阅读说明：本技术 一种采用稀相输送床和湍流流化床进行催化裂解的方法和系统 (Method and system for catalytic cracking by adopting dilute phase conveying bed and turbulent fluidized bed ) 是由 张执刚 龚剑洪 魏晓丽 于 2018-07-16 设计创作，主要内容包括：本发明涉及一种采用稀相输送床和湍流流化床进行催化裂解的方法和系统,该方法包括：将预热的劣质重油从稀相输送床的下部引入稀相输送床中与催化裂解催化剂接触并由下至上进行第一催化裂解反应,得到第一反应产物和半待生催化剂；将所得第一反应产物和半待生催化剂引入湍流流化床的底部并由下至上进行第二催化裂解反应,得到第二反应产物和待生催化剂；将所得第二反应产物进行分离,得到干气、液化气、裂解汽油、裂解柴油和裂解重油；将待生催化剂送入再生器进行烧焦再生,至少将部分所得再生催化剂作为所述催化裂解催化剂返回稀相输送床的底部。本发明方法和系统进行催化裂解的干气和焦炭产率低,产品分布好。(The invention relates to a method and a system for catalytic cracking by adopting a dilute phase conveying bed and a turbulent fluidized bed, wherein the method comprises the following steps: introducing preheated poor-quality heavy oil into the dilute-phase conveying bed from the lower part of the dilute-phase conveying bed to contact with a catalytic cracking catalyst and perform a first catalytic cracking reaction from bottom to top to obtain a first reaction product and a semi-spent catalyst; introducing the obtained first reaction product and the semi-spent catalyst into the bottom of a turbulent fluidized bed and carrying out a second catalytic cracking reaction from bottom to top to obtain a second reaction product and a spent catalyst; separating the obtained second reaction product to obtain dry gas, liquefied gas, pyrolysis gasoline, pyrolysis diesel oil and pyrolysis heavy oil; and (3) feeding the spent catalyst into a regenerator for coke burning regeneration, and returning at least part of the obtained regenerated catalyst serving as the catalytic cracking catalyst to the bottom of the dilute phase conveying bed. The method and the system have the advantages of low yield of dry gas and coke for catalytic cracking and good product distribution.)

1. A process for catalytic cracking using a dilute phase transport bed and a turbulent fluidized bed, the process comprising:

introducing preheated poor-quality heavy oil into the dilute-phase conveying bed from the lower part of the dilute-phase conveying bed to contact with a catalytic cracking catalyst and perform a first catalytic cracking reaction from bottom to top to obtain a first reaction product and a semi-spent catalyst;

the obtained first reaction product and a semi-spent catalystIntroducing the mixture into the bottom of the turbulent fluidized bed and carrying out a second catalytic cracking reaction from bottom to top to obtain a second reaction product and a spent catalyst; wherein the density of the catalyst in the turbulent fluidized bed is 300-450 kg/m³；

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

and (3) feeding the spent catalyst into a regenerator for coke burning regeneration, and returning at least part of the obtained regenerated catalyst serving as the catalytic cracking catalyst to the bottom of the dilute phase conveying bed.

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 900-³The carbon residue is 2-10 wt%, the total content of nickel and vanadium is 2-30ppm, and the characteristic factor K value is less than 12.1.

3. 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 910-³The 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.

4. 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.

5. The process of claim 1 wherein the catalytic cracking catalyst comprises, on a dry basis and based on the weight of the 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;

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.

6. A process as claimed in claim 5, wherein the medium pore size zeolite comprises from 0 to 50 wt% of the total weight of zeolite on a dry basis.

7. A process as claimed in claim 5, wherein the medium pore size zeolite comprises from 0 to 20 wt% of the total weight of zeolite on a dry basis.

8. The method of claim 1, wherein the conditions of the first catalytic cracking reaction comprise: the reaction temperature is 500-600 ℃, the reaction time is 0.05-5 seconds, and the weight ratio of the catalyst to the oil is (1-50): 1, the weight ratio of water to oil is (0.03-0.5): 1, the catalyst density is 20-100 kg/m³Gas linear speed of 4-18 m/s, reaction pressure of 130-450 kPa, and catalyst mass flow rate G_sIs 180-500 kg/(meter)²Seconds);

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 (3-60): 1, the weight ratio of water to oil is (0.03-0.8): 1, catalyst density of 320-440 kg/m³The gas linear speed is 0.4-0.8 m/s, and the reaction pressure is 130-450 kPa.

