst-the carbon content of the regenerated catalyst), the unit of the coke generation speed is kilogram/second, and the carbon content of the spent catalyst and the carbon content of the regenerated catalyst are both weight contents;

coke formation rate ═ flue gas mass × (CO)₂% + CO%) +/-Vm × M; vm is the molar volume of gas and takes the value of 22.4 multiplied by 10^-3Rice and its production process³M is the molar mass of carbon and takes the value of 12 multiplied by 10^-3Kilogram/mole;

flue gas amount (regeneration air amount × 79 vol%)/(1-CO)₂％-CO％-O₂%) of the amount of regenerated air in meters³Second, the unit of smoke is meter³Second, CO₂％、CO％、O₂% of CO in the flue gas₂、CO、O₂Volume percent of (c).

The invention provides a catalytic cracking method for producing propylene and light aromatic hydrocarbon, which comprises the following steps:

In the invention, the fast fluidized bed refers to a reactor in which a catalyst is in fast fluidization, the fast fluidization is bubble-free gas-solid contact fluidization, and the important characteristic is that solid particles tend to move in an agglomeration manner. The axial solid fraction epsilon distribution of all heights in the fast fluidized bed meets the following requirements: epsilon is more than or equal to 0.1 and less than or equal to 0.2, the catalyst in the rapid fluidized bed is distributed in a full-dense phase, the catalyst is prevented from being distributed in a dilute-down-dense mode, the actual agent-oil ratio above and below the rapid fluidized bed is kept consistent, the dry gas coke yield is reduced, and the target product yield is improved.

According to the invention, the catalyst is regulated to be in full-dense phase distribution by regulating the linear velocity of gas in the fast fluidized bed and arranging a gas distributor at the feeding position of the fast fluidized bed, wherein the gas distributor can be one or more of gas distributors which are common in industry, such as flat plates, arches, discs, rings and umbrella-shaped gas distributors, so that the raw oil atomized by atomized steam is contacted with the catalyst in a uniform concentration in the axial direction of the reactor to carry out catalytic cracking reaction, and the generation of catalyst-oil ratio coke and thermal reaction coke caused by overhigh or overlow concentration of the catalyst is reduced.

According to the invention, the inferior heavy oil can be introduced into the fast fluidized bed at one feeding position or can be introduced into the fast fluidized bed from two or more feeding positions according to the same or different proportions.

According to the present invention, the inferior heavy oil refers to heavy oil which is less suitable for catalytic cracking process than conventional heavy oil, and for example, the inferior heavy oil may satisfy one, two, three or four of the following properties: a density of 900-. In particular, the low-grade heavy oil can be heavy petroleum hydrocarbons and/or other mineral oils; the heavy petroleum hydrocarbon may be one or more selected from Vacuum Residue (VR), low-grade Atmospheric Residue (AR), low-grade hydrogenated residue, coker gas oil, deasphalted oil, vacuum wax oil, high acid number crude oil, and high metal crude oil, and the other mineral oil may be one or more selected from coal liquefied oil, oil sand oil, and shale oil. The carbon residue in the inferior heavy oil is measured by adopting an ASTMD-189 Conradson carbon residue experimental method.

