Method and system for producing light aromatic hydrocarbon

Abstract

Provided in the present invention are a method and system for producing a light aromatic hydrocarbon. A catalytic cracking cycle oil undergoes a two-stage hydrogenation treatment in the presence of hydrogen and a hydrogenation catalyst, and the hydrogenated oil gas passes through a first rectification tower device to obtain an overhead light fraction, a light aromatic hydrocarbon-rich fraction, a first recycle fraction having a specific distillation range, and a second recycle fraction having a specific distillation range; and the first recycle fraction enters a first reaction zone of a catalytic cracking device for a catalytic cracking reaction, and the second recycle fraction enters a second reaction zone of the catalytic cracking device for a reaction or returns to a second section of the hydrogenation treatment for recycling. Furthermore, a reaction product obtained from catalytic cracking is separated by a second rectification tower device to obtain a first fraction oil gas, a second fraction oil, and an oil slurry, the second fraction oil returns to a first section of the hydrogenation treatment, and optionally, the first fraction oil gas along with the hydrogenated oil gas enters the first rectification tower device for separation. In the present invention, a catalytic conversion product and a hydrogenation product are effectively separated in the same separation system and then undergo catalytic conversion in different zones, thereby saving the separation equipment and device investment while improving the yield of the produced light aromatic hydrocarbon.

Description

Method and system for producing light aromatics Technical Field

[0001] The present invention relates to the field of petrochemical industry, and in particular to a method and system for producing light aromatic hydrocarbons. Background Art

[0002] With the trend toward heavier crude oil and the rapidly growing market demand for lighter, higher-quality products, catalytic cracking (FCC) technology, a method for lightening heavy oil, has rapidly developed in my country. However, the quality of FC diesel (also known as FC light cycle oil) has been relatively poor, characterized by high density, high aromatics content, and a low cetane number. Even diesel hydro-reforming technology has struggled to meet increasingly stringent diesel specifications. Light aromatics (benzene, toluene, and xylene, collectively known as BTX), represented by paraxylene, are important raw materials for the production of chemical products. Market demand for these hydrocarbons is increasing, and demand currently exceeds supply. Therefore, producing light aromatics through FC diesel is a market-driven approach.

[0003] US4585545A discloses a catalytic conversion method for producing gasoline rich in monocyclic aromatics by first subjecting the whole fraction of catalytically cracked light cycle oil to hydrogenation treatment and then subjecting the obtained hydrogenated diesel to catalytic cracking.

[0004] CN01807978.4 (CN1422327A) discloses a method for upgrading catalytic cracking light cycle oil, which involves deep hydrogenation of catalytic light cycle oil produced by a first catalytic cracking unit using heavy oil as raw material, and the resulting hydrogenated diesel is fed to a second catalytic cracking unit.

[0005] CN01808113.4 (CN1466619A) discloses a method for converting catalytic cracking light cycle oil, which divides the catalytic cracking riser reaction zone into two reaction zones, upstream and downstream, wherein heavy oil is injected into the downstream zone, and the hydrogenated cycle oil obtained by hydrogenating the catalytic cracking product light cycle oil is injected into the upstream zone.

[0006] CN103923698B discloses a method for producing aromatic compounds. The diesel fraction with a distillation range of 250 to 450°C obtained by catalytic cracking is sent to a hydrotreating unit, and the catalytic cracking heavy cycle oil is hydrogenated and then catalytically cracked to produce benzene, toluene, and xylene.

[0007] CN104560187B discloses a catalytic conversion method for producing aromatics-rich gasoline, which cuts catalytic cracking light cycle oil into a light fraction and a heavy fraction. The heavy fraction is hydrogenated and then separately fed into different catalytic cracking units together with the light fraction to maximize the production of catalytic gasoline rich in benzene, toluene, and xylene.

[0008] From the above-disclosed literature, it can be found that one of the important ways to process catalytic cracking light cycle oil is to hydrotreat it before catalytic cracking. However, it is worth noting that the selection of fractions before and after hydrogenation is crucial to both hydrogen consumption and the yield of light aromatics from catalytic cracking. Summary of the Invention

[0009] The object of the present invention is to provide a method and system for producing light aromatics based on the prior art.

[0010] To achieve the above objectives, the present invention provides a method and system for producing light aromatics, wherein catalytic cracking cycle oil is subjected to two-stage hydrogenation treatment in the presence of hydrogen and a hydrogenation catalyst. The hydrogenated oil and gas pass through a first distillation tower device to obtain a light fraction at the top of the tower, a fraction rich in light aromatics, a first recycled fraction with a specific distillation range, and a second recycled fraction with a specific distillation range; the first recycled fraction enters the first reaction zone of the catalytic cracking unit for catalytic cracking reaction, and the second recycled fraction enters the second reaction zone of the catalytic cracking unit for reaction or returns to the second stage of hydrogenation treatment for recycling. Furthermore, the reaction products obtained by catalytic cracking are separated by a second distillation tower device to obtain a first distillate oil and gas, a second distillate oil, and an oil slurry. The second distillate oil is returned to the first stage of the above-mentioned hydrogenation treatment. Optionally, the first distillate oil and gas and the hydrogenated oil and gas can enter the above-mentioned first distillation tower device together for separation.

[0011] A first aspect of the present invention provides a method for producing light aromatics, comprising the steps of:

[0012] (1) Catalytic cracking cycle oil is subjected to two-stage hydrogenation treatment in the presence of hydrogen and a hydrogenation catalyst to obtain a hydrogenated product;

[0013] (2) hydrogen is separated from the hydrogenated product to obtain hydrogenated oil gas, which is passed through a first distillation tower to obtain a light fraction at the top of the tower, a fraction rich in light aromatics, a first recycled fraction, and a second recycled fraction. The first recycled fraction has an initial distillation point of 140 to 170° C. and a final distillation point of 220 to 300° C., and the second recycled fraction has a distillation range starting point temperature of any value between 220 and 300° C.;

[0014] (3) The first recycled fraction obtained in step (2) enters the first reaction zone of the catalytic cracking unit for catalytic cracking reaction. Optionally, the second recycled fraction enters the second reaction zone of the catalytic cracking unit or returns to the second stage of hydrogenation for recycling;

[0015] (4) The reaction product obtained from the catalytic cracking unit is separated into a first fraction oil gas, a second fraction oil and an oil slurry by a second distillation tower device, and the second fraction oil is returned to the first hydrogenation treatment stage. Optionally, the first fraction oil gas is separated from the hydrogenated oil gas by entering the first distillation tower device of step (2);

[0016] (5) Separating light aromatics from the light aromatics-rich fraction.

[0017] According to the method of the first aspect, wherein, in step (1), the hydrotreatment reaction conditions satisfy at least one of the following conditions:

[0018] The reaction conditions of the first stage of hydrogenation treatment are: hydrogen partial pressure 5.0-20.0 MPa, reaction temperature 300-450°C, volume space velocity 0.5-5.0 h -1 , Hydrogen to oil volume ratio 300~1600Nm 3 / m 3 and / or

[0019] The reaction conditions of the second stage hydrotreatment are as follows: hydrogen partial pressure 5.0-20.0 MPa, reaction temperature 10-50°C higher than the first stage hydrotreatment temperature, volume space velocity 0.1-3.0 h -1 , Hydrogen to oil volume ratio 400~2000Nm 3 / m 3 .

[0020] According to the method of the first aspect, wherein, in step (1), in the two-stage hydroprocessing reaction, the hydrogenation catalysts are the same or different;

[0021] The hydrogenation catalyst of the first stage of the hydroprocessing comprises an active metal component and a support, and the hydrogenation catalyst of the second stage of the hydroprocessing comprises an active metal, a molecular sieve, and a support, wherein the active metal component is selected from one or more of Group VIB metals and / or Group VIII non-noble metals, the support is selected from one or more of alumina, silica, and amorphous silica-alumina, and the molecular sieve is one or more selected from Y-type molecular sieve and USY molecular sieve;

[0022] More preferably, the active metal component is selected from any one of the following metal combinations: nickel-tungsten, nickel-tungsten-cobalt, nickel-molybdenum, and cobalt-molybdenum.

[0023] According to the method of the first aspect, in step (1), the initial boiling point of the catalytic cracking cycle oil is 160-250°C, and the final boiling point is 340-400°C; the initial boiling point of the hydrogenated oil gas is 0-60°C, and the final boiling point is 280-360°C.

[0024] According to the method of the first aspect, wherein, in step (2), the first distillation tower device is a next-wall distillation tower, and at least one dividing wall is provided in the next-wall distillation tower, and the dividing wall divides the next-wall distillation tower into four areas: a pre-separation area, a common distillation section, a common stripping section and a side-line distillation area, and the feed position of the hydrogenated oil and gas is located in the pre-separation area.

[0025] According to the method of the first aspect, wherein, in step (2), the number of trays in the pre-separation zone in the dividing wall distillation tower is 10 to 40, preferably 20 to 35;

[0026] The number of plates in the common distillation section is 2 to 10, preferably 3 to 8;

[0027] The number of trays in the common stripping section is 2 to 10, preferably 3 to 8;

[0028] The number of trays in the side distillation zone is 10 to 45, preferably 25 to 40;

[0029] And the dividing wall distillation column satisfies at least one of the following conditions:

[0030] The feed position of the hydrogenated oil and gas is above the 4th to 22nd plates at the top of the pre-separation zone, preferably above the 8th to 20th plates;

[0031] The top pressure of the dividing wall distillation tower is 0.1-1.0 MPa, preferably 0.2-0.7 MPa, and the top temperature is 50-120° C., preferably 60-100° C.; and / or

[0032] The top reflux ratio is 3 to 10, preferably 4 to 8.

[0033] According to the method of the first aspect, wherein, in step (2), the dividing wall distillation tower satisfies at least one of the following conditions:

[0034] The side distillation zone of the distillation tower is provided with a middle section circulation heat extraction, which is the 20th to 45th tray of the side distillation zone, more preferably the 30th to 40th tray;

[0035] Preferably, the pre-separation zone of the dividing wall distillation tower is provided with a gas phase withdrawal, the withdrawal position is located at the 20th to 40th plate of the pre-separation zone, more preferably at the 25th to 35th plate, the gas phase return position is the 30th to 45th plate of the side line distillation zone, more preferably at the 35th to 40th plate, and the return position is the same as the withdrawal position or is one plate above or below the withdrawal position;

[0036] The mass flow ratio of the liquid phase from the common distillation section into the pre-separation zone and the side distillation zone is 0.1 to 10, preferably 0.5 to 5;

[0037] The mass flow ratio of the gas phase from the common stripping section into the pre-separation zone and the side-line distillation zone is 1 to 15, preferably 2 to 10; and / or

[0038] The ratio of the total gas phase mass flow rate at the extraction position of the pre-separation zone to the extracted gas phase mass flow rate is 10-150, preferably 30-120.

[0039] According to the method of the first aspect, wherein, in step (2), the dividing wall distillation tower satisfies at least one of the following conditions:

[0040] The light fraction at the top of the tower is separated from the common distillation section of the distillation tower, with an initial distillation point of 20-40°C and a final distillation point of 55-65°C;

[0041] The light aromatic hydrocarbon-rich fraction is extracted from the 5th to 30th plates of the side distillation zone of the dividing wall distillation tower, preferably the 10th to 20th plates, with an initial distillation point of 55 to 65° C. and a final distillation point of 140 to 170° C.

[0042] The first recycled fraction is produced from any position below the production position of the light aromatics-rich fraction in the side distillation zone, preferably the 25th to 35th tray; and / or

[0043] The second recycled fraction is separated from the lower part of the common stripping section of the dividing wall distillation tower.

[0044] According to the method of the first aspect, wherein, in step (3), the conditions for catalytic conversion in the first reaction zone meet at least one of the following conditions:

[0045] The weight hourly space velocity of oil and gas is 4 to 20h -1 , preferably 6 to 12 hours -1 ;

[0046] The reaction time is 0.5 to 8 s, preferably 1 to 5 s; and / or

[0047] The weight ratio of agent to oil is 6 to 26, preferably 8 to 12.

