A method and system for the production of light olefins and aromatics by catalytic cracking in a cyclonic reactor

Abstract

The application provides a method for producing low-carbon olefins and aromatic hydrocarbons by catalytic cracking, which comprises: in a tubular fluidized bed reactor, carrying out catalytic cracking reaction on reaction materials containing fluidizing medium, catalyst and raw oil in a state of rotational fluidization, then carrying out gas-solid separation and stripping to obtain spent catalyst and reaction product oil gas, and returning the spent catalyst to the catalytic cracking reaction after regeneration. The application also provides a system for producing low-carbon olefins and aromatic hydrocarbons by catalytic cracking. Through the above technical scheme, the application strengthens the gas-solid contact between the raw oil and the catalyst through rotational fluidization, thereby effectively improving the conversion rate and selectivity of the catalytic cracking for producing low-carbon olefins and aromatic hydrocarbons.

Description

Technical Field

[0001] The present invention relates to the field of petrochemical industry, in particular to a method for producing light olefins and aromatics by catalytic cracking in a cyclone reactor and a system for producing light olefins and aromatics by catalytic cracking. Background Art

[0002] With the upgrading of gasoline and diesel quality and shifting demand structures, the diesel-to-gasoline ratio continues to decline, leading to an increasing problem of diesel surplus in refineries. Light olefins and aromatics are important organic chemical raw materials, with strong domestic demand and a high degree of external dependence. The combination of overcapacity in refining and a shortage of chemical raw materials has made the transition from refining to chemical processing an inevitable trend.

[0003] Catalytic cracking is a process in which heavy crude oil undergoes cracking reaction under the action of heat and catalyst, converting it into cracked gas, gasoline and diesel, etc. It is an important means of using refined products to produce light olefins and aromatics.

[0004] CN101362963A, CN101747928A, and CN1667089A disclose methods for producing more propylene and aromatics by catalytic cracking of heavy feedstocks. By recycling distillate oils such as difficult-to-crack feedstocks or recycled cracking feedstocks, the propylene yield can reach over 40% by weight, and BTX is extracted from the aromatics-rich fraction using aromatics extraction technology.

[0005] However, the conversion rate and selectivity of existing catalytic cracking methods still need to be further improved. Summary of the Invention

[0006] The purpose of the present invention is to further improve the conversion rate and selectivity of catalytic cracking to produce light olefins and aromatics.

[0007] To achieve the above-mentioned object, the present invention provides a method for producing light olefins and aromatics by catalytic cracking, the method comprising: in a tubular fluidized bed reactor, subjecting a reaction material containing a fluidizing medium, a catalyst and a feedstock oil to a catalytic cracking reaction under a rotating fluidized state, then performing gas-solid separation and steam stripping to obtain a spent catalyst and reaction product oil and gas, and then regenerating the spent catalyst and returning it to the catalytic cracking reaction.

[0008] The present invention also provides a system for producing light olefins and aromatics by catalytic cracking, the system comprising a tubular fluidized bed reactor, a gas-solid separation device, a stripper and a regenerator connected in sequence, the discharge port of the regenerator being connected to the catalyst inlet of the fluidized bed reactor; wherein a cyclone fluidization device is provided in the fluidized bed reactor.

[0009] Through the above technical solution, the present invention strengthens the gas-solid contact between the feed oil and the catalyst through swirl fluidization, thereby effectively improving the conversion rate and selectivity of catalytic cracking to produce light olefins and aromatics.

[0010] Other features and advantages of the present invention will be described in detail in the following detailed description. BRIEF DESCRIPTION OF THE DRAWINGS

[0011] The accompanying drawings are used to provide a further understanding of the present invention and constitute a part of the specification. Together with the following detailed description, they are used to explain the present invention but do not constitute a limitation of the present invention. In the accompanying drawings:

[0012] Figure 1 This is a schematic diagram of the system structure in a preferred embodiment of the present invention.

[0013] Figure 2 This is a schematic structural diagram of a swirl distribution plate in a preferred embodiment of the present invention.

[0014] Figure 3 for Figure 2 Schematic diagram of the structure of the distribution blades in the swirl distribution plate.