9. The method of claim 1, wherein the conditions of the first catalytic cracking reaction comprise: the reaction temperature is 520-580 ℃, the reaction time is 1-3 seconds, and the weight ratio of the catalyst to the oil is (5-25): 1, the weight ratio of water to oil is (0.05-0.3): 1;

the conditions of the second catalytic cracking reaction include: the reaction temperature is 550-620 ℃, and the weight ratio of the catalyst to the oil is (10-40): 1, the weight ratio of water to oil is (0.05-0.5): 1, the gas linear speed is 0.6-0.9 m/s.

10. The method of claim 1, further comprising: introducing a C4 hydrocarbon fraction and/or a C5-C6 light gasoline fraction into the dilute phase transport bed and/or turbulent fluidized bed for catalytic cracking reaction.

11. The process of claim 10 wherein said C4 hydrocarbon fraction and/or C5-C6 light gasoline fraction are introduced before the introduction of the low grade heavy oil into the dilute phase transport bed at the feed location; and/or

The C4 hydrocarbon fraction and/or the C5-C6 light gasoline fraction are introduced into the bottom of a turbulent fluidized bed.

12. The method of claim 1, wherein the method further comprises: supplementing a catalyst into the turbulent fluidized bed to perform the second catalytic cracking reaction together with the first reaction product and the semi-spent catalyst; wherein the carbon content of the supplemented catalyst is 0 to 1.0 wt.%.

13. The process of claim 12 wherein the make-up catalyst comprises from 0 to 50 wt% of the total circulating amount of dilute phase transport bed and turbulent fluidized bed catalyst.

14. The process of claim 12 wherein the make-up catalyst comprises from 5 to 30 wt% of the total circulating amount of dilute phase transport bed and turbulent fluidized bed catalyst.

15. A catalytic cracking system comprises a dilute phase fluidized bed, a turbulent fluidized bed, an oil agent separation device, a reaction product separation device and a regenerator;

said dilute phase fluidized bed being in fluid communication with a turbulent fluidized bed and said dilute phase fluidized bed being upstream of said turbulent fluidized bed in the direction of flow of the reactant materials;

the dilute phase conveying bed is provided with a catalyst inlet at the bottom and an inferior heavy oil inlet at the lower part, the turbulent fluidized bed is provided with an oil agent outlet at the top, the oil agent separation equipment is provided with an oil agent 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;

the catalyst inlet of the dilute phase conveying bed is in fluid communication with the catalyst outlet of the regenerator, the oil agent outlet of the turbulent fluidized bed is in fluid communication with the oil agent inlet of the oil agent separation device, the reaction product outlet of the oil agent separation device is in fluid communication with the reaction product inlet of the reaction product separation device, and the catalyst outlet of the oil agent separation device is in fluid communication with the catalyst inlet of the regenerator.

16. The system of claim 15, wherein the turbulent fluidized bed is coaxially disposed above and below the dilute phase transport bed and the turbulent fluidized bed is above the dilute phase transport bed.

17. The system of claim 15, wherein the ratio of the inner diameter of the dilute phase transport bed to the inner diameter of the turbulent fluidized bed is (0.2-0.7): 1.

Technical Field

The invention relates to a method and a system for catalytic cracking by adopting a dilute phase conveying bed and a turbulent fluidized bed.

Background

The low-carbon olefin represented by ethylene and propylene is the most basic raw material in chemical industry, and natural gas or light petroleum fraction is mostly used as raw material at home and abroad, and the low-carbon olefin is produced by adopting a steam cracking process in an ethylene combined device. Benzene, toluene, and xylene (BTX) are important basic chemical raw materials, wherein para-xylene (PX) accounts for about 45% of the total BTX consumption. 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%.

Chinese patent CN01119807.9 discloses a method for increasing the yield of ethylene and propylene by catalytic conversion of heavy petroleum hydrocarbon, which is to make 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 a second riser for further reaction, and the catalyst is a mixture of USY molecular sieve and ZSM-5 molecular sieve.