[0048] According to the method of the first aspect, wherein, in step (3), the conditions for catalytic conversion in the second reaction zone meet at least one of the following conditions:

[0049] The weight hourly space velocity of oil and gas is 4 to 20h -1 , preferably 8 to 12 hours -1 ;

[0050] The reaction time is 1 to 10 seconds, preferably 2 to 6 seconds; and / or

[0051] The weight ratio of agent to oil is 4 to 14, preferably 6 to 10.

[0052] According to the method of the first aspect, wherein, in step (4), the top pressure of the distillation tower is 0.15-0.4 MPa, preferably 0.18-0.35 MPa, and the top temperature is 100-160°C, preferably 115-145°C.

[0053] According to the method of the first aspect, the second distillation tower device in step (4) satisfies at least one of the following conditions:

[0054] The first fraction oil and gas has an initial boiling point of 0-40°C and a final boiling point of 220-250°C, and the first fraction oil and gas is degassed after being combined with the hydrogenated oil and gas and then enters the first distillation tower device of step (2) for separation;

[0055] The initial boiling point of the second fraction oil is 220-250°C, the final boiling point is 340-360°C, and its initial boiling point is not lower than the initial boiling point of the first fraction oil gas; and / or

[0056] The initial boiling point of the oil slurry is 340-360°C.

[0057] A second aspect of the present invention provides a system for producing light aromatics, comprising:

[0058] A hydrogenation unit, wherein the hydrogenation unit is a two-stage hydrogenation unit, provided with a hydrogen inlet and a hydrogenation product outlet, a catalytic cracking cycle oil inlet provided on the first stage of the hydrogenation unit, and an optional second recycled fraction inlet provided on the second stage of the hydrogenation unit;

[0059] A catalytic cracking unit comprising a catalytic cracking reactor, wherein the catalytic cracking reactor is provided with a first reaction zone and an optional second reaction zone, wherein the first reaction zone is provided with a first recycled fraction inlet, and the optional second reaction zone is provided with a second recycled fraction inlet;

[0060] a second distillation tower device, wherein the product distillation tower device is provided with a catalytic cracking reaction product inlet, a first distillate oil and gas outlet, a second distillate oil outlet and an oil slurry outlet, wherein the second distillate oil outlet is connected to the catalytic cracking cycle oil inlet of the hydrogenation unit;

[0061] a degassing and separation device, wherein the degassing and separation device is provided with an oil and gas inlet, a rich gas outlet, and a distillate oil outlet, the oil and gas inlet of the degassing and separation device being connected to the first distillate oil and gas outlet of the second rectifying tower device and the hydrogenated product outlet of the hydrogenation device, so that the first distillate oil and gas from the second rectifying tower device and the hydrogenated oil and gas from the hydrogenation device are degassed and separated into rich gas and distillate oil;

[0062] a first distillation tower device, the distillation tower device being provided with a distillate inlet, an outlet for a light fraction at the top of the tower, an outlet for a fraction rich in light aromatics, a first recycled fraction outlet, and a second recycled fraction outlet, the distillate inlet being connected to the distillate outlet of the degassing separation device, the first recycled fraction outlet being connected to the first recycled fraction inlet of the catalytic cracking device, the second recycled fraction outlet being connected to a second recycled fraction inlet provided in an optional second reaction zone of the catalytic cracking device, or the second recycled fraction outlet being connected to a second recycled fraction inlet optionally provided in the second section of the hydrogenation device;

[0063] The separation device is provided with an inlet for a light aromatic hydrocarbon-rich fraction and an outlet for a light aromatic hydrocarbon-rich fraction, and the inlet for the light aromatic hydrocarbon-rich fraction is connected to the outlet for the light aromatic hydrocarbon-rich fraction of the first distillation tower device.

[0064] In the second aspect of the present invention, the first distillation tower device is a dividing wall distillation tower, and the dividing wall distillation tower is provided with at least one dividing wall, and the dividing wall divides the dividing wall distillation tower into four areas: a common stripping section provided at the bottom of the tower, a common distillation section provided at the top of the tower, a pre-separation zone provided in the middle of the tower, and a side distillation zone; the pre-separation zone is provided with a distillate oil inlet, and the distillate oil inlet is connected to the distillate oil outlet of the degassing separation device; the common distillation section is provided with a light fraction outlet at the top of the tower, the side distillation zone is provided with a light aromatics-rich fraction outlet and a first recycle fraction outlet, and the bottom of the common stripping section is provided with a second recycle fraction outlet.

[0065] At this time, the first recycle fraction outlet of the side distillation zone of the distillation tower is connected to the first recycle fraction inlet of the catalytic cracking reactor.

[0066] The second recycle fraction outlet at the bottom of the common stripping section of the dividing wall distillation tower is connected to the second recycle fraction inlet optionally provided on the second reaction zone of the first distillation tower device, or is connected to the second recycle fraction inlet optionally provided on the second section of the hydrogenation device.

[0067] In the second aspect of the present invention, the system further comprises a high-low separation device, the system further comprises a high-low separation device, the high-low separation device is provided with an inlet for hydrogenated products, a hydrogen outlet and a hydrogenated oil and gas outlet, and the high-low separation device separates the hydrogenated oil and gas from the hydrogenated products from the hydrogenation device;

[0068] At this time, the hydrogenation product inlet is connected to the hydrogenation product outlet of the hydrogenation device, and the oil and gas inlet of the degassing separation device is connected to the first fraction oil and gas outlet of the second distillation tower device and the hydrogenation oil and gas outlet of the high and low fraction device.

[0069] The system according to the second aspect, wherein the first reaction zone and the second reaction zone of the catalytic cracking reactor are arranged on the same or different reactors;

[0070] Preferably, the reactor is selected from one or more of the following: a constant diameter riser, a constant linear speed riser, a variable diameter riser, a variable linear speed riser, a fluidized bed, and a composite reactor. The composite reactor is composed of a constant diameter or variable diameter riser and a fluidized bed. Preferably, the reactor is a riser reactor.

[0071] More preferably, the first reaction zone and the second reaction zone of the catalytic cracking reactor are arranged on a double-riser reactor, wherein the main riser is the first reaction zone and the secondary riser is the second reaction zone.

[0072] Through the above technical solution, the effective components of each specific distillation range after hydrogenation and catalytic cracking are efficiently separated and then enter the catalytic cracking unit or hydrogenation unit for zoned conversion, greatly improving the yield of light aromatics.

[0073] Other features and advantages of the present invention will be described in detail in the following detailed description. BRIEF DESCRIPTION OF THE DRAWINGS

[0074] 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:

[0075] FIG1 is a schematic diagram of a system structure in a preferred embodiment of the present invention.

[0076] FIG2 is a schematic diagram of the system structure in another preferred embodiment of the present invention.

[0077] LIST OF NUMERALS AND SIGNS 100. Hydrogenation unit; 200. Dividing-wall distillation column; 300. Catalytic cracking unit; 400. Distillation column; 101. Catalytic cracking cycle oil inlet; 102. Hydrogenation product pipeline; 103. Hydrogenation oil and gas pipeline; 201. Distillate oil pipeline; 202. Overhead light fraction pipeline; 203. Light aromatics-rich fraction pipeline; 204. First recycle fraction pipeline; 205. Second recycle fraction pipeline; A. Pre-separation zone; B. Common distillation section; C. Side-stream distillation section; D. Common stripping section; E. Dividing wall; 301. First reaction zone of catalytic cracking unit; 302. Second reaction zone of catalytic cracking unit; 303. Feed oil inlet for first reaction zone; 304. Feed oil inlet for second reaction zone; 305. Catalytic product pipeline; 401. First fraction oil and gas pipeline; 402. Second fraction oil pipeline; 403. Slurry oil pipeline; DETAILED DESCRIPTION

[0078] The present application will be further described in detail below through the accompanying drawings and examples, through which the features and advantages of the present application will become more clear and distinct.

[0079] The word "exemplary" is used exclusively herein to mean "serving as an example, example, or illustration." Any embodiment described herein as "exemplary" is not necessarily to be construed as preferred or advantageous over other embodiments. Although various aspects of the embodiments are shown in the drawings, the drawings are not necessarily drawn to scale unless otherwise indicated.

[0080] Any specific numerical value disclosed in this application (including the endpoints of a numerical range) is not limited to the exact value of the numerical value, but should be understood to also include values ​​close to the exact value, such as all possible values ​​within the range of ±5% of the exact value. Moreover, for a disclosed numerical range, the values ​​between the endpoints of the range, between the endpoints and the specific points in the range, and between the specific points can be arbitrarily combined to form one or more new numerical ranges, and these new numerical ranges should also be considered to be specifically disclosed herein.

[0081] In this application, the terms "upstream" and "downstream" are used in relation to the direction of flow of the reactants. For example, when the reactants flow from bottom to top, "upstream" refers to a position below, while "downstream" refers to a position above.

[0082] Unless otherwise specified, the terms used herein have the same meaning as commonly understood by those skilled in the art. If a term is defined herein and its definition is different from the commonly understood meaning in the art, the definition herein shall prevail.

[0083] 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.

[0084] In addition, the technical features involved in different embodiments of the present application described below can be combined with each other as long as they do not conflict with each other.

[0085] In the context of the present invention, unless otherwise specified, the physical property values ​​of a substance (such as boiling point) are all measured values ​​at normal temperature / room temperature (25° C.) and normal pressure (101325 Pa).

[0086] Before describing the technical solution of the present invention, the terms used herein are defined as follows:

[0087] The term "initial distillation point" refers to the temperature at the starting point of the distillation range.

[0088] The term "end point" refers to the temperature at the end of the distillation range, as determined using ASTM D86, an internationally accepted standard method.

[0089] The light aromatic hydrocarbons mentioned in the present invention refer to C6 to C8 aromatic hydrocarbons, specifically benzene, toluene, xylene, and ethylbenzene.

[0090] The present invention provides a method for producing light aromatics, comprising the following steps:

[0091] (1) Catalytic cracking cycle oil is subjected to two-stage hydrogenation treatment in the presence of hydrogen and a hydrogenation catalyst to obtain a hydrogenated product;

[0092] (2) hydrogen is separated from the hydrogenated product to obtain hydrogenated oil gas, which is passed through a first distillation tower to obtain a light fraction at the top of the tower, a fraction rich in light aromatics, a first recycled fraction, and a second recycled fraction. The first recycled fraction has an initial distillation point of 140 to 170° C. and a final distillation point of 220 to 300° C., and the second recycled fraction has a distillation range starting point temperature of any value between 220 and 300° C.;

[0093] (3) The first recycled fraction obtained in step (2) enters the first reaction zone of the catalytic cracking unit for catalytic cracking reaction. Optionally, the second recycled fraction enters the second reaction zone of the catalytic cracking unit or returns to the second stage of hydrogenation for recycling;

[0094] (4) The reaction product obtained from the catalytic cracking unit is separated into a first fraction oil gas, a second fraction oil and an oil slurry by a second distillation tower device, and the second fraction oil is returned to the first hydrogenation treatment stage. Optionally, the first fraction oil gas is separated from the hydrogenated oil gas by entering the first distillation tower device of step (2);

[0095] (5) Separating light aromatics from the light aromatics-rich fraction.

[0096] The inventors have discovered that effectively separating the catalytic cracking products from the hydrogenation products and then returning the heavy aromatics fraction to the catalytic cracking riser for zoned conversion can increase the yield of light aromatics from catalytic cracking. This invention effectively separates the catalytic conversion products and hydrogenation products within the same separation system, followed by zoned catalytic conversion. This improves the yield of light aromatics while reducing investment in separation equipment.

[0097] In the present invention, the catalytic cracking cycle oil used as a raw material can be catalytic cracking cycle oil (light cycle oil, LCO) known in the art. Preferably, the catalytic cracking cycle oil of the present invention has an initial boiling point of 160-250°C and a final boiling point of 340-400°C.

[0098] In the present invention, the conditions and catalysts known in the art can be used for the hydrotreatment of catalytic cracking cycle oil.

[0099] In one embodiment, in step (1), the reaction conditions of the first stage of the hydrotreatment are: hydrogen partial pressure 5.0-20.0 MPa, reaction temperature 300-450°C, volume space velocity 0.5-5.0 h - 1 , Hydrogen to oil volume ratio 300~1600Nm 3 / m 3 .