[0015] Description of Reference Numerals

[0016] Figure 1 In the figure, 1 is the fluidizing medium; 2 is the swirl distribution plate; 3 is the slide valve to be regenerated; 4 is the oil inlet nozzle; 5 is the riser reactor; 6 is the cyclone separator; 7 is the stripper; 8 is the product oil and gas discharge pipeline; 9 is the stripping steam; 10 is the regeneration slide valve; 11 is the regenerator; 12 is the regenerator cyclone separator; 13 is the flue gas outlet pipeline; and 14 is the regeneration medium. DETAILED DESCRIPTION

[0017] The following is a detailed description of the specific embodiments of the present invention. It should be understood that the specific embodiments described herein are only used to illustrate and explain the present invention and are not intended to limit the present invention.

[0018] The present invention provides a method for producing light olefins and aromatics by catalytic cracking, which comprises: in a tubular fluidized bed reactor, subjecting reaction materials containing a fluidizing medium, a catalyst and a feedstock oil to a catalytic cracking reaction under a rotating fluidized state, then performing gas-solid separation and steam stripping to obtain a spent catalyst and reaction product oil and gas, and then regenerating the spent catalyst and returning it to perform the catalytic cracking reaction.

[0019] The inventors of the present invention have discovered that harsh reaction conditions and high gas-solid contact efficiency are beneficial to improving the conversion rate and selectivity of producing light olefins and aromatics through catalytic cracking reactions. However, the existing riser reactor itself has the defects of uneven gas-solid distribution, with radially dilute gas in the middle and concentrated gas on both sides, and axially dilute gas at the top and concentrated gas at the bottom, which affects the conversion rate and selectivity of the catalytic cracking reaction. The existing downer reactor has a low bed particle concentration and a weak backmixing degree, resulting in low gas-solid contact efficiency, which also affects the conversion rate and selectivity of the catalytic cracking reaction. The inventors of the present invention have further discovered that if the catalytic cracking reaction is carried out under a rotating fluidized state, a flow pattern distribution with semi-free vortex and forced vortex can be formed in the reaction zone, thereby effectively enhancing the gas-solid contact efficiency, thereby allowing the use of more harsh reaction conditions (higher temperature and larger catalyst-to-oil ratio) and extending the gas-solid residence time, thereby improving the reaction conversion rate and selectivity, and thus increasing the yield of olefin and aromatic products.

[0020] Among them, a variety of swirl fluidization devices can be used to rotate the fluidized medium. According to a preferred embodiment of the present invention, the fluidized medium is rotated by a swirl distribution plate, which then drives the catalyst and the feedstock oil to reach a swirl fluidized state. Specifically, the fluidized medium flows into the tubular reactor from the inlet of the tubular reactor through the swirl distribution plate, and the catalytic cracking catalyst from the regenerator enters the tubular reactor from the upper part of the swirl distribution plate. The fluidized medium can be mixed with the catalyst flowing out of the catalyst delivery pipe after passing through the swirl distribution plate, and the catalyst also forms a swirl fluidized state under the action of the swirl fluidization field formed by the fluidized medium. After preheating, the feedstock oil enters the inlet of the tubular reactor through a nozzle and mixes with the catalyst. The feedstock oil and the catalyst are carried by the fluidized medium in the swirl fluidized state and also reach a swirl fluidized state. A catalytic cracking reaction is carried out in the tubular reactor, and then moves to the outlet of the tubular fluidized bed reactor and enters the gas-solid separator for gas-solid separation.

[0021] The swirl distribution plate can have various structures suitable for use in the method of the present invention. As a preferred embodiment, the swirl distribution plate includes a concentrically arranged outer ring and an inner ring, the diameter of the outer ring is 0.01-0.9 times, preferably 0.1-0.8 times, the inner diameter of the tubular fluidized bed reactor; the diameter of the inner ring is 0.01-0.5 times, preferably 0.1-0.4 times, the diameter of the outer ring; a fixed connection is provided between the outer ring and the inner wall of the tubular fluidized bed reactor, the inner ring forms a sealed connection with the catalyst input pipe, and a plurality of propeller-shaped distribution blades are radially spaced between the outer ring and the inner ring, the plurality of distribution blades have the same inclination direction and an inclination angle of 5°-85°, preferably 20°-70°; the gap between two adjacent distribution blades is used to allow the fluidized medium to pass obliquely and flow along the inner wall of the reactor to generate rotation. In this preferred embodiment, the swirl distribution plate has the advantages of large operating load and strong stability, can be applied to different operating gas velocities, and can achieve uniform distribution of swirl fields in different fluidization states.