Catalytic cracking catalysts are well known to those skilled in the art in accordance with the present invention, and for the process of the present invention, the first catalytic cracking catalyst may comprise, on a dry basis and based on the weight of the first catalytic cracking catalyst on a dry basis, from 1 to 50 wt% of a zeolite, from 5 to 99 wt% of an inorganic oxide, and from 0 to 70 wt% of a clay; the second catalytic cracking catalyst may include, on a dry basis and by weight of the second catalytic cracking catalyst on a dry basis, from 1 to 50 wt% of a zeolite, from 5 to 99 wt% of an inorganic oxide, and from 0 to 70 wt% of a clay; the content and type of zeolite, inorganic oxide and clay in the first catalytic cracking catalyst and the second catalytic cracking catalyst may be the same or different, preferably the same. The zeolite may comprise, as active components, a medium pore zeolite and optionally a large pore zeolite, and the medium pore zeolite in the first catalytic cracking catalyst may comprise from 0 to 50 wt%, preferably from 0 to 20 wt%, of the total weight of zeolite on a dry basis; the medium pore zeolite in the second catalytic cracking catalyst may be present in an amount of from 0 to 50 wt%, preferably from 0 to 20 wt%, based on the total weight of zeolite on a dry basis. The medium and large pore zeolites are defined as conventional in the art, i.e., the medium pore zeolite has an average pore size of 0.5 to 0.6nm and the large pore zeolite has an average pore size of 0.7 to 1.0 nm. The large-pore zeolite can be selected from one or more of Rare Earth Y (REY), Rare Earth Hydrogen Y (REHY), ultrastable Y obtained by different methods and high-silicon Y. The medium pore zeolite may be selected from zeolites having an MFI structure, such as ZSM-series zeolites and/or ZRP zeolites, which may also be modified with non-metallic elements such as phosphorus and/or transition metal elements such as iron, cobalt, nickel, as described more fully in connection with ZRP, see U.S. Pat. No. 5,232,675, the ZSM-series zeolites are preferably selected from mixtures of one or more of ZSM-5, ZSM-11, ZSM-12, ZSM-23, ZSM-35, ZSM-38, ZSM-48 and other zeolites of similar structure, and the ZSM-5 is described more fully in connection with U.S. Pat. No. 5,232,6753,702,886. The inorganic oxide is preferably silicon dioxide (SiO) as a binder₂) And/or aluminum oxide (Al)₂O₃). The clay acts as a matrix (i.e., carrier) and is preferably selected from kaolin and/or halloysite.

According to the invention, the light raw oil is one, two or three selected from C4 fraction, gasoline fraction and diesel oil fraction. The C4 fraction refers to low molecular weight hydrocarbon which is in gas form at normal temperature and normal pressure and contains C4 fraction as main component, including alkane, alkene and alkyne with carbon number of 4 in molecule. It includes the gaseous hydrocarbon product produced by the apparatus of the present invention which is enriched in a C4 fraction, and may also include gaseous hydrocarbons produced by other plant processes which are enriched in a C4 fraction, with the C4 fraction produced by the apparatus of the present invention being preferred. The C4 fraction is preferably an olefin-rich C4 fraction, wherein the C4 olefin content is greater than 50 wt%, preferably greater than 60 wt%, and most preferably greater than 70 wt%. The gasoline fraction is selected from one or more of the catalytically cracked gasoline, straight-run gasoline, coker gasoline, thermally cracked gasoline and hydrogenated gasoline obtained by the method, wherein the catalytically cracked gasoline, the straight-run gasoline, the coker gasoline, the thermally cracked gasoline and the hydrogenated gasoline are gasoline from the outside of the device. The diesel fraction is one or more of the mixture of catalytic cracking diesel, straight-run diesel, coking diesel, thermal cracking diesel and hydrogenation diesel obtained by the method, wherein the catalytic cracking diesel, the straight-run diesel, the coking diesel, the thermal cracking diesel and the hydrogenation diesel are diesel from outside the device.