[0100] In one embodiment, in step (1), the reaction conditions of the second stage of the hydrotreatment are: hydrogen partial pressure 5.0-20.0 MPa, reaction temperature 350-500°C, preferably 10-50°C higher than the temperature of the first stage of the hydrotreatment, volume space velocity 0.1-3.0 h -1 , Hydrogen to oil volume ratio 400~2000Nm 3 / m 3 .

[0101] According to the method of the first aspect, in step (1), in the two-stage hydroprocessing reaction, the hydrogenation catalysts are the same or different.

[0102] According to the method of the first aspect, wherein, in step (1), the hydrogenation catalyst of the first stage of the hydroprocessing comprises an active metal component and a support, and the hydrogenation catalyst of the second stage of the hydroprocessing comprises an active metal, a molecular sieve and a support, wherein the active metal component is selected from one or more of Group VIB metals and / or Group VIII non-noble metals, the support is selected from one or more of alumina, silica, and amorphous silica-alumina, and the molecular sieve is selected from one or more of Y-type molecular sieve and USY molecular sieve;

[0103] More preferably, the active metal component is selected from any one of the following metal combinations: nickel-tungsten, nickel-tungsten-cobalt, nickel-molybdenum, and cobalt-molybdenum.

[0104] In one embodiment, in step (1), the initial boiling point of the catalytic cracking cycle oil is 160-250°C, and the final boiling point is 340-400°C.

[0105] In one embodiment, in step (2), the initial boiling point of the hydrogenated oil gas is 0-60°C, and the final boiling point is 280-350°C.

[0106] In one embodiment, in step (2), the first distillation tower device is a bulkhead distillation tower, and at least one dividing wall is provided in the bulkhead distillation tower, and the dividing wall divides the bulkhead distillation tower into four areas: a pre-separation area, a common distillation section, a common stripping section and a side distillation area, and the feed position of the hydrogenated oil and gas is located in the pre-separation area.

[0107] In one embodiment, in step (2), the first distillation tower device is a dividing wall distillation tower, and the number of plates in the pre-separation zone in the dividing wall distillation tower is 10 to 40, preferably 20 to 35;

[0108] The number of plates in the common distillation section is 2 to 10, preferably 3 to 8;

[0109] The number of trays in the common stripping section is 2 to 10, preferably 3 to 8;

[0110] The number of trays in the side distillation zone is 10 to 45, preferably 25 to 40;

[0111] The feed position of the hydrogenated oil gas is above the 4th to 22nd plates at the top of the pre-separation zone, preferably above the 8th to 20th plates.

[0112] In one embodiment, in step (2), the top pressure of the dividing wall distillation tower is 0.1 to 1.0 MPa, preferably 0.2 to 0.7 MPa, and the top temperature is 50 to 120° C., preferably 60 to 100° C.; and / or

[0113] The top reflux ratio is 3 to 10, preferably 4 to 8.

[0114] In one embodiment, in step (2), the dividing wall distillation column satisfies at least one of the following conditions:

[0115] The side distillation zone of the distillation tower is provided with a mid-stage circulation heat extraction, which is the 20th to 45th tray of the side distillation zone, more preferably the 30th to 40th tray, and the return position is the same as the extraction position or is one tray above or below the extraction position;

[0116] The pre-separation zone of the dividing wall distillation tower is provided with a gas phase extraction, the extraction position is located at the 20th to 40th plate of the pre-separation zone, more preferably the 25th to 35th plate, and the gas phase return position is the 30th to 45th plate of the side line distillation zone, more preferably the 35th to 40th plate;

[0117] The mass flow ratio of the liquid phase from the common distillation section into the pre-separation zone and the side distillation zone is 0.1 to 10, preferably 0.5 to 5;

[0118] The mass flow ratio of the gas phase from the common stripping section into the pre-separation zone and the side-line distillation zone is 1 to 15, preferably 2 to 10; and / or

[0119] The ratio of the total gas phase mass flow rate at the extraction position of the pre-separation zone to the extracted gas phase mass flow rate is 10-150, preferably 30-120.

[0120] In one embodiment, in step (2), the dividing wall distillation column satisfies at least one of the following conditions:

[0121] The light fraction at the top of the tower is separated from the common distillation section of the distillation tower, with an initial distillation point of 20-40°C and a final distillation point of 55-65°C;

[0122] The light aromatic hydrocarbon-rich fraction is extracted from the 5th to 30th plates of the side distillation zone of the dividing wall distillation tower, preferably the 10th to 20th plates, with an initial distillation point of 55 to 65° C. and a final distillation point of 140 to 170° C.

[0123] The first recycled fraction is produced from any position below the production position of the light aromatics-rich fraction in the side distillation zone, preferably the 25th to 35th tray; and / or

[0124] The second recycled fraction is separated from the lower portion of the common stripping section of the dividing wall distillation tower. In one embodiment, in step (3), the conditions for catalytic conversion in the first reaction zone meet at least one of the following conditions:

[0125] The weight hourly space velocity of oil and gas is 4 to 20h -1 , preferably 6 to 12 hours -1 ;

[0126] The reaction time is 0.5 to 8 s, preferably 1 to 5 s; and / or

[0127] The weight ratio of agent to oil is 6 to 26, preferably 8 to 12.

[0128] In one embodiment, in step (3), the conditions for catalytic conversion in the second reaction zone meet at least one of the following conditions:

[0129] The weight hourly space velocity of oil and gas is 4 to 20h -1 , preferably 8 to 12 hours -1 ;

[0130] The reaction time is 1 to 10 seconds, preferably 2 to 6 seconds; and / or

[0131] The weight ratio of agent to oil is 4 to 14, preferably 6 to 10.

[0132] In the present invention, the catalytic cracking catalyst used in the catalytic cracking unit can be a Y molecular sieve catalytic cracking catalyst known in the art.

[0133] In one embodiment, the second distillation tower device in step (4) satisfies at least one of the following conditions:

[0134] The top pressure of the distillation tower device is 0.15-0.4 MPa, preferably 0.18-0.35 MPa, and the top temperature is 100-160°C, preferably 115-145°C.

[0135] The first fraction oil and gas has an initial boiling point of 0-40°C and a final boiling point of 220-250°C, and the first fraction oil and gas is degassed after being combined with the hydrogenated oil and gas and then enters the first distillation tower device of step (2) for separation;

[0136] The initial boiling point of the second fraction oil is 220-250°C, the final boiling point is 340-360°C, and its initial boiling point is not lower than the initial boiling point of the first fraction oil gas; and / or

[0137] The initial boiling point of the oil slurry is 340-360°C.

[0138] In one embodiment, in step (2), the initial distillation point of the first recycled fraction is 140-170°C, the final distillation point is 220-250°C, and the distillation range starting point temperature of the second recycled fraction is any value between 220-250°C.

[0139] In one embodiment, in step (2), the initial distillation point of the first recycled fraction is 140-170°C, the final distillation point is 250-300°C, and the distillation range starting point temperature of the second recycled fraction is any value between 250-300°C.

[0140] In one embodiment, when the distillation range starting point temperature of the second recycled fraction is any value between 220° C. and 250° C., in step (3), the second recycled fraction enters the second reaction zone of the catalytic cracking unit.

[0141] In one embodiment, when the distillation range starting point temperature of the second recycled fraction is any value between 270° C. and 300° C., in step (3), the second recycled fraction is returned to the second stage of the hydrotreatment for recycling.

[0142] In one embodiment, when the distillation range starting temperature of the second recycled fraction is any value between 250°C and 270°C, in step (3), the second recycled fraction enters the second reaction zone of the catalytic cracking unit or returns to the second stage of the hydroprocessing for recycling.

[0143] In one embodiment, in step (5), light aromatics are separated from the light aromatics-rich fraction. The separation method can be a method known in the art, including but not limited to fractionation, extraction, extraction, adsorption, membrane separation, etc.

[0144] The present invention also provides a system for producing light aromatics, comprising:

[0145] A hydrogenation unit, wherein the hydrogenation unit is a two-stage hydrogenation unit, provided with a hydrogen inlet and a hydrogenation product outlet, a catalytic cracking cycle oil inlet provided on the first stage of the hydrogenation unit, and an optional second recycled fraction inlet provided on the second stage of the hydrogenation unit;

[0146] A catalytic cracking unit comprising a catalytic cracking reactor, wherein the catalytic cracking reactor is provided with a first reaction zone and an optional second reaction zone, wherein the first reaction zone is provided with a first recycled fraction inlet, and the optional second reaction zone is provided with a second recycled fraction inlet;

[0147] a second distillation tower device, wherein the product distillation tower device is provided with a catalytic cracking reaction product inlet, a first distillate oil and gas outlet, a second distillate oil outlet and an oil slurry outlet, wherein the second distillate oil outlet is connected to the catalytic cracking cycle oil inlet of the hydrogenation unit;

[0148] a degassing and separation device, wherein the degassing and separation device is provided with an oil and gas inlet, a rich gas outlet, and a distillate oil outlet, the oil and gas inlet of the degassing and separation device being connected to the first distillate oil and gas outlet of the second rectifying tower device and the hydrogenated product outlet of the hydrogenation device, so that the first distillate oil and gas from the second rectifying tower device and the hydrogenated oil and gas from the hydrogenation device are degassed and separated into rich gas and distillate oil;

[0149] a first rectifying tower device, the rectifying tower device being provided with a distillate inlet, an outlet for a light fraction at the top of the tower, an outlet for a fraction rich in light aromatics, a first recycled fraction outlet, and a second recycled fraction outlet, the distillate inlet being connected to the distillate outlet of the degassing separation device, the first recycled fraction outlet being connected to the first recycled fraction inlet of the first rectifying tower device, the second recycled fraction outlet being connected to a second recycled fraction inlet provided in an optional second reaction zone of the first rectifying tower device, or the second recycled fraction outlet being connected to a second recycled fraction inlet optionally provided in the second section of the hydrogenation device;

[0150] The separation device is provided with an inlet for a light aromatic hydrocarbon-rich fraction and an outlet for a light aromatic hydrocarbon-rich fraction, and the inlet for the light aromatic hydrocarbon-rich fraction is connected to the outlet for the light aromatic hydrocarbon-rich fraction of the first distillation tower device.

[0151] In one embodiment, in the system of the present invention, the first distillation tower device is a dividing wall distillation tower, and at least one dividing wall is provided in the dividing wall distillation tower, and the dividing wall divides the dividing wall distillation tower into four regions: a common stripping section provided at the bottom of the tower, a common distillation section provided at the top of the tower, a pre-separation zone provided in the middle of the tower, and a side distillation zone; the pre-separation zone is provided with a distillate oil inlet, and the distillate oil inlet is connected to the distillate oil outlet of the degassing separation device; the common distillation section is provided with a light fraction outlet at the top of the tower, the side distillation zone is provided with a light aromatics-rich fraction outlet and a first recycle fraction outlet, and the bottom of the common stripping section is provided with a second recycle fraction outlet.

[0152] At this time, the first recycle fraction outlet of the side distillation zone of the distillation tower is connected to the first recycle fraction inlet of the catalytic cracking reactor.

[0153] The second recycle fraction outlet at the bottom of the common stripping section of the dividing wall distillation tower is connected to the second recycle fraction inlet optionally provided on the second reaction zone of the first distillation tower device, or is connected to the second recycle fraction inlet optionally provided on the second section of the hydrogenation device.

[0154] In one embodiment, the system of the present invention further comprises a high-low separation device, wherein the high-low separation device is provided with an inlet for hydrogenated products, a hydrogen outlet, and a hydrogenated oil and gas outlet, and the high-low separation device separates the hydrogenated oil and gas from the hydrogenated products from the hydrogenation device;

[0155] At this time, the hydrogenation product inlet is connected to the hydrogenation product outlet of the hydrogenation device, and the oil and gas inlet of the degassing separation device is connected to the first fraction oil and gas outlet of the second distillation tower device and the hydrogenation oil and gas outlet of the high and low fraction device.