[0022] Typically, the linear velocity of the swirling fluidized gas can be 0.1-10 m / s, preferably 1-5 m / s, and the angular velocity of the swirling fluidized gas can be 0.01-10 s -1 , preferably 0.1-5s -1 The angular velocity of the rotating fluidized gas can be adjusted by the tilt direction of the distribution blades, the linear velocity of the gas, the inner diameter of the tubular fluidized bed reactor, etc.

[0023] Optionally, the tubular fluidized bed reactor is a riser reactor, a horizontal tube reactor or a downtube reactor, preferably a riser reactor; the length-to-diameter ratio of the tubular fluidized bed reactor is 1-20:1, preferably 2-10:1.

[0024] Optionally, the catalytic cracking conditions include: reaction temperature of 500-800°C, preferably 550-650°C; reaction pressure of 0.1-2.0 MPa, preferably 0.3-1 MPa; catalyst-oil weight ratio of 5-100, preferably 12-30; residence time of 1-15 seconds, preferably 4-10 seconds.

[0025] Optionally, the fluidizing medium is at least one of water vapor and dry gas.

[0026] Optionally, the catalyst contains 10-60wt% of a molecular sieve, 1-40wt% of a binder and 1-90wt% of a carrier; the molecular sieve is a modified Y-type molecular sieve and / or an unmodified Y-type molecular sieve; the binder is a silica binder and / or an alumina binder; and the carrier can be selected from one or more of silica, kaolin, montmorillonite, diatomaceous earth, halloysite, saponite, rectorite, sepiolite, attapulgite, hydrotalcite and bentonite.

[0027] The method of the present invention is highly adaptable to feedstocks and is applicable not only to light catalytic cracking feedstocks such as hydrogenated LCO, but also to deep catalytic cracking of low-quality heavy feedstocks with high density and low hydrogen content. Optionally, the feedstock oil is at least one of unhydrogenated LCO, hydrogenated LCO, straight-run diesel, coker diesel, and hydrorefined straight-run diesel.

[0028] Optionally, in the modified Y-type molecular sieve, calculated as oxide, the rare earth element content is 3-15wt%, the sodium element content is 0.01-0.8wt%, and the zinc element content is 0.2-4.5wt%; the rare earth element may be but is not limited to one or more of La, Ce, Pr and Nd.

[0029] On the other hand, reference Figure 1 The present invention also provides a system for producing light olefins and aromatics by catalytic cracking. The system comprises a tubular fluidized bed reactor, a gas-solid separation device, a stripper, and a regenerator connected in sequence, wherein the regenerator outlet is connected to the catalyst inlet of the fluidized bed reactor; wherein the fluidized bed reactor is provided with a cyclonic fluidization device. This system is particularly suitable for use in the method of the present invention for producing light olefins and aromatics by catalytic cracking.

[0030] Preferably, the swirl fluidization device is a swirl distribution plate.

[0031] Wherein, the swirl distribution plate can have various structures suitable for the method of the present invention. As a preferred embodiment, refer to Figure 2 and Figure 3The swirl distribution plate includes an outer ring and an inner ring arranged concentrically, the diameter of the outer ring is 0.01-0.9 times the diameter of the fluidized bed reactor, preferably 0.1-0.8 times; the diameter of the inner ring is 0.01-0.5 times the diameter of the outer ring, preferably 0.1-0.4 times the diameter of the outer ring, a fixed connection is provided between the outer ring and the inner wall of the fluidized bed reactor, the inner ring forms a sealed connection with the catalyst input pipe, a plurality of propeller-shaped distribution blades are radially spaced between the outer ring and the inner ring, the inclination direction of the plurality of distribution blades is the same, and the inclination angle is 5°-85°, preferably 20°-70°; the gap between two adjacent distribution blades is used to allow the fluidized medium to pass obliquely and flow along the inner wall of the reactor to generate rotation.

[0032] Optionally, the fluidized bed reactor is a riser reactor, a horizontal tube reactor or a downtube reactor, preferably a riser reactor.