In the present invention, catalytic cracking is well known to those skilled in the art, and different reaction conditions can be selected for different reactors, and the conditions of the first catalytic cracking reaction can include: the reaction temperature is 550-700 ℃, the reaction time is 0.1-5 seconds, and the weight ratio of the catalyst to the oil is (5-30): 1, the weight ratio of water to oil is (0.05-2): 1, the catalyst density is 20-100 kg/m³Gas linear speed of 4-18 m/s, reaction pressure of 0.2-1.2 MPa, and catalyst mass flow rate G_sIs 180-500 kg/(meter)²Seconds); the conditions of the second catalytic cracking reaction may include: the reaction temperature is 510-650 ℃, and the weight hourly space velocity is 1-20 h^-1The weight ratio of the agent oil is (5-30): 1, catalyst density of 300-³The gas linear speed is 0.4-0.8 m/s, and the reaction pressure is 130-450 MPa; the conditions of the third catalytic cracking reaction may include: the reaction temperature is 510-650 ℃, the reaction time is 1-20 seconds, and the weight ratio of the catalyst to the oil is (3-50): 1, the weight ratio of water to oil is (0.03-0.8): 1, catalyst density of 120-³Gas linear speed of 0.8-2.5 m/s, reaction pressure of 130-450 kPa, and catalyst mass flow rate G_sIs 15-150 kg/(meter)²Seconds); the conditions of the third catalytic cracking reaction preferably include: the reaction temperature is 550-620 ℃, the reaction time is 3-15 seconds, and the weight ratio of the catalyst to the oil is (10-30): 1, the weight ratio of water to oil is (0.05-0.5): 1, catalyst density of 150-³Gas linear velocity of 1-1.8 m/s, catalyst mass flow rate G_sIs 20-130 kg/(meter)²Seconds).

The separation of the reaction products from the spent catalyst according to the present invention is well known to those skilled in the art, and may be performed, for example, in a settler using a cyclone separator, and the further separation of the reaction products to obtain dry gas, liquefied gas, pyrolysis gasoline, pyrolysis diesel oil and pyrolysis heavy oil is also well known to those skilled in the art, and the dry gas and liquefied gas may be further separated to obtain the desired products such as ethylene, propylene, etc. by separation means conventional in the art.

The coke-burning regeneration of the spent catalyst according to the present invention is well known to those skilled in the art and can be carried out in a regenerator, an oxygen-containing gas such as air can be introduced into the regenerator to contact the spent catalyst, and the flue gas from the coke-burning regeneration can be separated from the catalyst in the regenerator and then sent to a subsequent energy recovery system.

According to the invention, the method may further comprise: replenishing the catalyst into the dilute phase conveying bed and/or the fast fluidized bed; wherein the carbon content of the supplemented catalyst is 0 to 1.0 wt%, and may be, for example, one or more selected from the group consisting of a regenerated catalyst, a spent catalyst and a semi-regenerated catalyst. The catalyst replenished in the dilute phase conveying bed accounts for 1-50 wt%, preferably 5-30 wt% of the catalyst circulating amount in the dilute phase conveying bed; the fast fluidized bed is replenished with 1 to 50 wt%, preferably 5 to 30 wt% of the fast fluidized bed catalyst recycle. The distance between the catalyst replenishing position of the dilute phase conveying bed and the bottom of the dilute phase conveying bed can account for 0-2/3 of the total height of the dilute phase conveying bed; the distance between the catalyst replenishing position of the fast fluidized bed and the bottom of the fast fluidized bed can be 0-2/3 of the total height of the fast fluidized bed. Since the catalytic cracking reaction is a volume expansion reaction, in order to maintain the density at the inlet of the reactor and the density at the outlet of the reactor to be substantially equal or increased, the supplementary spent catalyst in the fast fluidized bed can adjust or maintain the density of the reactor within a wide range to ensure the time required for the cracking reaction. The temperature of the replenished catalyst can be adjusted according to the reaction temperature requirements, for example cold and/or hot regenerated catalyst can be introduced. In addition, the catalyst density uniformity of the reactor can be maintained as much as possible by supplementing the catalytic cracking catalyst in the dilute phase conveying bed and/or the fast fluidized bed reactor, the density distribution of the catalyst is effectively adjusted, the cracking reaction is fully and effectively carried out, and the selectivity of a target product is improved.

According to the invention, the regenerated catalyst obtained by the coke-burning regeneration of the regenerator can be cooled to 600-680 ℃ by a cooler and then returned to the bottom of the dilute phase conveying bed and/or the fast fluidized bed. The hot regenerated catalyst returns to the reactor after being cooled, which is beneficial to reducing the contact temperature of the oil agent, improving the contact state of the raw oil and the catalyst and improving the selectivity of dry gas and coke formation.