[0156] In one embodiment, in the system of the present invention, the first reaction zone and the second reaction zone of the catalytic cracking reactor are arranged on the same or different reactors;

[0157] Preferably, the reactor is selected from one or more of the following: a constant diameter riser, a constant linear speed riser, a variable diameter riser, a variable linear speed riser, a fluidized bed, and a composite reactor. The composite reactor is composed of a constant diameter or variable diameter riser and a fluidized bed. Preferably, the reactor is a riser reactor.

[0158] More preferably, the first reaction zone and the second reaction zone of the catalytic cracking reactor are arranged on a double-riser reactor, wherein the main riser is the first reaction zone and the secondary riser is the second reaction zone.

[0159] In one embodiment, the system for producing light aromatics of the present invention comprises:

[0160] A hydrogenation unit, wherein the hydrogenation unit is a two-stage hydrogenation unit, the hydrogenation unit is provided with a hydrogen inlet and a hydrogenation product outlet, and the first stage of the hydrogenation unit is provided with a catalytic cracking cycle oil inlet;

[0161] A catalytic cracking unit comprising a catalytic cracking reactor, wherein the catalytic cracking reactor is provided with a first reaction zone and a second reaction zone, wherein the first reaction zone is provided with a first recycled fraction inlet, and the second reaction zone is provided with a second recycled fraction inlet;

[0162] a second distillation tower device, wherein the product distillation tower device is provided with a catalytic cracking reaction product inlet, a first distillate oil and gas outlet, a second distillate oil outlet and an oil slurry outlet, wherein the second distillate oil outlet is connected to the catalytic cracking cycle oil inlet of the hydrogenation unit;

[0163] a degassing and separation device, wherein the degassing and separation device is provided with an oil and gas inlet, a rich gas outlet, and a distillate oil outlet, the oil and gas inlet of the degassing and separation device being connected to the first distillate oil and gas outlet of the second rectifying tower device and the hydrogenated product outlet of the hydrogenation device, so that the first distillate oil and gas from the second rectifying tower device and the hydrogenated oil and gas from the hydrogenation device are degassed and separated into rich gas and distillate oil;

[0164] a first distillation tower device, the distillation tower device being provided with a distillate inlet, a light fraction outlet at the top of the tower, a fraction outlet rich in light aromatics, a first recycled fraction outlet, and a second recycled fraction outlet, the distillate inlet being connected to the distillate outlet of the degassing separation device, the first recycled fraction outlet being connected to the first recycled fraction inlet of the catalytic cracking device, and the second recycled fraction outlet being connected to the second recycled fraction inlet provided in the second reaction zone of the catalytic cracking device;

[0165] The separation device is provided with an inlet for a light aromatic hydrocarbon-rich fraction and an outlet for a light aromatic hydrocarbon-rich fraction, and the inlet for the light aromatic hydrocarbon-rich fraction is connected to the outlet for the light aromatic hydrocarbon-rich fraction of the first distillation tower device.

[0166] In one embodiment, the system for producing light aromatics of the present invention comprises:

[0167] A hydrogenation unit, wherein the hydrogenation unit is a two-stage hydrogenation unit, provided with a hydrogen inlet and a hydrogenation product outlet, a catalytic cracking cycle oil inlet provided on the first stage of the hydrogenation unit, and a second recycled fraction inlet provided on the second stage of the hydrogenation unit;

[0168] A catalytic cracking unit comprising a catalytic cracking reactor, wherein the catalytic cracking reactor is provided with a first reaction zone, and the first reaction zone is provided with a first recycled fraction inlet;

[0169] a second distillation tower device, wherein the product distillation tower device is provided with a catalytic cracking reaction product inlet, a first distillate oil and gas outlet, a second distillate oil outlet and an oil slurry outlet, wherein the second distillate oil outlet is connected to the catalytic cracking cycle oil inlet of the hydrogenation unit;

[0170] a degassing and separation device, wherein the degassing and separation device is provided with an oil and gas inlet, a rich gas outlet, and a distillate oil outlet, the oil and gas inlet of the degassing and separation device being connected to the first distillate oil and gas outlet of the second rectifying tower device and the hydrogenated product outlet of the hydrogenation device, so that the first distillate oil and gas from the second rectifying tower device and the hydrogenated oil and gas from the hydrogenation device are degassed and separated into rich gas and distillate oil;

[0171] a first distillation tower device, the distillation tower device being provided with a distillate inlet, a light fraction outlet at the top of the tower, a fraction outlet rich in light aromatics, a first recycled fraction outlet, and a second recycled fraction outlet, the distillate inlet being connected to the distillate outlet of the degassing separation device, the first recycled fraction outlet being connected to the first recycled fraction inlet of the catalytic cracking device, and the second recycled fraction outlet being connected to the second recycled fraction inlet provided on the second section of the hydrogenation device;

[0172] The separation device is provided with an inlet for a light aromatic hydrocarbon-rich fraction and an outlet for a light aromatic hydrocarbon-rich fraction, and the inlet for the light aromatic hydrocarbon-rich fraction is connected to the outlet for the light aromatic hydrocarbon-rich fraction of the first distillation tower device.

[0173] The present invention will be described with reference to FIG1 . As a preferred embodiment of the present invention, catalytic cracking circulating oil with a distillation range of 220 to 370°C enters the hydrogenation unit 100 through inlet 101, contacts with hydrogen and hydrogenation catalyst to undergo a two-stage hydrogenation reaction, and the hydrogenated product enters the high-low fractionation unit through pipeline 102 to separate circulating hydrogen and low-fraction gas. The hydrogenated oil and gas then enter the degassing separation unit through pipeline 103, and the separated distillate oil enters the pre-separation zone A of the next-door distillation tower 200 as the second distillation tower unit through pipeline 201. The top common distillation section B separates a light fraction with a temperature less than 60°C and flows out through pipeline 202. The bottom common stripping section D separates a second recycled fraction with a temperature greater than 240°C and enters the feedstock oil inlet 304 of the secondary riser 302 of the catalytic cracking unit 300 through pipeline 205. The catalytic cracking reaction is carried out as the feedstock for the second reaction zone. A light aromatics-rich fraction with a distillation range of 80-150°C is separated from the upper portion of the side distillation zone C and flows out via pipeline 203. A first recycled fraction with a distillation range of 150-240°C is separated from the lower portion and enters the feedstock oil inlet 303 of the first riser 301 of the catalytic cracking unit 300 via pipeline 204 as the feedstock for the first reaction zone. The resulting catalytic cracking products are then passed through product pipeline 305 to the second distillation tower 400 for separation. The separated first fraction oil and gas with a distillation range below 240°C are combined with the hydrogenation oil and gas pipeline 103 via pipeline 401 and enter the degassing separation unit. The second fraction oil with a distillation range of 240-350°C is returned to the inlet of the hydrogenation unit 101 via pipeline 402 for recycling. The catalytic cracking unit utilizes a dual-riser reactor, with the first reaction zone located in the first riser and the second reaction zone located in the second riser. A separation unit is then used to separate light aromatics from the light aromatics-rich fraction.

[0174] Referring to FIG2 , the present invention will be described. As another preferred embodiment of the present invention, catalytic cracking cycle oil with a boiling range of 220 to 370°C enters the hydrogenation device 100 through the inlet 101, contacts with hydrogen and hydrogenation catalyst to carry out a two-stage hydrogenation reaction, and the hydrogenated product enters the high-low fractionation device through the 102 pipeline to separate the circulating hydrogen and low-fraction gas, and the hydrogenated oil gas enters the degassing separation device through the pipeline 103, and the separated distillate oil enters the pre-separation zone A of the next-door distillation tower 200 as the second distillation tower device through the pipeline 201, and the light fraction below 60°C is separated from the top common distillation section B at the top of the tower and flows out through the pipeline 202, and the second recycled fraction above 280°C is separated from the bottom common stripping section D at the bottom of the tower and enters the hydrogenation reaction of the hydrogenation device 100 through the pipeline 205. In the second stage, a light aromatics-rich fraction with a distillation range of 80-150°C is separated from the upper portion of the side-stream distillation zone C and flows out via pipeline 203. A first recycled fraction with a distillation range of 150-280°C is separated from the lower portion and flows via pipeline 204 to the feedstock inlet 303 of the first riser 301 of the catalytic cracking unit 300, where it serves as the feedstock for the first reaction zone and undergoes catalytic cracking. The resulting catalytic cracking products flow through product pipeline 305 and enter the second distillation tower 400 for separation. The separated first distillate oil and gas with a distillation range below 240°C are combined with the hydrogenation oil and gas pipeline 103 via pipeline 401 and enter the degassing separation unit. The second distillate oil with a distillation range of 240-350°C is returned to the inlet 101 of the hydrogenation unit 100 via pipeline 402 for recycling. The catalytic cracking unit utilizes a single riser reactor. A separation unit is then used to separate light aromatics from the light aromatics-rich fraction.

[0175] In one embodiment of the present invention, the following scheme I is provided:

[0176] Scheme I-1. A method for producing more light aromatics, characterized in that the method comprises:

[0177] (1) Catalytic cracking feedstock oil enters the first reaction zone of the catalytic cracking unit and contacts with the catalyst for catalytic cracking to obtain first reaction oil gas;

[0178] (2) After the first reaction oil and gas are separated from the catalyst gas and solid, they enter the first distillation tower to separate into the first distillate oil and gas, the second distillate oil and oil slurry;

[0179] (3) The second distillate oil obtained in step (2) is subjected to hydrogenation treatment, and hydrogen is separated from the hydrogenated product obtained by the hydrogenation treatment to obtain hydrogenated oil gas, and the hydrogenated oil gas and the first distillate oil gas are degassed and separated to obtain rich gas and the first distillate oil;

[0180] (4) The first distillate oil obtained in step (3) enters a second distillation tower to separate and obtain a naphtha fraction rich in light aromatics, a heavy aromatics fraction, and a diesel fraction;

[0181] (5) The heavy aromatics fraction obtained in step (4) is fed into the second reaction zone of a catalytic cracking unit, and the diesel fraction is fed into the first reaction zone of a catalytic cracking unit for catalytic conversion;

[0182] Scheme I-2. The method according to Scheme I-1 is characterized in that, in step (1), the catalytic cracking feedstock oil is selected from at least one of hydrotreated diesel, straight-run diesel, straight-run wax oil, hydrogenated wax oil, hydrocracking tail oil, and hydrotreated residual oil, and the diesel fraction separated from the second distillation tower.

[0183] Scheme I-3. The method according to Scheme I-1, characterized in that, in step (1), the conditions for catalytic conversion in the first reaction zone are:

[0184] The weight hourly space velocity of oil and gas is 4 to 20h -1 , preferably 6 to 12 hours -1 ;

[0185] The reaction time is 1 to 10 seconds, preferably 2 to 6 seconds; and / or

[0186] The weight ratio of agent to oil is 4 to 14, preferably 6 to 10;

[0187] Scheme I-4. The method according to Scheme I-1, characterized in that, in step (2), the initial distillation point of the first fraction of oil and gas is 0 to 40 ° C, and the final distillation point is 220 to 250 ° C;

[0188] The second distillate oil has an initial boiling point of 220-250° C. and a final boiling point of 340-360° C.; and / or

[0189] The initial boiling point of the slurry oil is 340-360°C;

[0190] Scheme I-5. The method according to Scheme I-1, characterized in that, in step (3), the hydrotreating is hydrofining, hydroreforming or hydrocracking; and / or

[0191] The initial boiling point of the hydrogenated oil gas is 0-60°C, and the final boiling point is 280-350°C;

[0192] Scheme I-6. The method according to Scheme I-1, characterized in that, in step (4), the initial distillation point of the naphtha fraction is 20 to 40 ° C, and the final distillation point is 140 to 160 ° C;

[0193] The initial boiling point of the heavy aromatic fraction is 140-160°C, and the final boiling point is 210-250°C;

[0194] The initial boiling point of the diesel fraction is 210-250°C, and the final boiling point is 280-350°C;

[0195] Scheme I-7. The method according to Scheme I-1, characterized in that, in step (5), the conditions for catalytic conversion in the second reaction zone are:

[0196] The weight hourly space velocity of oil and gas is 4 to 20h -1 , preferably 8 to 16 hours -1 ;

[0197] The reaction time is 0.1 to 5 s, preferably 0.5 to 3 s; and / or

[0198] The agent-oil weight ratio is 6 to 26, preferably 10 to 22;