U.S. Pat. Nos. 2002195373 and WO2017223310 use a downflow reactor operating at high temperature (1020-1200 ℃ F./550-650 ℃ C.), short contact time (<0.5 seconds) and large oil-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 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 CN 01130984.9 discloses a catalytic conversion method for preparing ethylene and propylene. A riser and a dense-phase fluidized bed reactor are connected in series, light raw materials are injected into the riser to react under higher severity, reaction products and carbon deposition catalyst enter a fluidized bed to continuously react under relatively mild conditions, and the method has higher yield of ethylene, propylene and butylene.

Chinese patent CN200910210330 discloses a catalytic cracking method, comprising the steps of contacting a heavy raw material with a catalyst in a first riser reactor comprising at least two reaction zones to perform a cracking reaction, and contacting a light raw material and cracked heavy oil with the catalyst in a second riser reactor and a fluidized bed reactor to perform the 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.

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. Catalytic cracking devices used as bridges for oil refining and chemical engineering face unprecedented pressure and challenge. 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 method and a system for catalytic cracking by adopting a dilute phase conveying bed and a turbulent fluidized bed.

In order to achieve the above object, the present invention provides a method for catalytic cracking using a dilute phase transport bed and a turbulent fluidized bed, the method comprising:

introducing preheated poor-quality heavy oil into the dilute-phase conveying bed from the lower part of the dilute-phase conveying bed to contact with a catalytic cracking catalyst and perform a first catalytic cracking reaction from bottom to top to obtain a first reaction product and a semi-spent catalyst;

introducing the obtained first reaction product and the semi-spent catalyst into the bottom of a turbulent fluidized bed and carrying out a second catalytic cracking reaction from bottom to top to obtain a second reaction product and a spent catalyst; wherein the density of the catalyst in the turbulent fluidized bed is 300-450 kg/m³；

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

and (3) feeding the spent catalyst into a regenerator for coke burning regeneration, and returning at least part of the obtained regenerated catalyst serving as the catalytic cracking catalyst to the bottom of the dilute phase conveying bed.

Optionally, the properties of the inferior heavy oil meet one, two, three or four of the following criteria: the density at 20 ℃ is 900-³The carbon residue is 2-10 wt%, the total content of nickel and vanadium is 2-30ppm, and the characteristic factor K value is less than 12.1.

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 catalytic cracking catalyst comprises, on a dry basis and based on the weight of the 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;

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 comprises from 0 to 50 wt% of the total weight of the zeolite on a dry basis.

Optionally, the medium pore zeolite comprises from 0 to 20 wt% of the total weight of the zeolite on a dry basis.

Optionally, the conditions of the first catalytic cracking reaction include: the reaction temperature is 500-600 ℃, the reaction time is 0.05-5 seconds, and the weight ratio of the catalyst to the oil is (1-50): 1, the weight ratio of water to oil is (0.03-0.5): 1, the catalyst density is 20-100 kg/m³Gas linear speed of 4-18 m/s, reaction pressure of 130-450 kPa, and catalyst mass flow rate G_sIs 180-500 kg/(meter)²Seconds);

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 (3-60): 1, the weight ratio of water to oil is (0.03-0.8): 1, catalyst density of 320-440 kg/m³The gas linear speed is 0.4-0.8 m/s, and the reaction pressure is 130-450 kPa.

Optionally, the conditions of the first catalytic cracking reaction include: the reaction temperature is 520-580 ℃, the reaction time is 1-3 seconds, and the weight ratio of the catalyst to the oil is (5-25): 1, the weight ratio of water to oil is (0.05-0.3): 1;

the conditions of the second catalytic cracking reaction include: the reaction temperature is 550-620 ℃, and the weight ratio of the catalyst to the oil is (10-40): 1, the weight ratio of water to oil is (0.05-0.5): 1, the gas linear speed is 0.6-0.9 m/s.

Optionally, the method further includes: introducing a C4 hydrocarbon fraction and/or a C5-C6 light gasoline fraction into the dilute phase transport bed and/or turbulent fluidized bed for catalytic cracking reaction.

Optionally, introducing the C4 hydrocarbon fraction and/or the C5-C6 light gasoline fraction before introducing the low grade heavy oil into the dilute phase transport bed at the feed location; and/or

The C4 hydrocarbon fraction and/or the C5-C6 light gasoline fraction are introduced into the bottom of a turbulent fluidized bed.