According to the present invention, the oil separating device and the reaction product separating device are well known to those skilled in the art, for example, the oil separating device may include a cyclone, a settler, a stripper, and the like, and the reaction product separating device may be a fractionating tower, and the like.

The invention will be further illustrated by means of specific embodiments in the following description with reference to the drawings, but the invention is not limited thereto.

As shown in figure 1, a pre-lifting medium enters the bottom of a pre-lifting section 2 through a pipeline 1, a regenerated catalyst from a regeneration inclined tube 11 enters the bottom of the pre-lifting section 2, and moves upwards in an accelerating manner along a rapid fluidized bed I under the lifting action of the pre-lifting medium, inferior heavy oil is injected into the lower part of the rapid fluidized bed I through a pipeline 14, is mixed and contacted with existing material flow in the rapid fluidized bed I and carries out a first catalytic cracking reaction, a first reaction product and a first catalyst to be generated enter a cyclone separator 6 in a settler 4 through an outlet section 3, the separation of the first catalyst to be generated and the first reaction product is realized, the first reaction product enters a gas collection chamber 7, and catalyst fine powder returns to the settler through a dipleg. The first spent catalyst in the settler flows to the stripping section 5. The first reaction product oil gas extracted from the first catalyst to be generated enters the gas collection chamber 7 after passing through the cyclone separator. The stripped first catalyst to be regenerated enters a regenerator 10 through a spent inclined tube 9, air is distributed by an air distributor 13 and then enters the regenerator 10, coke on the catalyst to be regenerated in a dense bed layer at the bottom of the regenerator 10 is burned off, the inactivated first catalyst to be regenerated is regenerated, and flue gas enters a subsequent energy recovery system through a flue gas pipeline 12. The reaction product oil gas in the gas collection chamber 7 enters a subsequent separation system through a large oil gas pipeline 8. Wherein the pre-lifting medium may be dry gas, water vapor or a mixture thereof.

The regenerated catalyst from a pipeline 15 enters the lower part of the dilute phase conveying bed II and moves upwards along the dilute phase conveying bed under the lifting action of a pre-lifting medium in an accelerated manner, light raw oil is injected into the bottom of the dilute phase conveying bed II through a pipeline 16, the light raw oil is subjected to a second catalytic cracking reaction on the hot regenerated catalyst, a second reaction product and a semi-spent catalyst with carbon deposit move upwards in an accelerated manner and enter a dense phase fluidized bed 18 for a third catalytic cracking reaction, the obtained third reaction product and the second spent catalyst are separated in a settler and a stripper, the third reaction product enters a subsequent separation system through a pipeline 17 and a large oil pipe line 8, and the second spent catalyst enters a regenerator for coke burning regeneration like the first spent catalyst.

The following examples further illustrate the process but do not limit the invention.

The inferior heavy oil used in the examples and comparative examples was hydrogenated residual oil, the light feedstock was catalytically cracked gasoline, and the catalyst used was a commercial catalyst sold under the trade designation DMMC-2.

Comparative example 1

The test is carried out on a medium-sized device, and the reactor is in the type of a riser series dense-phase fluidized bed which is more advanced to produce propylene at present. The preheated inferior heavy oil enters the lower part of a lifting pipe to contact with a regenerated catalyst and carry out catalytic cracking reaction, reaction oil, water vapor and a carbon deposition catalyst enter a dense-phase fluidized bed and are mixed with injected light raw oil to continue reaction, and the feeding weight ratio of the light raw oil to the inferior heavy oil is 1: 10, feeding the reacted material flow into a closed cyclone separator, quickly separating reaction oil gas and a spent catalyst, and cutting the reaction oil gas in a separation system according to a distillation range; the spent catalyst enters a stripping section under the action of gravity, hydrocarbon products adsorbed on the spent catalyst are stripped by steam, and the stripped catalyst directly enters a regenerator without heat exchange and is in contact with air for regeneration; the regenerated catalyst returns to the bottom of the riser reactor for recycling; the operating conditions and the product distribution are listed in tables 2 and 3.