[0199] Scheme I-8. The method according to Scheme I-1, characterized in that, in step (2), the top pressure of the first distillation tower is 0.1 to 0.5 MPa, preferably 0.2 to 0.4 MPa;

[0200] Scheme I-9. The method according to Scheme I-1, characterized in that, in step (4), the second distillation tower is a conventional distillation tower or a dividing wall distillation tower;

[0201] Preferably, the top pressure of the second distillation tower is 0.1 to 1.0 MPa, more preferably 0.2 to 0.7 MPa; and / or

[0202] The top temperature of the second distillation tower is 50-120°C, more preferably 60-100°C;

[0203] Further preferably, the number of trays in the second distillation tower is 30 to 50, preferably 35 to 45; the naphtha fraction rich in light aromatics is extracted from the 10th to 40th tray, preferably the 15th to 30th tray, at the top of the side line of the second distillation tower;

[0204] Option I-10. The method according to Option I-9, characterized in that the second distillation tower is a dividing wall distillation tower, and at least one dividing wall is provided in the dividing wall distillation tower, the dividing wall dividing the dividing wall distillation tower into four regions: a pre-separation zone, a common distillation section, a common stripping section, and a side distillation section, and the oil and gas feed position is located in the pre-separation zone;

[0205] Preferably, the number of plates in the pre-separation zone of the dividing wall distillation tower is 10 to 40, preferably 20 to 35; the number of plates in the common distillation section is 3 to 20, preferably 5 to 15; the number of plates in the common stripping section is 2 to 10, preferably 3 to 8; and / or the number of plates in the side distillation zone is 10 to 45, preferably 25 to 40;

[0206] More preferably, the mass flow ratio of the liquid phase from the common distillation section into the pre-separation zone and the side distillation zone is 0.1 to 10, preferably 0.5 to 5; and / or

[0207] More preferably, the mass flow ratio of the gas phase entering the pre-separation zone and the side distillation zone from the common stripping section is 1 to 15, preferably 2 to 10;

[0208] Scheme I-11. The method according to Scheme I-1, characterized in that in step (1), the catalyst is a catalytic cracking catalyst, the catalyst comprising 10 to 60 parts by weight of a molecular sieve, 1 to 40 parts by weight of a binder and 1 to 90 parts by weight of a carrier;

[0209] Wherein, the molecular sieve is selected from one or more of ZSM molecular sieve, Y molecular sieve, HY molecular sieve, USY molecular sieve, and Beta molecular sieve. Optionally, the molecular sieve contains rare earth, and the rare earth is one or more of La, Ce, Pr, and Nd.

[0210] The binder is a silicon oxide binder and / or an aluminum oxide binder; and / or

[0211] The carrier is selected from one or more of silicon dioxide, kaolin, montmorillonite, diatomaceous earth, halloysite, saponite, rectorite, sepiolite, attapulgite, hydrotalcite and bentonite;

[0212] Scheme I-12. A system for producing more light aromatics, comprising:

[0213] Catalytic cracking unit, including:

[0214] A reactor is provided with a first reaction zone and a second reaction zone, wherein the first reaction zone and the second reaction zone are arranged in series, and the second reaction zone is arranged upstream of the first reaction zone, and the catalyst contacts the feedstock oil through the second reaction zone and the first reaction zone in sequence;

[0215] a gas-solid separator, the gas-solid separator being connected to the outlet of the oil-agent mixture of the first reaction zone of the reactor, so as to separate the first reaction oil gas from the first reaction zone from the catalyst;

[0216] a first distillation tower, wherein the first distillation tower is provided with a first distillate oil and gas outlet, a second distillate oil outlet, and an oil slurry outlet, and the inlet of the first distillation tower is connected to the outlet of the catalytic cracking unit through a pipeline;

[0217] A diesel hydrotreating unit, wherein the diesel hydrotreating unit is provided with a diesel inlet, a hydrogen inlet, and a hydrogenation product outlet, and the diesel inlet of the diesel hydrotreating unit is connected to the second distillate oil outlet of the first distillation tower;

[0218] A high-low separation device is provided with a hydrogenation product inlet, a hydrogen outlet, and a hydrogenated oil and gas outlet, the hydrogenation product inlet is connected to the hydrogenation product outlet of the diesel hydrotreating device, and the high-low separation device separates the hydrogenated oil and gas from the hydrogenation product from the diesel hydrotreating device;

[0219] a degassing and separation device, wherein the degassing and separation device is provided with an oil and gas inlet, a rich gas outlet, and a first distillate oil outlet, the oil and gas inlet of the degassing and separation device being connected to the first distillate oil and gas outlet of the first rectifying tower and the hydrogenated oil and gas outlet of the high-low fractionation device, so that the first distillate oil and gas from the first rectifying tower and the hydrogenated oil and gas from the high-low fractionation device are degassed and separated into rich gas and first distillate oil;

[0220] a second distillation tower, wherein the second distillation tower is provided with a first distillate oil inlet, a naphtha fraction outlet, a heavy aromatic fraction outlet, and a diesel fraction outlet, wherein the first distillate oil inlet is connected to the first distillate oil outlet of the degassing separation device, the heavy aromatic fraction outlet is connected to the second reaction zone of the reactor of the catalytic cracking device, and the diesel fraction outlet is connected to the first reaction zone of the reactor of the catalytic cracking device;

[0221] Option I-13. The system according to Option I-12, wherein the reactor is selected from one or two of the following: a constant diameter riser, a constant linear velocity riser, a variable diameter riser, a variable linear velocity riser, a fluidized bed, and a composite reactor, wherein the composite reactor comprises a constant diameter or variable diameter riser and a fluidized bed.

[0222] The first reaction zone and the second reaction zone are arranged in the same or different reactors;

[0223] Option I-14. The system of Option I-12, wherein the first distillation tower is a conventional distillation tower having 20 to 40 trays, preferably 25 to 35 trays, and the second distillate is produced from the top of the sideline of the first distillation tower at trays 5 to 30, preferably 10 to 25.

[0224] Preferably, the first distillation column is heated by gaseous feed and the bottom of the column is stripped by steam; and / or

[0225] Preferably, the side line of the first distillation tower can be provided with one or more mid-section circulation heat extraction.

[0226] Scheme I-15. A system according to any one of Schemes I-12 to I-14, characterized in that the first distillation tower and / or the second distillation tower is a constant diameter distillation tower or a variable diameter distillation tower.

[0227] In one embodiment of the present invention, the following scheme II is provided:

[0228] Scheme II-1. A catalytic conversion method for producing light aromatics, characterized in that the method comprises:

[0229] (1) Catalytic cracking cycle oil enters the hydrogenation unit and contacts hydrogen in the presence of a hydrogenation catalyst to undergo a two-stage hydrogenation reaction to obtain hydrogenated oil gas;

[0230] (2) The hydrogenated oil and gas enter the distillation tower to separate into gas, hydrogenated light naphtha, hydrogenated heavy naphtha, hydrogenated light diesel and hydrogenated heavy diesel;

[0231] (3) the hydrogenated heavy diesel obtained in step (2) is returned to the second stage of the hydrogenation unit for recycling, and the hydrogenated light diesel obtained in step (2) is fed into a catalytic cracking unit to undergo a catalytic cracking reaction in the presence of a catalytic cracking catalyst, and the catalytic cracking reaction products are separated to obtain dry gas, liquefied gas, catalytic naphtha rich in light aromatics, a catalytic heavy aromatics fraction, catalytic cracking cycle oil, and slurry oil;

[0232] (4) the catalytic cracking cycle oil obtained in step (3) is returned to the first stage of the hydrogenation unit for recycling; optionally, the catalytic heavy aromatics fraction obtained in step (3) is returned to the catalytic cracking unit for recycling conversion;

[0233] Scheme II-2. The catalytic conversion method according to Scheme II-1, characterized in that, in step (1), the hydrotreatment reaction conditions are:

[0234] Preferably, the first stage hydrotreatment reaction conditions are: hydrogen partial pressure 5.0-20.0 MPa, reaction temperature 300-450°C, volume space velocity 0.5-5.0 h -1 , Hydrogen to oil volume ratio 300~1600Nm 3 / m 3 and / or

[0235] Preferably, the second stage hydrotreatment reaction conditions are hydrogen partial pressure 5.0-20.0 MPa, reaction temperature 350-500°C, volume space velocity 0.1-3.0 h -1 , Hydrogen to oil volume ratio 400~2000Nm 3 / m 3 ;

[0236] Scheme II-3. The catalytic conversion method according to Scheme II-1, characterized in that in step (1), in the two-stage hydrotreatment reaction, the hydrogenation catalysts are the same or different;

[0237] Preferably, the hydrogenation catalyst comprises an active metal component and a support, wherein the active metal component is selected from one or more of Group VIB metals and / or Group VIII non-noble metals, and the support is selected from one or more of alumina, silica, and amorphous silica-alumina;

[0238] More preferably, the active metal component is selected from any one of the following metal combinations: nickel-tungsten, nickel-tungsten-cobalt, nickel-molybdenum, cobalt-molybdenum;

[0239] Scheme II-4. The catalytic conversion method according to Scheme II-1, characterized in that in step (2), at least one dividing wall is provided in the bulkhead distillation tower, the dividing wall dividing the bulkhead distillation tower into four regions: a pre-separation zone, a common distillation section, a common stripping section, and a side distillation zone, and the hydrogenated oil and gas feed position is located in the pre-separation zone;

[0240] Scheme II-5. The method according to Scheme II-4, characterized in that the top pressure of the dividing wall distillation tower is 0.1 to 1.0 MPa, preferably 0.2 to 0.7 MPa;

[0241] The tower top temperature is 50-120°C, preferably 60-100°C;

[0242] The mass flow ratio of the liquid phase from the common distillation section into the pre-separation zone and the side distillation zone is 0.1 to 10, preferably 0.5 to 5;

[0243] The mass flow ratio of the gas phase from the common stripping section into the pre-separation zone and the side-line distillation zone is 1 to 15, preferably 2 to 10; and / or

[0244] The ratio of the total mass flow rate of the gas phase at the extraction position of the pre-separation zone to the mass flow rate of the extracted gas phase is 10 to 150, preferably 30 to 120;

[0245] Scheme II-6. The catalytic conversion method according to Scheme II-4, characterized in that the number of trays in the pre-separation zone in the dividing wall distillation tower is 10 to 40, preferably 20 to 35;

[0246] The number of plates in the common rectifying section is 2 to 10, preferably 3 to 8; the number of plates in the common stripping section is 2 to 10, preferably 3 to 8; and / or

[0247] The number of trays in the side distillation zone is 10 to 45, preferably 25 to 40;

[0248] Scheme II-7. The catalytic conversion method according to Scheme II-6, characterized in that the side distillation zone is heavy, and the extraction and return positions are the 20th to 45th trays of the side distillation zone, preferably the 30th to 40th trays, and preferably, the side distillation zone is further provided with a mid-stage circulation heat extraction; and / or

[0249] Preferably, the pre-separation zone is provided with a gas phase extraction, the extraction position is located at the 20th to 40th plate of the pre-separation zone, preferably the 25th to 35th plate, and the return tower position is the 30th to 45th plate of the side distillation zone, preferably the 30th to 40th plate;

[0250] Scheme II-8. The catalytic conversion method according to any one of Schemes II-4 to II-7, characterized in that in step (2), the hydrogenated light naphtha is separated from the common distillation section of the dividing wall distillation column, and has an initial distillation point of 20-40°C and a final distillation point of 55-65°C;

[0251] The hydrogenated heavy naphtha is produced from the 5th to 30th plates of the side distillation zone of the dividing wall distillation tower, preferably the 10th to 20th plates, with an initial distillation point of 55 to 65°C and a final distillation point of 140 to 170°C;

[0252] The hydrogenated light diesel is produced from any position below the hydrogenated heavy naphtha production position in the side distillation zone, preferably the 25th to 35th plate, with an initial boiling point of 140 to 170° C. and a final boiling point of 230 to 280° C.; and / or

[0253] The hydrogenated heavy diesel is separated from the lower part of the common stripping section of the distillation tower, and the initial boiling point is 230-280°C;