Optionally, the method further includes: supplementing a catalyst into the turbulent fluidized bed to perform the second catalytic cracking reaction together with the first reaction product and the semi-spent catalyst; wherein the carbon content of the supplemented catalyst is 0 to 1.0 wt.%.

Optionally, the make-up catalyst comprises from 0 to 50 wt% of the total catalyst circulation of the dilute phase transport bed and the turbulent fluidized bed.

Optionally, the make-up catalyst comprises 5 to 30 wt% of the total catalyst circulation of the dilute phase transport bed and the turbulent fluidized bed.

The invention also provides a catalytic cracking system, which comprises a dilute phase fluidized bed, a turbulent fluidized bed, an oil agent separation device, a reaction product separation device and a regenerator;

said dilute phase fluidized bed being in fluid communication with a turbulent fluidized bed and said dilute phase fluidized bed being upstream of said turbulent fluidized bed in the direction of flow of the reactant materials;

the dilute phase conveying bed is provided with a catalyst inlet at the bottom and an inferior heavy oil inlet at the lower part, the turbulent fluidized bed is provided with an oil agent outlet at the top, the oil agent separation equipment is provided with an oil agent 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;

the catalyst inlet of the dilute phase conveying bed is in fluid communication with the catalyst outlet of the regenerator, the oil agent outlet of the turbulent fluidized bed is in fluid communication with the oil agent inlet of the oil agent separation device, the reaction product outlet of the oil agent separation device is in fluid communication with the reaction product inlet of the reaction product separation device, and the catalyst outlet of the oil agent separation device is in fluid communication with the catalyst inlet of the regenerator.

Optionally, the turbulent fluidized bed and the dilute phase conveying bed are coaxially arranged up and down, and the turbulent fluidized bed is located above the dilute phase conveying bed.

Optionally, the ratio of the internal diameter of the dilute phase transport bed to the internal diameter of the turbulent fluidized bed is (0.2-0.7): 1.

the invention adopts the turbulent fluidized bed with specific catalyst density to effectively control the density of the reaction catalyst, thereby greatly improving the ratio of the instantaneous catalyst in the reactor to the raw oil compared with a dilute phase conveying bed, controlling relatively longer reaction time, leading the catalyst to be capable of fully reacting with inferior heavy oil, improving the reaction conversion rate and the yield of low-carbon olefin and light aromatic hydrocarbon compared with the conventional dense phase fluidized bed (the density of the catalyst is more than 450 kg/m)³The density of the catalyst in the dense fluidized bed of DCC is generally greater than 500 kg/m³) Compared with the prior art, the method can effectively reduce the generation of dry gas and coke, and improve 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 dilute phase conveying bed II turbulent fluidized bed

1 pipeline 2 pre-lift 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 make-up line

16 pipeline

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 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。

(2) 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。

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

(4) 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 method for catalytic cracking by adopting a dilute phase conveying bed and a turbulent fluidized bed, which comprises the following steps:

introducing preheated poor-quality heavy oil into the dilute-phase conveying bed from the lower part of the dilute-phase conveying bed to contact with a catalytic cracking catalyst and perform a first catalytic cracking reaction from bottom to top to obtain a first reaction product and a semi-spent catalyst;

introducing the obtained first reaction product and the semi-spent catalyst into the bottom of a turbulent fluidized bed and carrying out a second catalytic cracking reaction from bottom to top to obtain a second reaction product and a spent catalyst; wherein the density of the catalyst in the turbulent fluidized bed is 300-450 kg/m³；

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

and (3) feeding the spent catalyst into a regenerator for coke burning regeneration, and returning at least part of the obtained regenerated catalyst serving as the catalytic cracking catalyst to the bottom of the dilute phase conveying bed.

According to the invention, the inferior heavy oil can be introduced into the dilute phase transport bed at one feeding position, or can be introduced into the dilute phase transport 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: the density at 20 ℃ is 900-³Preferably 910-³Carbon residue of 2 to 10 wt.%, preferably 3 to 8 wt.%, a total nickel and vanadium content of 2 to 30ppm, preferably 5 to 20ppm, and a characteristic factor K value of less than 12.1, preferably less than 12.0. 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 may include, for the process of the present invention, from 1 to 50 wt% zeolite, from 5 to 99 wt% inorganic oxide, and from 0 to 70 wt% clay, on a dry basis and based on the weight of the catalytic cracking catalyst on a dry basis; the zeolite may comprise, as active components, a medium pore zeolite and optionally a large pore zeolite, and the medium pore zeolite in the catalytic cracking catalyst may comprise from 0 to 50 wt%, preferably from 0 to 20 wt%, of the total weight of the 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 preferably being selected from one or more mixtures of ZSM-5, ZSM-11, ZSM-12, ZSM-23, ZSM-35, ZSM-38, ZSM-48 and other zeolites of similar structure, as described more fully in connection with ZSM-5, see U.S. Pat. No. 3,702,886. 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.