As can be seen from the results of table 3, the yield of propylene was about 13.1 wt%, the yield of light aromatic hydrocarbons was about 10.1 wt%, and the yields of dry gas and coke were 11.3 wt% and 9.3 wt%, respectively.

Example 1

The test is carried out according to the flow of figure 1, the inferior raw oil is hydrogenated residual oil, and the light raw oil is catalytically cracked gasoline. The test is carried out on a medium-sized device, the reactor is divided into a first reactor (a rapid fluidized bed) and a second reactor (a dilute phase conveying bed is connected with a dense phase fluidized bed in series), preheated poor-quality raw oil enters the lower part of the rapid fluidized bed to contact with a regenerated catalyst and carry out catalytic cracking reaction, the catalyst in the rapid fluidized bed is controlled to be in full-dense phase distribution by adjusting the linear velocity of gas and arranging an umbrella-shaped gas distributor at the feeding position, and the axial solid fraction epsilon distribution in the rapid fluidized bed is all within the range of 0.1-0.2 from bottom to top. The light raw oil enters the lower part of a dilute phase conveying bed of a second reactor to contact with a regenerated catalyst and carry out catalytic cracking reaction, a reaction product and a spent catalyst enter a dense phase fluidized bed to continue the catalytic cracking reaction, and the reaction product and the spent catalyst enter a subsequent separation system. Cutting the reaction product in a separation system according to the distillation range; the spent catalyst enters a stripping section under the action of gravity, hydrocarbon products adsorbed on the spent catalyst are stripped by steam, and the stripped catalyst directly enters a regenerator without heat exchange and is in contact with air for regeneration; the regenerated catalyst returns to the dilute phase transport bed for reaction and is recycled; the operating conditions and the product distribution are listed in tables 2 and 3.

As can be seen from table 3, the yield of propylene was 21.4 wt%, the yield of light aromatics was about 11%, and the yields of dry gas and coke were 9.8% and 9.1%, respectively.

Comparative example 2

Comparative example 2 is substantially the same as example 1 except that no gas distributor is provided, the axial solid fraction epsilon distribution in the fast fluidized bed is gradually increased from top to bottom from 0.1 → 0.2 → 0.3, the operating conditions are the same as example 1, and the product distribution is shown in Table 3.

The preferred embodiments of the present invention have been described in detail, however, the present invention 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 invention within the technical idea of the present invention, and these simple modifications are within the protective scope of the present invention.

It should be noted that the various technical features described in the above embodiments can be combined in any suitable manner without contradiction, and the invention is not described in any way for the possible combinations in order to avoid unnecessary repetition.

In addition, any combination of the various embodiments of the present invention is also possible, and the same should be considered as the content of the present invention as long as it does not depart from the gist of the present invention.

TABLE 1

Raw materials	CatalysisCracked gasoline	Hydrogenated residual oil
			Density (20 deg.C)/g.cm^-3	0.7179	0.9342
Refractive index/70 deg.C	1.4158	1.5075
			Basic nitrogen/microgram g^-1	/	767
Carbon residue/weight%	/	5.58
			Value of characteristic factor K	/	11.6
Distillation range/. degree.C
			5% by volume	/	365
10% by volume	55	403
			30% by volume	67	479
50% by volume	89	545
			70% by volume	119	617
90% by volume	152	/
			End point of distillation	180	/
Metal content/microgram g^-1
			Fe	/	26.3
Ni	/	9
			Ca	/	5.5
V	/	8
			Na	/	1.2