[0254] Scheme II-9. The catalytic conversion method according to Scheme II-1, characterized in that, in step (3), the initial boiling point of the catalytic naphtha is 20 to 40 ° C, and the final boiling point is 140 to 160 ° C;

[0255] The initial boiling point of the catalytic heavy aromatic fraction is 140-160° C., and the final boiling point is 210-250° C.; and / or

[0256] The initial boiling point of the catalytic cracking cycle oil is 210-250°C, and the final boiling point is 320-360°C;

[0257] Scheme II-10. The catalytic conversion method according to Scheme II-1, characterized in that, in step (3), the conditions of the catalytic cracking reaction include:

[0258] The weight hourly space velocity of oil and gas is 4 to 20h -1 , preferably 8 to 16 hours -1 ;

[0259] The catalytic cracking reaction temperature is 520-650°C, preferably 560-640°C; and / or

[0260] The agent-oil weight ratio is 6 to 26, preferably 8 to 20;

[0261] Scheme II-11. The catalytic conversion method according to Scheme II-1, characterized in that, in step (3), the catalytic cracking catalyst contains 10-60 parts by weight of a molecular sieve, 1-40 parts by weight of a binder, and 1-90 parts by weight of a carrier;

[0262] The molecular sieve is selected from one or more of ZSM molecular sieve, Y molecular sieve, HY molecular sieve, USY molecular sieve and Beta molecular sieve. Optionally, the molecular sieve contains rare earth, and the rare earth is one or more of La, Ce, Pr and Nd;

[0263] The binder is selected from silicon oxide binder and / or aluminum oxide binder;

[0264] The carrier is selected from one or more of silicon dioxide, kaolin, montmorillonite, diatomaceous earth, halloysite, saponite, rectorite, sepiolite, attapulgite, hydrotalcite and bentonite;

[0265] Scheme II-12. A catalytic conversion system for producing light aromatics, comprising:

[0266] A hydrogenation unit, wherein the hydrogenation unit is a two-stage hydrogenation unit, provided with a hydrogen inlet and a hydrogenated oil gas outlet, a catalytic cracking cycle oil inlet being provided on the first section of the hydrogenation unit, and a hydrogenated heavy diesel inlet being provided on the second section of the hydrogenation unit;

[0267] A bulkhead distillation tower, wherein at least one dividing wall is provided in the bulkhead distillation tower, and the dividing wall divides the bulkhead distillation tower into four regions: a common stripping section provided at the bottom of the tower, a common distillation section provided at the top of the tower, a pre-separation zone provided in the middle section of the tower, and a side-line distillation zone; the pre-separation zone is provided with a hydrogenated oil and gas inlet, and the hydrogenated oil and gas inlet is connected to the hydrogenated oil and gas outlet of the hydrogenation unit; the common distillation section is provided with a low-boiling point fraction outlet, the side-line distillation zone is provided with a hydrogenated heavy naphtha outlet rich in light aromatics and a hydrogenated light diesel outlet, and the bottom of the common stripping section is provided with a hydrogenated heavy diesel outlet, and the hydrogenated heavy diesel outlet is connected to the hydrogenated heavy diesel inlet on the second section of the hydrogenation unit;

[0268] A catalytic cracking unit, the catalytic cracking unit being provided with a hydrogenated light diesel inlet, a dry gas outlet, a liquefied gas outlet, a catalytic naphtha outlet rich in light aromatics, a catalytic heavy aromatics outlet, a catalytic cracking cycle oil outlet, and a slurry oil outlet, the catalytic heavy aromatics outlet being connected to a catalytic heavy aromatics fractionation pipeline, the catalytic cracking cycle oil outlet being connected to a catalytic cracking cycle oil inlet on the first section of the hydrogenation unit, and optionally, the catalytic heavy aromatics fractionation pipeline being connected to the hydrogenated light diesel inlet;

[0269] Scheme II-13. The catalytic conversion system according to Scheme II-12, characterized in that the catalytic cracking unit is a single reactor or a dual reactor;

[0270] Preferably, the reactor is selected from one or more of the following: a constant diameter riser, a constant linear velocity riser, a variable diameter riser, a variable linear velocity riser, a fluidized bed, and a composite reactor, wherein the composite reactor is composed of a constant diameter or variable diameter riser and a fluidized bed.

[0271] In one embodiment of the present invention, the present invention provides the following Scheme III:

[0272] Scheme III-1. A method for producing light aromatics, characterized in that the method comprises:

[0273] (1) Catalytic cracking cycle oil is subjected to a hydrogenation reaction in the presence of hydrogen and a hydrogenation catalyst to obtain hydrogenated oil gas;

[0274] (2) The hydrogenated oil and gas are separated in a dividing wall fractionation tower to obtain gas, hydrogenated light naphtha, hydrogenated heavy naphtha rich in light aromatics, hydrogenated light diesel and hydrogenated heavy diesel;

[0275] (3) the hydrogenated light diesel obtained in step (2) enters the first reaction zone of a catalytic cracking unit, and the hydrogenated heavy diesel enters the second reaction zone of a catalytic cracking unit, and catalytic cracking reactions are respectively carried out in the presence of a catalytic cracking catalyst, and the reaction products are separated to obtain dry gas, liquefied gas, catalytic naphtha, catalytic heavy aromatics fraction, catalytic cracking cycle oil and slurry oil;

[0276] (4) the catalytic cracking cycle oil obtained in step (3) is returned to the hydrogenation unit for recycling; optionally, the catalytic heavy aromatics fraction obtained in step (3) is returned to the third reaction zone of the catalytic cracking unit for recycling conversion, wherein the third reaction zone is located upstream of the first reaction zone and / or the second reaction zone;

[0277] Scheme III-2. The method according to Scheme III-1, characterized in that, in step (1), the hydrotreatment reaction conditions are: hydrogen partial pressure 5.0~20.0MPa, reaction temperature 330~450℃, volume space velocity 0.1~3.0h -1 , Hydrogen to oil volume ratio 350~2000Nm 3 / m 3 and / or

[0278] The initial boiling point of the catalytic cracking cycle oil is 210-250°C, and the final boiling point is 340-400°C;

[0279] Scheme III-3. The method according to Scheme III-1, characterized in that in step (2), at least one dividing wall is provided in the bulkhead distillation column, the dividing wall dividing the bulkhead distillation column into four regions: a pre-separation zone, a common distillation section, a common stripping section, and a side distillation section, and the feed position of the hydrogenated oil gas is located in the pre-separation zone;

[0280] Scheme III-4. The method according to Scheme III-3, characterized in that, in step (2), the number of plates in the pre-separation zone in the dividing wall distillation tower is 10 to 40, preferably 20 to 35;

[0281] The number of plates in the common distillation section is 2 to 10, preferably 3 to 8;

[0282] The number of trays in the common stripping section is 2 to 10, preferably 3 to 8; and / or

[0283] The number of trays in the side distillation zone is 10 to 45, preferably 25 to 40;

[0284] Preferably, the side distillation zone of the next-wall distillation tower is provided with a mid-stage circulation heat extraction, which is the 20th to 45th tray of the side distillation zone, more preferably the 30th to 40th tray; and / or

[0285] Preferably, the pre-separation zone of the dividing wall distillation tower is provided with a gas phase extraction, the extraction position is located at the 20th to 40th plate of the pre-separation zone, more preferably the 25th to 35th plate, and the return tower position is the 30th to 45th plate of the side line distillation zone, more preferably the 35th to 40th plate;

[0286] Scheme III-5. The method according to Scheme III-3, characterized in that the top pressure of the dividing wall distillation tower is 0.1 to 1.0 MPa, preferably 0.2 to 0.7 MPa;

[0287] The tower top temperature is 50-120°C, preferably 60-100°C;

[0288] The mass flow ratio of the liquid phase from the common distillation section into the pre-separation zone and the side distillation zone is 0.1 to 10, preferably 0.5 to 5;

[0289] The mass flow ratio of the gas phase from the common stripping section into the pre-separation zone and the side-line distillation zone is 1 to 15, preferably 2 to 10; and / or

[0290] The ratio of the total gas phase mass flow rate at the extraction position of the pre-separation zone to the extracted gas phase mass flow rate is 10 to 150, preferably 30 to 120;

[0291] Scheme III-6. The method according to Scheme III-3, characterized in that the hydrogenated light naphtha is separated from the common distillation section of the dividing wall distillation column, with an initial distillation point of 20 to 40°C and a final distillation point of 55 to 65°C;

[0292] The heavy naphtha is produced from the 5th to 30th plates of the side distillation zone of the dividing wall distillation tower, preferably the 10th to 20th plates, with an initial distillation point of 55 to 65°C and a final distillation point of 140 to 170°C;

[0293] The hydrogenated light diesel is produced from any position below the heavy naphtha production position in the side distillation zone, preferably the 25th to 35th plate, with an initial boiling point of 140 to 170°C and a final boiling point of 220 to 250°C; and / or

[0294] The hydrogenated heavy diesel is separated from the lower part of the common stripping section of the distillation tower, and the starting point temperature of the distillation range is any value between 220 and 250°C;

[0295] Scheme III-7. The method according to Scheme III-1, characterized in that, in step (3), the conditions for catalytic conversion in the first reaction zone include:

[0296] The weight hourly space velocity of oil and gas is 4 to 20h -1 , preferably 6 to 12 hours -1 ;

[0297] The reaction time is 0.5 to 8 s, preferably 1 to 5 s; and / or

[0298] The agent-oil weight ratio is 6 to 26, preferably 8 to 12;

[0299] Scheme III-8. The method according to Scheme III-1, characterized in that, in step (3), the conditions for catalytic conversion in the second reaction zone include:

[0300] The weight hourly space velocity of oil and gas is 4 to 20h -1 , preferably 8 to 12 hours -1 ;

[0301] The reaction time is 1 to 10 seconds, preferably 2 to 6 seconds; and / or

[0302] The weight ratio of agent to oil is 4 to 14, preferably 6 to 10;

[0303] Scheme III-9. The method according to Scheme III-1, characterized in that, in step (4), the conditions for catalytic conversion in the third reaction zone include:

[0304] The weight hourly space velocity of oil and gas is 4 to 20h -1 , preferably 8 to 15 hours -1 ;

[0305] The reaction time is 0.1 to 5 s, preferably 0.5 to 3 s; and / or

[0306] The agent-oil weight ratio is 6 to 26, preferably 10 to 22;

[0307] Scheme III-10. The method according to Scheme III-1, characterized in that, in step (3), the initial boiling point of the catalytic naphtha is 20 to 40 ° C and the final boiling point is 140 to 160 ° C; and / or

[0308] The initial boiling point of the heavy aromatic fraction is 140-160°C, and the final boiling point is 210-250°C;

[0309] Scheme III-11. The method according to Scheme III-1, characterized in that the hydrogenation catalyst comprises an active metal and a support, wherein the active metal component is selected from one or more of Group VIB metals and / or Group VIII non-noble metals, and the support is selected from any one or more of alumina, silica, and amorphous silica-alumina.