Catalytic cracking is well known to those skilled in the art in accordance with the present invention, and for the catalytic cracking of dilute phase transport beds and turbulent fluidized beds of the present application, the conditions of the first catalytic cracking reaction may include: the reaction temperature is 500-600 ℃, the reaction time is 0.05-5 seconds, and the weight ratio of the catalyst to the oil is (1-50): 1, the weight ratio of water to oil is (0.03-0.5): 1, the catalyst density is 20-100 kg/m³Gas linear speed of 4-18 m/s, reaction pressure of 130-450 kPa, and catalyst mass flow rate G_sIs 180-500 kg/(meter)²Seconds); the conditions of the first catalytic cracking reaction preferably include: the reaction temperature is 520-580 ℃, the reaction time is 1-3 seconds, and the weight ratio of the catalyst to the oil is (5-25): 1, the weight ratio of water to oil is (0.05-0.3): 1; the conditions of the second catalytic cracking reaction may include: the reaction temperature is 510-650 ℃, the weight ratio of the catalyst to the oil is (3-60): 1, the weight ratio of water to oil is (0.03-0.8): 1, catalyst density of 300-450 kg/m³Preferably 320-440 kg/m³Further preferably 320-400 kg/m³More preferably 330-³The gas linear speed is 0.4-0.8 m/s, and the reaction pressure is 130-450 kPa; the conditions of the second catalytic cracking reaction preferably include: the reaction temperature is 550-620 ℃, and the weight ratio of the catalyst to the oil is (10-40): 1, the weight ratio of water to oil is (0.05-0.5): 1, the gas linear speed is 0.6-0.9 m/s.

The separation of the second reaction product 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 second reaction product 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.

According to the invention, the method also preferably comprises: introducing a C4 hydrocarbon fraction and/or a C5-C6 light gasoline fraction into the dilute phase transport bed and/or turbulent fluidized bed for catalytic cracking reaction. The C4 hydrocarbon fraction refers to low molecular weight hydrocarbons containing C4 fraction as main component and existing in gas form at normal temperature and pressure, including alkanes, alkenes and alkynes with 4 carbon atoms in the molecule, which may include gaseous hydrocarbon products (such as liquefied gas) produced by the present invention method and rich in C4 hydrocarbon fraction, and may also include gaseous hydrocarbons produced by other devices and rich in C4 fraction, wherein C4 hydrocarbon fraction produced by the present invention method is preferred. The C4 hydrocarbon fraction is preferably an olefin-rich C4 hydrocarbon fraction, and the content of C4 olefins may be greater than 50 wt%, preferably greater than 60 wt%, more preferably above 70 wt%. The C5-C6 light gasoline fraction can comprise pyrolysis gasoline produced by the method of the invention, and can also comprise gasoline fractions produced by other devices, such as at least one C5-C6 fraction selected from catalytic pyrolysis gasoline, catalytic cracking gasoline, straight run gasoline, coker gasoline, thermal cracking gasoline and hydrogenated gasoline. The present invention preferably introduces said C4 hydrocarbon fraction and/or C5-C6 light gasoline fraction prior to introducing the low grade heavy oil into the dilute phase transport bed at the feed location; and/or preferably the C4 hydrocarbon fraction and/or the C5-C6 light gasoline fraction is introduced into the bottom of the turbulent fluidized bed.