[0033] Optionally, the length-to-diameter ratio of the tubular fluidized bed reactor is 1-20:1, preferably 2-10:1.

[0034] According to a particularly preferred embodiment of the present invention, Figure 1-3 The method for producing light olefins and aromatics by catalytic cracking of the present invention comprises: introducing a fluidizing medium 1 into the bottom of a riser reactor 5, generating rotation through a swirl distribution plate 2, and allowing a catalyst to enter the riser reactor 5 through a slide valve 3 to be regenerated. The crude oil enters the riser reactor 5 through an oil inlet nozzle 4. The rotating fluidizing medium 1 drives the catalyst and the crude oil to reach a rotating fluidized state to carry out a catalytic cracking reaction. Gas-solid separation is then carried out through a cyclone separator 6, and stripping is then carried out through a stripper 7 into which stripping steam 9 is introduced to obtain a regenerated catalyst and reaction product oil and gas 8. The regenerated catalyst enters a regenerator 11 through a regeneration slide valve, is regenerated under the condition of introducing a regeneration medium 14, is separated through a cyclone separator 12 of the regenerator to obtain a regenerated catalyst, and the regenerated flue gas is discharged through a flue gas outlet pipeline 13.

[0035] The present invention is further described in detail below by way of examples. Unless otherwise specified, the raw materials used in the examples can be obtained from commercial sources.

[0036] The properties of the raw material hydrogenated LCO used in the examples and comparative examples are shown in Table 1.

[0037] Table 1

[0038] Raw oil name Hydrogenated LCO <![CDATA[密度(20℃),千克 / 米 3 ]]> 888.7 Carbon content, weight % 88.37 Hydrogen content, weight % 11.63 Hydrocarbon mass composition Paraffins, weight % 13.0 Total cycloalkanes 34.4 Total aromatics 52.6 Gum, weight % 0 Total weight, weight % 100

[0039] The catalyst SLA used in the examples and comparative examples is the same, and the catalyst properties are listed in Table 2. The preparation method is as follows: (1) NaY molecular sieve is ion exchanged with a rare earth salt solution (cerium nitrate) at a temperature of 40°C for 100 minutes. The mass ratio of the NaY molecular sieve, rare earth salt, and solvent water is 1:0.1:10; the masses of the NaY molecular sieve and rare earth salt are calculated on a dry basis and rare earth oxide, respectively. (2) The molecular sieve after ion exchange is calcined at a temperature of 400°C in an atmosphere with a water vapor content of 40% by volume for 6 hours. (3) The calcined molecular sieve is reacted with silicon tetrachloride at a temperature of 500°C for 3 hours. The mass ratio of the silicon tetrachloride to the calcined molecular sieve is 0.5:1. The mass of the calcined molecular sieve is calculated on a dry basis. (4) The molecular sieve was impregnated with a zinc salt solution (zinc nitrate), and the impregnated molecular sieve was calcined at an impregnation temperature of 40° C., a calcination temperature of 500° C., and a calcination time of 3 hours to obtain a modified Y molecular sieve having a rare earth element (Ce) content of 9 wt%, a sodium content of 0.5 wt%, and a zinc content of 2 wt%. (5) 20 wt% of the modified Y molecular sieve, 10 wt% of a binder (silicon oxide binder), 70 wt% of a carrier (kaolin), and water were formed into a slurry (solid content of 40 wt%), and the slurry was spray-dried to obtain the catalyst.

[0040] Table 2

[0041] Catalyst No. SLA Micro-anti-activity 80 Specific surface area, m2 / g 2 / g 124 Pore ​​volume, ml / g 0.26 Sieve composition, weight % 0~40 microns 11.8 40-80 microns 53.2 >80 microns 35

[0042] Example 1

[0043] This embodiment follows Figure 1 The device and process were tested in a small riser reactor. The hydrogenated LCO in Table 1 was used as the feed oil. The test was conducted on the riser reactor using SLA catalyst with a catalyst activity of 80. The catalyst properties are listed in Table 2. The reaction and regeneration process conditions were: hydrogenated LCO preheating temperature of 200°C, reaction outlet temperature of 520°C, reaction pressure of 0.2 MPa, catalyst-oil ratio of 10, residence time of 5.5 seconds, swirl fluidization gas linear velocity of 4 m / s, gas angular velocity of 2s -1 , the regenerator outlet temperature is 700℃, the regenerator pressure is 0.6MPa, the regenerator medium is air, and the weight ratio of water vapor to total raw materials is 0.15.