[0310] Preferably, the active metal component is selected from any one of the following metal combinations: nickel-tungsten, nickel-tungsten-cobalt, nickel-molybdenum, cobalt-molybdenum;

[0311] Scheme III-12. The method according to Scheme III-1, characterized in that the catalytic cracking catalyst contains 10-60 parts by weight of a molecular sieve, 1-40 parts by weight of a binder and 1-90 parts by weight of a carrier;

[0312] The molecular sieve is selected from one or more of ZSM molecular sieve, Y molecular sieve, HY molecular sieve, USY molecular sieve and Beta molecular sieve. Optionally, the molecular sieve contains rare earth, and the rare earth is one or more of La, Ce, Pr and Nd;

[0313] The binder is selected from silicon oxide binder and / or aluminum oxide binder;

[0314] The carrier is selected from one or more of silicon dioxide, kaolin, montmorillonite, diatomaceous earth, halloysite, saponite, rectorite, sepiolite, attapulgite, hydrotalcite and bentonite;

[0315] Scheme III-13. A system for producing light aromatics, comprising:

[0316] A hydrogenation unit, wherein the hydrogenation unit is provided with a hydrogen inlet, a catalytic cracking circulating oil inlet and a hydrogenated oil and gas outlet;

[0317] A bulkhead distillation tower, wherein at least one dividing wall is provided in the bulkhead distillation tower, and the dividing wall divides the bulkhead distillation tower into four regions: a common stripping section provided at the bottom of the tower, a common distillation section provided at the top of the tower, a pre-separation zone provided in the middle of the tower, and a side-line distillation zone; the pre-separation zone is provided with a hydrogenated oil and gas inlet, and the hydrogenated oil and gas inlet is connected to the hydrogenated oil and gas outlet of the hydrogenation unit; the common distillation section is provided with a low-boiling point fraction outlet, the side-line distillation zone is provided with a hydrogenated heavy naphtha outlet rich in light aromatics and a hydrogenated light diesel outlet, and the bottom of the common stripping section is provided with a hydrogenated heavy diesel outlet;

[0318] Catalytic cracking unit, including:

[0319] A catalytic cracking reactor, wherein the catalytic cracking reactor is provided with a first reaction zone and a second reaction zone, the first reaction zone is provided with a hydrogenated light diesel inlet, the second reaction zone is provided with a hydrogenated heavy diesel inlet, and the hydrogenated heavy diesel inlet is connected to the hydrogenated heavy diesel outlet at the bottom of the common stripping section of the distillation tower;

[0320] a product separation device, wherein the product separation device is provided with a catalytic cracking reaction product inlet, a dry gas outlet, a liquefied gas outlet, a catalytic naphtha outlet, a catalytic heavy aromatics outlet, a catalytic cracking cycle oil outlet, and an oil slurry outlet, wherein the catalytic cracking cycle oil outlet is connected to the catalytic cracking cycle oil inlet of the hydrogenation unit;

[0321] Optionally, the catalytic cracking reactor is further provided with a third reaction zone, the third reaction zone is provided with a catalytic heavy aromatics inlet, and the third reaction zone is located upstream of the first reaction zone and / or the second reaction zone, and the catalytic heavy aromatics inlet is connected to the catalytic heavy aromatics outlet of the product separation device;

[0322] Option III-14. The system according to Option III-13, characterized in that the third reaction zone, the first reaction zone, and the second reaction zone of the catalytic cracking reactor are arranged in the same or different reactors;

[0323] Preferably, the riser reactor is selected from one or more of the following: a constant diameter riser, a constant linear speed riser, a variable diameter riser, a variable linear speed riser, a fluidized bed, and a composite reactor, wherein the composite reactor is composed of a constant diameter or variable diameter riser and a fluidized bed;

[0324] More preferably, the third reaction zone, the first reaction zone and the second reaction zone of the catalytic cracking reactor are arranged on a double riser reactor, wherein the main riser is the first reaction zone, and the third reaction zone and the second reaction zone are arranged on the secondary riser, and the third reaction zone is located upstream of the second reaction zone.

[0325] Example

[0326] 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.

[0327] The hydrogenation catalysts used in the examples and comparative examples were purchased from brands FZC and FC-70, which are hydrorefining catalysts and hydrocracking catalysts, respectively.

[0328] The catalytic cracking catalyst used in the examples and comparative examples was purchased from SLA-1.

[0329] The catalysts are all produced by Sinopec Catalyst Co., Ltd.

[0330] The properties of the catalytic cracking cycle oil used in the Examples and Comparative Examples are shown in Table 1.

[0331] Table 1

[0332] Example 1

[0333] 1 , the hydrogenation unit 100 utilizes an upper and lower fixed-bed reactor connected in series, with FZC and FC-70 hydrogenation catalysts placed therein, respectively. Catalytic cracking cycle oil enters the hydrogenation unit 100 through inlet 101 and hydrogen enters the hydrogenation unit 100 through the hydrogen inlet, and the reactors sequentially pass through the upper and lower sections to contact the hydrogenation catalysts for a two-stage hydrogenation reaction.

[0334] A dividing wall E is installed within the distillation tower 200, separating the tower radially. The pre-separation zone A, common distillation section B, common stripping section D, and side distillation section C have 34, 6, 4, and 36 trays, respectively. The top reflux ratio is 6. The feed is located at the 17th tray from the top of the pre-separation zone. The light aromatics-rich fraction B is withdrawn from the 15th tray from the top of the side distillation section C, and the first recycled fraction is withdrawn from the 32nd tray from the top of the side distillation section C. The top pressure of the distillation tower 200 is 0.4 MPa, and the top temperature is 67°C. The side distillation section of the distillation tower is equipped with mid-stage heat extraction, with the side distillation withdrawal and return points located at trays 31 and 30 of the side distillation zone. The pre-separation zone of the distillation tower is equipped with a gas phase withdrawal point located at tray 25 of the pre-separation zone and the return point at tray 30 of the side distillation zone. The liquid mass flow ratio from the common distillation section to the pre-separation zone and the side distillation zone was 2.64; the gas mass flow ratio from the common stripping section to the pre-separation zone and the side distillation zone was 4.22; and the ratio of the total gas mass flow rate at the pre-separation zone withdrawal point to the withdrawal gas mass flow rate was 69.36. The catalytic cracking unit 300 employed a dual-riser reactor, using SLA-1 catalytic cracking catalyst for both. The primary riser served as the first reaction zone, and the secondary riser served as the second reaction zone.

[0335] The catalytic cracking cycle oil with a distillation range of 240-367°C is passed through the catalytic cracking cycle oil inlet 101 and the hydrogen is passed through the hydrogenation unit 100 in sequence to contact and react with the hydrogenation catalyst. The hydrogen partial pressure is 8.0 MPa, the upper and lower reaction temperatures are 360°C and 380°C respectively, and the volume space velocity is 2.0 h -1 and 1.0h -1 The volume ratio of hydrogen to oil is 800 and 1300 respectively. The hydrogenated product enters the high-low fraction device through pipeline 102 to separate the circulating hydrogen and low-fraction gas. The hydrogenated oil and gas enter the degassing separation device through pipeline 103. The separated distillate oil enters the distillation tower 200 through pipeline 201 for separation. Four fractions with boiling points from low to high are separated. The light fraction below 60°C is extracted from the top of the tower. The second recycled fraction above 240°C is separated from the bottom of the tower and enters the catalytic cracking unit 300 through 205. The reaction time is 3s, the catalyst-oil mass ratio is 10, and the weight hourly space velocity is 10h -1 Under the following conditions, a catalytic cracking reaction is carried out. A light aromatics-rich fraction with a temperature of 57 to 150° C. is extracted from the upper part of the side distillation zone C (the 15th plate from the top down). A first recycled fraction with a temperature of 150 to 240° C. is extracted from the lower part (the 32nd plate from the top down). The first recycled fraction enters the first reaction zone of the first riser of the catalytic cracking unit 300 through the first recycled fraction pipeline 204. The reaction time is 1.2 s, the catalyst-oil mass ratio is 10, and the weight hourly space velocity is 10h -1 The catalytic cracking reaction was carried out under the following conditions. The reaction products were separated via product line 305 into a second distillation column 400. The separated first fraction oil and gas with a distillation temperature below 240°C was combined with the hydrogenation oil and gas line 103 via line 401 and then fed into a degassing separation unit. The second fraction oil with a distillation range of 240-350°C was returned to the inlet of the hydrogenation unit 101 via line 402 for recycling. The slurry oil with a distillation temperature above 350°C was discharged via line 403. The product distribution and distillation ranges of each fraction are shown in Tables 2 and 3.

[0336] Comparative Example 1

[0337] The difference from Example 1 is that the catalytic cracking reaction products enter the second distillation tower 400 for separation. The final boiling point of the separated first fraction oil and gas is 210°C, and the distillation range of the second fraction oil is 210-350°C. The rest is the same as Example 1. The product distribution and the distillation range of each fraction are shown in Tables 2 and 3.

[0338] Comparative Example 2

[0339] The difference from Example 1 is that the catalytic cracking reaction products enter the second distillation tower 400 for separation. The final boiling point of the separated first fraction oil and gas is 260°C, and the distillation range of the second fraction oil is 260-350°C. The rest is the same as Example 1. The product distribution and the distillation range of each fraction are shown in Tables 2 and 3.

[0340] Comparative Example 3

[0341] The difference from Example 1 is that the first recycled fraction and the second recycled fraction separated by the distillation tower 200 are both fed into the first reaction zone of the catalytic cracking unit 300. The reaction time is 3s, the catalyst-oil mass ratio is 10, and the weight hourly space velocity is 10h -1 The catalytic cracking reaction was carried out under the same conditions as in Example 1. The product distribution and the distillation range of each fraction are shown in Tables 2 and 3.

[0342] Comparative Example 4

[0343] The difference from Example 1 is that the operating conditions of the dividing wall distillation column are adjusted, the first recycle fraction is withdrawn from the 36th plate at the top of the side distillation zone, the gas-liquid mass ratio entering the side distillation zone is reduced, the distillation range of the light aromatics-rich fraction is 57-180°C, the distillation range of the first recycle fraction is 180-310°C, and the distillation range of the second recycle fraction is 310-350°C. The rest are the same as in Example 1. The product distribution and the distillation range of each fraction are shown in Tables 2 and 3.

[0344] Comparative Example 5

[0345] Unlike Example 1, tower 200 is composed of two conventional distillation towers, functioning as a dividing-wall distillation tower. Distillation tower A has 40 stages and a reflux ratio of 0.8. Gas and a light fraction are separated at the top, a light aromatic fraction is separated in the middle, and a heavy fraction is separated at the bottom and fed to distillation tower B. Distillation tower B has 20 stages and a reflux ratio of 1.8. A first recycled fraction is separated at the top, and a second recycled fraction is separated at the bottom. All other aspects are the same as in Example 1. The product distribution and distillation ranges of each fraction are shown in Tables 2 and 3.

[0346] Comparative Example 6

[0347] After the catalytic cracking circulating oil is treated with two-stage hydrogenation, the hydrogenation product is separated into circulating hydrogen through high and low fractions. The hydrogenated oil gas enters the first distillation tower through the hydrogenation oil and gas pipeline. The tower has 40 plates and a reflux ratio of 0.8. The gas and the light fraction at the top of the tower are separated. The light aromatics-rich fraction is separated in the tower. The hydrogenated diesel fraction is separated at the bottom of the tower. The hydrogenated diesel fraction is sent to the catalytic cracking unit. The reaction time is 3s, the catalyst-oil mass ratio is 10, and the weight hourly space velocity is 10h -1Catalytic cracking reaction was carried out under the following conditions. The reaction products were separated in a second distillation tower. A first fraction of oil and gas with a distillation range of <210°C was separated at the top of the tower. A second fraction of oil with a distillation range of 210-350°C was separated in the middle of the tower. Catalytic oil slurry with a distillation range of >350°C was separated at the bottom of the tower. The second fraction of oil with a distillation range of 210-350°C was returned to the hydrogenation unit for recycling. The first fraction of oil and gas was separated in a third distillation tower for light aromatics. The third distillation tower had 40 stages and a reflux ratio of 2. A light fraction was separated at the top of the tower, a light aromatic fraction was separated in the middle of the tower, and a heavy fraction was separated at the bottom of the tower. The product distribution and the distillation ranges of each fraction are shown in Tables 2 and 3.

[0348] Example 2

[0349] 2 , the hydrogenation unit 100 uses an upper and lower fixed-bed reactor connected in series, with FZC and FC-70 hydrogenation catalysts placed in the upper and lower sections respectively. Catalytic cracking cycle oil enters the hydrogenation unit through the catalytic cracking cycle oil inlet 101 and passes through the first and second sections in sequence.