TABLE 2

	Comparative example 1	Example 1
			Conditions of fast fluidized bed
Outlet temperature of	/	555
			Reaction time in seconds	/	5
Weight ratio of solvent to oil	/	0.25
			Water oilWeight ratio of	/	10
Catalyst density in kg/m³	/	200
			Linear velocity of gas, m/s		2
Reaction pressure, kPa		210
			Gs, kg/(meter)²Second)		80
Riser/dilute phase transport bed conditions
			Outlet temperature of	565	620
Reaction time in seconds	2	1.5
			Water to oil weight ratio	0.25	0.05
Weight ratio of solvent to oil	10	15
			Catalyst density in kg/m³	60	50
Linear velocity of gas, m/s	12	15
			Reaction pressure, kPa	210	210
Gs, kg/(meter)²Second)	320	300

TABLE 3

Dense-phase fluidized bed	Comparative example 1	Example 1	Comparative example 2
				Bed temperature of low DEG C	555	605	/
Weight speed airspeed hourly^-1	4	4	/
				Catalyst density in kg/m³	460	480	/
Linear velocity of gas, m/s	0.6	0.6	/
				Product distribution, weight%
Dry gas	11.3	9.8	10.3
				Liquefied gas	27.0	41.5	36.1
Wherein propylene is	13.1	21.4	17.9
				Gasoline (gasoline)	30.4	26.2	27.5
Wherein the light aromatic hydrocarbons	9.9	11	10.6
				Diesel oil	14.6	10.4	11.3
Heavy oil	7.4	3	5.6
				Coke	9.3	9.1	9.2
Total up to	100.0	100.0	100.0

Claims

1. A catalytic cracking process for producing propylene and light aromatic hydrocarbons, the process comprising:

2. The method of claim 1 wherein the properties of the low quality heavy oil meet one, two, three or four of the following criteria: the density at 20 ℃ is 910-³Carbon residue of 3-8 wt%, total of Ni and VThe content is 5-20ppm, and the characteristic factor K value is less than 12.0.

3. The method of claim 1 wherein the low quality heavy oil is heavy petroleum hydrocarbons and/or other mineral oils; the heavy petroleum hydrocarbon is one or more selected from vacuum residue, poor atmospheric residue, poor hydrogenated residue, coker gas oil, deasphalted oil, vacuum wax oil, high acid value crude oil and high metal crude oil, and the other mineral oil is one or more selected from coal liquefied oil, oil sand oil and shale oil.

4. The process of claim 1 wherein the first catalytic cracking catalyst comprises, on a dry basis and based on the weight of the first catalytic cracking catalyst on a dry basis, from 1 to 50 wt% zeolite, from 5 to 99 wt% inorganic oxide, and from 0 to 70 wt% clay;

5. The process of claim 4, wherein the medium pore zeolite of the first catalytic cracking catalyst comprises from 0 to 50 wt% of the total weight of zeolite on a dry basis;

6. The process of claim 4, wherein the medium pore zeolite of the first catalytic cracking catalyst comprises from 0 to 20 wt% of the total weight of zeolite on a dry basis;

7. The method of claim 1, wherein the conditions of the first catalytic cracking reaction comprise: the reaction temperature is 550-700 ℃, the reaction time is 0.1-5 seconds, and the weight ratio of the catalyst to the oil is (5-30): 1, the weight ratio of water to oil is (0.05-2): 1, the catalyst density is 20-100 kg/m³Gas linear speed of 4-18 m/s, reaction pressure of 0.2-1.2 MPa, and catalyst mass flow rate G_sIs 180-500 kg/(meter)²Seconds).

8. The method of claim 1, wherein the conditions of the second catalytic cracking reaction comprise: the reaction temperature is 510-650 ℃, and the weight hourly space velocity is 1-20 h^-1The weight ratio of the agent oil is (5-30): 1, catalyst density of 300-³The gas linear speed is 0.4-0.8 m/s, and the reaction pressure is 130-450 MPa.

9. The method of claim 1, wherein the conditions of the third catalytic cracking reaction comprise: the reaction temperature is 510-650 ℃, the reaction time is 1-20 seconds, and the weight ratio of the catalyst to the oil is (3-50): 1, the weight ratio of water to oil is (0.03-0.8): 1, catalyst density of 120-³Gas linear speed of 0.8-2.5 m/s, reaction pressure of 130-450 kPa, and catalyst mass flow rate G_sIs 15-150 kg/(meter)²Seconds).