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 further comprises: supplementing a catalyst into the turbulent fluidized bed to perform the second catalytic cracking reaction together with the first reaction product and the semi-spent catalyst; wherein the carbon content of the supplemented catalyst may be 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 make-up catalyst may be present in the range of from 0 to 50 wt%, preferably from 5 to 30 wt%, of the total amount of catalyst circulated through the dilute phase transport bed and the turbulent fluidized bed, and the point of make-up of the turbulent fluidized bed where the catalyst is present may be the bottom of the turbulent fluidized bed. Since the catalytic cracking reaction is a volume expansion reaction, in order to maintain a substantially equivalent or increased density of the inlet and outlet regions of the reactor, the additional spent catalyst in the turbulent fluidized bed can be adjusted or maintained over a wide range to ensure the time required for the cracking reaction. Meanwhile, more 'fresh' active centers are provided for the cracking reaction, the flexibility of adjusting the reaction temperature is enhanced, and the gradient of the temperature and the catalyst activity in the turbulent fluidized bed is obviously improved. In addition, the catalytic cracking catalyst is supplemented in the turbulent fluidized bed reactor, so that the density uniformity of the catalyst in the reactor can be maintained as much as possible, the density distribution of the catalyst is effectively adjusted, the cracking reaction is ensured to be fully and effectively carried out, and the selectivity of a target product is improved.

The invention also provides a catalytic cracking system, which comprises a dilute phase fluidized bed, a turbulent fluidized bed, an oil agent separation device, a reaction product separation device and a regenerator;

said dilute phase fluidized bed being in fluid communication with a turbulent fluidized bed and said dilute phase fluidized bed being upstream of said turbulent fluidized bed in the direction of flow of the reactant materials;

the dilute phase conveying bed is provided with a catalyst inlet at the bottom and an inferior heavy oil inlet at the lower part, the turbulent fluidized bed is provided with an oil agent outlet at the top, the oil agent separation equipment is provided with an oil agent 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;

the catalyst inlet of the dilute phase conveying bed is in fluid communication with the catalyst outlet of the regenerator, the oil agent outlet of the turbulent fluidized bed is in fluid communication with the oil agent inlet of the oil agent separation device, the reaction product outlet of the oil agent separation device is in fluid communication with the reaction product inlet of the reaction product separation device, and the catalyst outlet of the oil agent separation device is in fluid communication with the catalyst inlet of the regenerator.

According to the invention, the dilute phase transport bed is located upstream of the turbulent fluidized bed, in the present invention both "upstream" and "downstream" of the reactor, i.e. at the bottom or lower part of the reactor, with respect to the direction of flow of the reaction mass. Preferably, the turbulent fluidized bed and the dilute phase conveying bed are coaxially arranged up and down, the turbulent fluidized bed can be positioned above the dilute phase conveying bed, and the ratio of the inner diameter of the dilute phase conveying bed to the inner diameter of the turbulent fluidized bed can be (0.2-0.7): 1.

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 FIG. 1, a pre-lifting medium, which may be dry gas, steam or a mixture thereof, enters the bottom of the dilute-phase transport bed I from the bottom of the pre-lifting section 2 through a line 1. The regenerated catalyst from the regenerated inclined tube 11 enters the bottom of the dilute phase conveying bed I, moves upwards in an accelerated manner along the dilute phase conveying bed I under the lifting action of a pre-lifting medium, inferior heavy oil is injected into the lower part of the dilute phase conveying bed I through a pipeline 14, is mixed and contacted with the existing material flow in the dilute phase conveying bed and carries out a first catalytic cracking reaction, a first reaction product and a semi-spent catalyst move upwards in an accelerated manner, enters the turbulent flow fluidized bed II, is contacted with the supplemented regenerated catalyst from a supplement line 15 and carries out a second catalytic cracking reaction, a generated second reaction product and an inactivated spent catalyst enter the cyclone separator 6 in the settler 4 through the outlet section 3 to realize the separation of the spent catalyst and the second reaction product, the second reaction product enters the gas collection chamber 7, and catalyst fine powder returns to the settler through a dipleg. Spent catalyst in the settler flows to the stripping section 5. The reaction product stripped from the spent catalyst enters a gas collection chamber 7 after passing through a cyclone separator, and the reaction product in the gas collection chamber 7 enters a subsequent separation system through a large oil-gas pipeline 8. The stripped spent catalyst 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 spent catalyst in a dense bed layer at the bottom of the regenerator 10 is burned off, the inactivated spent catalyst is regenerated, and flue gas enters a subsequent energy recovery system through a flue gas pipeline 12.

The following examples further illustrate the process of the present invention but are not intended to limit the invention thereto.