[0044] The catalyst enters the riser reactor through the spent catalyst slide valve, is fluidized by the fluidizing medium through the cyclone distribution plate, and then enters the bottom of the riser reactor. The fluidizing medium generates a rotating gas flow, which drives the catalyst to rotate and fluidize. The oil feed is preheated, enters the bottom of the riser reactor through the oil injection nozzle, and is mixed and contacted with the catalyst. The oil feed gas and the catalyst rotate and flow upward into the riser reactor for catalytic cracking. The reaction product gas and the catalyst are separated by the cyclone separator at the top of the riser reactor. The separated product gas enters the product gas discharge pipeline and enters the fractionation unit. The separated catalyst enters the stripper and is stripped by the stripping gas. The stripped catalyst enters the regenerator through the regenerated catalyst slide valve, is regenerated by burning coke, is separated by the cyclone separator, and is settled. The flue gas generated by the regeneration is separated by the cyclone separator in the settler and is discharged from the flue gas outlet pipeline. The regenerated catalyst enters the reactor through the spent catalyst slide valve for the next cycle. The operating conditions and product distribution are shown in Table 3.

[0045] Comparative Example 1

[0046] The comparative example uses the existing conventional riser catalytic cracking process, and the reaction is carried out in a small riser reactor. The non-hydrogen operation mode is used, and the catalyst and the oil feed are the same as in Example 1. The process conditions for the reaction and regeneration are as follows: the preheating temperature of the hydrogenated LCO is 200°C, the reaction outlet temperature is 520°C, the reaction pressure is 0.2 MPa, the catalyst to oil ratio is 10, the residence time is 3.5 seconds, the fluidizing gas linear velocity is 4 m / s, the regenerator outlet temperature is 700°C, the regenerator pressure is 0.6 MPa, the regenerator medium is air, and the weight ratio of water vapor to total feed is 0.15. The operating conditions and product distribution are shown in Table 3.

[0047] Example 2

[0048] This example is tested according to the device and process of Figure 1 The reaction is carried out in a small riser reactor, and the hydrogenated LCO in Table 1 is used as the oil feed. The SLA catalyst is used in the riser reactor, and the catalyst activity is 75. The catalyst properties are shown in Table 2. The process conditions for the reaction and regeneration are as follows: the preheating temperature of the hydrogenated LCO is 200°C, the reaction outlet temperature is 580°C, the reaction pressure is 0.4 MPa, the catalyst to oil ratio is 15, the residence time is 4.5 seconds, the rotating fluidizing gas linear velocity is 4 m / s, the gas angular velocity is 2 s -1 , the regenerator outlet temperature is 700°C, the regenerator pressure is 0.6 MPa, the regenerator medium is air, and the weight ratio of water vapor to total feed is 0.15.

[0049] The catalyst enters the riser reactor through the spent catalyst slide valve, is fluidized by the fluidization medium through the cyclone distribution plate, and then enters the bottom of the riser reactor. After the fluidization medium generates a rotating gas flow, the catalyst is driven to rotate and fluidize. After the raw oil is preheated, it enters the bottom of the riser reactor through the oil injection nozzle and is mixed and contacted with the catalyst. The raw oil gas and the catalyst rotate and flow upward into the riser reaction section to perform catalytic cracking. The reaction product oil gas and the catalyst are separated by the cyclone separator at the top of the riser, the separated product oil gas enters the product oil gas discharge pipeline through the product oil gas discharge pipeline, and the catalyst separated by the cyclone separator enters the stripper and is stripped by the stripping gas. The stripped catalyst enters the regenerator through the regeneration slide valve, is coked and regenerated, is separated by the cyclone separator and settles, and the flue gas generated by the regeneration is separated by the cyclone separator in the settler and is discharged from the flue gas outlet pipeline. The regenerated catalyst enters the reactor through the spent catalyst slide valve for the next cycle. The operating conditions and product distribution are shown in Table 3.