10. The method of claim 1, wherein the conditions of the third catalytic cracking reaction comprise: the reaction temperature is 550-620 ℃, the reaction time is 3-15 seconds, and the weight ratio of the catalyst to the oil is (10-30): 1, the weight ratio of water to oil is (0.05-0.5): 1, catalyst density of 150-³Gas linear velocity of 1-1.8 m/s, catalyst mass flow rate G_sIs 20-130 kg/(meter)²Seconds).

11. The method of claim 1, further comprising: replenishing the catalyst into the dilute phase conveying bed and/or the fast fluidized bed; wherein the carbon content of the supplemented catalyst is 0-1.0 wt.%.

12. The process of claim 11 wherein the dilute phase transport bed is replenished with catalyst in the range of from 0 to 50 wt.% of the dilute phase transport bed catalyst circulation; the replenished catalyst of the fast fluidized bed accounts for 0-50 wt% of the circulating amount of the fast fluidized bed catalyst.

13. The process of claim 11 wherein the dilute phase transport bed is replenished with catalyst in the range of 5 to 30 wt.% of the dilute phase transport bed catalyst circulation; the replenished catalyst of the fast fluidized bed accounts for 5-30 wt% of the circulating amount of the fast fluidized bed catalyst.

14. The process of claim 11 wherein the dilute phase transport bed has a catalyst makeup location that is spaced from the bottom of the dilute phase transport bed by a distance that is between 0 and 2/3 of the total height of the dilute phase transport bed;

15. The method as claimed in claim 1, wherein the regenerated catalyst obtained by coke-burning regeneration of the regenerator is cooled to 680 ℃ by a cooler and then returned to the bottom of the dilute phase transport bed and/or the fast fluidized bed.

16. A catalytic cracking system comprises a dilute phase conveying bed, a dense phase fluidized bed, a fast fluidized bed, optional oil agent separation equipment, reaction product separation equipment and a regenerator, wherein a settler which is communicated with fluid is arranged above the dense phase fluidized bed, and a stripping section which is communicated with fluid is arranged below the dense phase fluidized bed;

catalyst inlets of the dilute phase conveying bed and the fast fluidized bed are both communicated with a catalyst outlet of the regenerator in a fluid mode, the dilute phase conveying bed extends into an oil agent inlet of the dense phase fluidized bed from bottom to top, an oil agent outlet of the dilute phase conveying bed is located in the dense phase fluidized bed, a reaction product outlet of the settler is communicated with an oil agent inlet of the oil agent separation equipment in a fluid mode, an oil agent outlet of the fast fluidized bed is communicated with the oil agent inlet of the settler in a fluid mode or communicated with the oil agent inlet of the oil agent separation equipment in a fluid mode, a reaction product outlet of the oil agent separation equipment is communicated with a reaction product inlet of the reaction product separation equipment in a fluid mode, and a catalyst outlet of the stripping section is communicated with a catalyst inlet of the regenerator in a fluid mode;

the catalyst in the fast fluidized bed is distributed in a full dense phase, and the distribution of the axial solid fraction epsilon in the fast fluidized bed meets the following requirements: epsilon is more than or equal to 0.1 and less than or equal to 0.2; the distance between the catalyst replenishing position of the dilute phase conveying bed and the bottom of the dilute phase conveying bed accounts for 0-2/3 of the total height of the dilute phase conveying bed; and/or the distance between the catalyst replenishing position of the fast fluidized bed and the bottom of the fast fluidized bed accounts for 0-2/3 of the total height of the fast fluidized bed; the catalyst is adjusted to be in full dense phase distribution by adjusting the linear velocity of gas in the fast fluidized bed and arranging a gas distributor at the feeding position of the fast fluidized bed.