[0050] Comparative Example 2

[0051] This comparative example uses the existing conventional riser catalytic cracking process, and the reaction is carried out in a small riser reactor. The non-hydrogen operation mode is used, and the catalyst and raw oil are the same as in Example 1. The process conditions for reaction and regeneration are as follows: the preheating temperature of hydrogenated LCO is 200°C, the reaction outlet temperature is 580°C, the reaction pressure is 0.4 MPa, the catalyst to oil ratio is 15, the residence time is 3.0 seconds, the fluidization gas linear velocity is 4 m / s, the regenerator outlet temperature is 700°C, the regenerator pressure is 0.6 MPa, the regenerator pressure regeneration medium is air, and the weight ratio of water vapor to total raw material is 0.15. The operating conditions and product distribution are shown in Table 3.

[0052] As can be seen from Table 3, the triene yield and BTX yield of Examples 1 and 2 are higher than those of Comparative Examples 1 and 2. The above results show that the cyclone riser reactor with high temperature and large catalyst to oil ratio has higher low carbon olefin and aromatic hydrocarbon yield. By strengthening the gas-solid contact efficiency, the advantages of high severity are fully utilized, and the selectivity of the products is increased.

[0053] Table 3

[0054] Comparative Example 1 Comparative Example 2 Example 1 Example 2 crude oil Hydrogenated LCO Hydrogenated LCO Hydrogenated LCO Hydrogenated LCO Reactor type Conventional riser Conventional riser Cyclone riser Cyclone riser Catalyst name SLA SLA SLA SLA Catalyst activity (MAT) 80 80 80 80 Reaction operating conditions Reaction pressure, MPa 0.2 0.4 0.2 0.4 Reaction zone outlet temperature, ℃ 520 580 520 580 Catalyst / feedstock oil weight ratio 10 15 10 15 Oil and gas residence time, s 3.5 3.0 5.5 4.5 Product yield, weight % dry gas 1.27 3.14 1.75 4.33 Liquefied gas 6.65 11.51 8.29 14.56 gasoline 51.79 49.23 50.04 47.09 diesel fuel 37.67 31.26 36.57 27.85 Slurry 1.34 1.72 1.68 1.79 coke 1.28 3.14 1.67 4.37 total 100 100 100.00 100.00 Ethylene + propylene + butene 3.02 5.27 3.92 7.31 BTX 9.57 13.42 11.95 17.45

[0055] The preferred embodiments of the application are described in detail above, but the application is not limited to the specific details in the above-described embodiments. Within the technical concept of the application, various simple modifications can be made to the technical solutions of the application, and these simple modifications all belong to the protection scope of the application.

[0056] It should be further noted that any technically feasible combination of the various technical features described in the above embodiments is possible, provided that the combination does not contradict itself. In order to avoid unnecessary repetition, the present application does not describe each and every possible combination of the various technical features.

[0057] Furthermore, any combination of the various embodiments of the present application is possible, provided that the combination does not contradict itself, and should be considered as being disclosed by the present application.

Claims

1. A method for producing light olefins and aromatics by catalytic cracking, characterized in that: The method comprises: in a tubular fluidized bed reactor, subjecting reaction materials containing a fluidizing medium, a catalyst and a feedstock oil to a catalytic cracking reaction under a rotating fluidized state, then performing gas-solid separation and steam stripping to obtain spent catalyst and reaction product oil and gas, and then regenerating the spent catalyst and returning it to the catalytic cracking reaction; The fluidized medium is caused to rotate through a swirl distribution plate, thereby driving the catalyst and the feedstock oil to reach a swirling fluidized state; the swirl distribution plate includes a concentrically arranged outer ring and an inner ring, the diameter of the outer ring is 0.01-0.9 times the inner diameter of the tubular fluidized bed reactor, and the diameter of the inner ring is 0.01-0.5 times the diameter of the outer ring, a fixed connection is provided between the outer ring and the inner wall of the tubular fluidized bed reactor, and the inner ring forms a sealed connection with the catalyst input pipe, a plurality of propeller-shaped distribution blades are radially spaced between the outer ring and the inner ring, and the plurality of distribution blades have the same inclination direction and an inclination angle of 5°-85°; the gap between two adjacent distribution blades is used to allow the fluidized medium to pass obliquely and flow along the inner wall of the tubular fluidized bed reactor to generate rotation.

2. The method according to claim 1, wherein The inclination angle of the plurality of distribution blades is 20°-70°.

3. The method according to claim 1 or 2, wherein: The linear velocity of the swirling fluidized gas is 0.1-10 m / s; the angular velocity of the swirling fluidized gas is 0.01-10 s -1 .

4. The method according to claim 3, wherein: The linear velocity of the swirling fluidized gas is 1-5 m / s; the angular velocity of the swirling fluidized gas is 0.1-5 s -1 .

5. The method according to claim 1 or 2, wherein: The tubular fluidized bed reactor is a riser reactor, a horizontal tube reactor or a descending tube reactor; The tubular fluidized bed reactor has an aspect ratio of 1-20:1; The catalytic cracking conditions include: reaction temperature of 500-800° C.; reaction pressure of 0.1-2.0 MPa; catalyst-oil weight ratio of 5-100; and residence time of 1-15 seconds.

6. The method according to claim 5, wherein: The tubular fluidized bed reactor is a riser reactor; the length-to-diameter ratio of the tubular fluidized bed reactor is 2-10:1; The catalytic cracking conditions include: reaction temperature of 550-650° C.; reaction pressure of 0.3-1 MPa; catalyst-to-oil weight ratio of 12-30; and residence time of 4-10 seconds.

7. The method according to claim 1 or 2, wherein: The fluidizing medium is steam and / or dry gas; The catalyst contains 10-60 wt% of a molecular sieve, 1-40 wt% of a binder, and 1-90 wt% of a carrier; the molecular sieve is a modified Y-type molecular sieve and / or an unmodified Y-type molecular sieve; the binder is selected from a silica binder and / or an alumina binder; and the carrier is selected from at least one of silica, kaolin, montmorillonite, diatomaceous earth, halloysite, saponite, rectorite, sepiolite, attapulgite, hydrotalcite, and bentonite; The raw oil is at least one of unhydrogenated LCO, hydrogenated LCO, straight-run diesel, coker diesel and hydrorefined diesel.

8. The method according to claim 7, wherein: In the modified Y-type molecular sieve, the rare earth element content, calculated as oxide, is 3-15wt%, the sodium element content is 0.01-0.8wt%, and the zinc element content is 0.2-4.5wt%; the rare earth element is one or more of La, Ce, Pr and Nd.

9. A system for the method for producing light olefins and aromatics by catalytic cracking according to claim 1, characterized in that: The system comprises a tubular fluidized bed reactor, a gas-solid separation device, a stripper and a regenerator connected in sequence, wherein the discharge port of the regenerator is connected to the catalyst inlet of the tubular fluidized bed reactor; wherein the tubular fluidized bed reactor is provided with a cyclonic fluidization device; In which, the swirl fluidization device is a swirl distribution plate; the swirl distribution plate includes an outer ring and an inner ring arranged concentrically, the diameter of the outer ring is 0.01-0.9 times the inner diameter of the tubular fluidized bed reactor, and the diameter of the inner ring is 0.01-0.5 times the diameter of the outer ring, a fixed connection is provided between the outer ring and the inner wall of the tubular fluidized bed reactor, the inner ring forms a sealed connection with the catalyst input pipe, and a plurality of propeller-shaped distribution blades are radially spaced between the outer ring and the inner ring, and the inclination direction of the plurality of distribution blades is the same, and the inclination angle is 5°-85°; the gap between two adjacent distribution blades is used to allow the fluidized medium to pass obliquely and flow along the inner wall of the tubular fluidized bed reactor to generate rotation.

10. The system according to claim 9, wherein: The inclination angle of the plurality of distribution blades is 20°-70°.

11. The system according to claim 9 or 10, wherein: The tubular fluidized bed reactor is a riser reactor, a horizontal tube reactor or a descending tube reactor.

12. The system according to claim 11, wherein The tubular fluidized bed reactor is a riser reactor.

13. The system according to claim 9 or 10, wherein: The tubular fluidized bed reactor has a length-to-diameter ratio of 1-20:

1.

14. The system according to claim 13, wherein: The length-to-diameter ratio of the tubular fluidized bed reactor is 2-10:1.

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