[0350] A dividing wall E is installed within the distillation tower 200, separating the tower radially. The pre-separation zone A, common distillation section B, common stripping section D, and side distillation section C have 34, 6, 4, and 36 trays, respectively. The top reflux ratio is 6. The feed is located at the 17th tray from the top of the pre-separation zone. The light aromatics fraction B is withdrawn from the 15th tray from the top of the side distillation section C, and the first recycled fraction is withdrawn from the 34th tray from the top of the side distillation section C. The top pressure of the distillation tower 200 is 0.4 MPa, and the top temperature is 67°C. The side distillation section of the distillation tower is equipped with mid-stage heat extraction, with the side distillation withdrawal and return points located at trays 32 and 33 of the side distillation zone. The pre-separation zone of the distillation tower is equipped with a gas phase withdrawal point located at tray 25 of the pre-separation zone, and the return point is tray 33 of the side distillation zone. The liquid mass flow ratio from the common distillation section to the pre-separation zone and the side distillation zone was 2.64; the gas mass flow ratio from the common stripping section to the pre-separation zone and the side distillation zone was 4.22; and the total gas mass flow rate at the pre-separation zone withdrawal point to the withdrawal gas mass flow rate was 69.36. The catalytic cracking unit 300 employed a single riser reactor.

[0351] Catalytic cracking cycle oil with a distillation range of 240-367°C is passed through the catalytic cracking cycle oil inlet 101, and hydrogen is passed through the hydrogen inlet in sequence through the upper and lower sections of the hydrogenation unit and the two hydrogenation catalysts for contact reaction. The hydrogen partial pressure is 8.0 MPa, the upper and lower reaction temperatures are 360°C and 380°C respectively, and the volume space velocity is 2.0 h -1 and 1.0h -1 , the hydrogen to oil volume ratios are 800 and 1300 respectively.

[0352] The hydrogenated product enters the high-low fractionation device through pipeline 102 to separate the circulating hydrogen and low-fraction gas. The hydrogenated oil and gas enter the degassing separation device through pipeline 103. The separated distillate oil enters the next-door distillation tower 200 through pipeline 201 for separation, and four fractions with increasing boiling points are separated. A light fraction with a boiling point of less than 60°C is extracted from the top of the tower. A second recycled fraction with a boiling point of more than 280°C is separated from the bottom of the tower and returned to the second inlet of the hydrogenation unit 100 through 205. A light aromatics-rich fraction with a boiling point of 57-150°C is extracted from the upper part of the sideline distillation zone C (the 15th plate from top to bottom), and a first recycled fraction with a boiling point of 150-280°C is extracted from the lower part (the 34th plate from top to bottom). The first recycled fraction enters the first reaction zone of the first riser of the catalytic cracking unit 300 through the first recycled fraction pipeline 204. The reaction time is 2s, the catalyst-oil mass ratio is 10, and the weight hourly space velocity is 10h -1 The catalytic cracking reaction was carried out under the following conditions. The reaction products were separated via product line 305 into a second distillation column 400. The separated first fraction oil and gas with a distillation temperature below 240°C was combined with the hydrogenation oil and gas line 103 via line 401 and then fed into a degassing separation unit. The second fraction oil with a distillation range of 240-350°C was returned to the inlet of the hydrogenation unit 101 via line 402 for recycling. The slurry oil with a distillation temperature above 350°C was discharged via line 403. The product distribution is shown in Table 2.

[0353] Comparative Example 7

[0354] The difference from Example 2 is that the operating conditions of the dividing wall distillation column are adjusted. The first recycle fraction is withdrawn from the 36th plate at the top of the side distillation zone. The gas-liquid mass ratio entering the side distillation zone is reduced. The distillation range of the light aromatics-rich fraction is 57-180°C, the distillation range of the first recycle fraction is 180-310°C, and the distillation range of the second recycle fraction is 310-350°C. The rest are the same as in Example 2. The product distribution and the distillation range of each fraction are shown in Tables 2 and 3.

[0355] Table 2

[0356] The separation process accuracy and energy consumption of the embodiment and the comparative example were calculated, with embodiment 1 as the benchmark, and the comparison is shown in Table 3.

[0357] Table 3

[0358] The number of plates and reflux ratios of the distillation tower / dividing-wall distillation tower in Examples and Comparative Examples are shown in Table 4 below.

[0359] Table 4

[0360] Compared with Comparative Examples 1, 2, 3, and 4, the separation energy consumption of Example 1 is comparable, and the yield of light aromatics is 2.2, 3.0, 3.8, and 3.5 percentage points higher, respectively.

[0361] Compared with Comparative Example 5, when the separated raw materials and separated products are the same, the total energy consumption of separation using the dividing wall distillation tower in Example 1 is 30% lower.

[0362] Compared with Comparative Example 7, the separation energy consumption of Example 2 is equivalent, and the yield of light aromatics is increased by 3.37 percentage points.

[0363] Compared with Comparative Example 6, the separation energy consumption of Example 1 and Example 2 is 30% lower and the light aromatics yield is 14.95 and 13.98 percentage points higher, respectively.

[0364] The present application has been described above in conjunction with preferred embodiments, but these embodiments are merely exemplary and serve only as an illustrative example. On this basis, various replacements and improvements can be made to the present application, all of which fall within the scope of protection of the present application.

Claims

1. A method for producing light aromatics, characterized in that: The method comprises the following steps: (1) Catalytic cracking cycle oil is subjected to two-stage hydrogenation treatment in the presence of hydrogen and a hydrogenation catalyst to obtain a hydrogenated product; (2) hydrogen is separated from the hydrogenated product to obtain hydrogenated oil gas, which is passed through a first distillation tower to obtain a light fraction at the top of the tower, a fraction rich in light aromatics, a first recycled fraction, and a second recycled fraction. The first recycled fraction has an initial distillation point of 140 to 170° C. and a final distillation point of 220 to 300° C., and the second recycled fraction has a distillation range starting point temperature of any value between 220 and 300° C.; (3) The first recycled fraction obtained in step (2) enters the first reaction zone of the catalytic cracking unit for catalytic cracking reaction. Optionally, the second recycled fraction enters the second reaction zone of the catalytic cracking unit or returns to the second stage of hydrogenation for recycling; (4) The reaction product obtained from the catalytic cracking unit is separated into a first fraction oil gas, a second fraction oil and an oil slurry by a second distillation tower device, and the second fraction oil is returned to the first hydrogenation treatment stage. Optionally, the first fraction oil gas is separated from the hydrogenated oil gas by entering the first distillation tower device of step (2); (5) Separating light aromatics from the light aromatics-rich fraction.

2. The method according to claim 1, characterized in that In step (1), the hydrotreatment reaction conditions satisfy at least one of the following conditions: The reaction conditions of the first stage of the hydrotreatment are: hydrogen partial pressure 5.0-20.0 MPa, reaction temperature 300-450°C, volume space velocity 0.5-5.0 h -1 , Hydrogen to oil volume ratio 300~1600Nm 3 / m 3 and / or The reaction conditions of the second stage of hydrotreatment are hydrogen partial pressure of 5.0-20.0 MPa, reaction temperature of 350-500°C, preferably 10-50°C higher than the temperature of the first stage of hydrotreatment, volume space velocity of 0.1-3.0 h -1 , Hydrogen to oil volume ratio 400~2000Nm 3 / m 3 .

3. The method according to claim 1 or 2, characterized in that In step (1), in the two-stage hydroprocessing reaction, the hydrogenation catalysts are the same or different; Preferably, the hydrogenation catalyst of the first stage of the hydroprocessing comprises an active metal component and a support, and the hydrogenation catalyst of the second stage of the hydroprocessing comprises an active metal, a molecular sieve and a support, wherein the active metal component is selected from one or more of Group VIB metals and / or Group VIII non-noble metals, the support is selected from one or more of alumina, silica, and amorphous silica-alumina, and the molecular sieve is selected from one or more of Y-type molecular sieve and USY molecular sieve; More preferably, the active metal component is selected from any one of the following metal combinations: nickel-tungsten, nickel-tungsten-cobalt, nickel-molybdenum, and cobalt-molybdenum.

4. The method according to any one of claims 1 to 3, characterized in that At least one of the following conditions is met, The catalytic cracking cycle oil has an initial boiling point of 160-250°C and a final boiling point of 340-400°C; and / or The initial distillation point of the hydrogenated oil gas is 0-60°C, and the final distillation point is 280-360°C.

5. The method according to claim 1, wherein In step (2), the first distillation tower device is a next-wall distillation tower, and at least one dividing wall is set in the next-wall distillation tower. The dividing wall divides the next-wall distillation tower into four areas: a pre-separation area, a common distillation section, a common stripping section and a side distillation area. The feed position of the hydrogenated oil and gas is located in the pre-separation area.

6. The method according to claim 5, characterized in that In step (2), The number of plates in the pre-separation zone of the dividing wall distillation tower is 10 to 40, preferably 20 to 35. The number of plates in the common distillation section is 2 to 10, preferably 3 to 8; The number of plates in the common stripping section is 2 to 10, preferably 3 to 8; The number of trays in the side distillation zone is 10 to 45, preferably 25 to 40; And the dividing wall distillation column satisfies at least one of the following conditions: The feed position of the hydrogenated oil and gas is above the 4th to 22nd plates at the top of the pre-separation zone, preferably above the 8th to 20th plates; The top pressure of the dividing wall distillation tower is 0.1-1.0 MPa, preferably 0.2-0.7 MPa, and the top temperature is 50-120° C., preferably 60-100° C.; and / or The top reflux ratio is 3 to 10, preferably 4 to 8.

7. The method according to claim 6, characterized in that In step (2), the dividing wall distillation tower satisfies at least one of the following conditions: The side distillation zone of the distillation tower is provided with a mid-section circulation heat extraction, the extraction position is the 20th to 45th tray of the side distillation zone, more preferably the 30th to 40th tray, and the return position is the same as the extraction position or is one tray above or below the extraction position; The pre-separation zone of the dividing wall distillation tower is provided with a gas phase extraction, the extraction position is located at the 20th to 40th plate of the pre-separation zone, more preferably the 25th to 35th plate, and the gas phase return position is the 30th to 45th plate of the side line distillation zone, more preferably the 35th to 40th plate; The mass flow ratio of the liquid phase from the common distillation section into the pre-separation zone and the side distillation zone is 0.1 to 10, preferably 0.5 to 5; The mass flow ratio of the gas phase from the common stripping section into the pre-separation zone and the side-line rectification zone is 1 to 15, preferably 2 to 10; and / or The ratio of the total gas phase mass flow rate at the extraction position of the pre-separation zone to the extracted gas phase mass flow rate is 10-150, preferably 30-120.

8. The method according to claim 6, characterized in that In step (2), the dividing wall distillation tower satisfies at least one of the following conditions: The light fraction at the top of the tower is separated from the common distillation section of the distillation tower, with an initial distillation point of 20-40°C and a final distillation point of 55-65°C; The light aromatic hydrocarbon-rich fraction is extracted from the 5th to 30th plates of the side distillation zone of the dividing wall distillation tower, preferably the 10th to 20th plates, with an initial distillation point of 55 to 65° C. and a final distillation point of 140 to 170° C. The first recycled fraction is produced from any position below the production position of the light aromatics-rich fraction in the side distillation zone, preferably the 25th to 35th tray; and / or The second recycled fraction is separated from the lower part of the common stripping section of the dividing wall distillation tower.

9. The method according to any one of claims 1 to 8, characterized in that In step (2), the initial distillation point of the first recycled fraction is 140-170°C, the final distillation point is 220-250°C, and the distillation range starting point temperature of the second recycled fraction is any value between 220-250°C; or In step (2), the initial distillation point of the first recycled fraction is 140-170°C and the final distillation point is 250-300°C, and the distillation range starting point temperature of the second recycled fraction is any value between 250-300°C.

10. The method according to any one of claims 1 to 8, characterized in that When the distillation range starting point temperature of the second recycled fraction is any value between 220 and 250° C., in step (3), the second recycled fraction enters the second reaction zone of the catalytic cracking unit; When the distillation range starting temperature of the second recycled fraction is any value between 270°C and 300°C, in step (3), the second recycled fraction is returned to the second stage of the hydrotreatment for recycling; When the distillation range starting temperature of the second recycled fraction is any value between 250° C. and 270° C., in step (3), the second recycled fraction enters the second reaction zone of the catalytic cracking unit or returns to the second stage of the hydrotreatment for recycling.

11. The method according to any one of claims 1 to 10, characterized in that In step (3), the conditions for catalytic conversion in the first reaction zone meet at least one of the following conditions: The weight hourly space velocity of oil and gas is 4 to 20h -1 , preferably 6 to 12 hours -1 ; The reaction time is 0.5 to 8 s, preferably 1 to 5 s; and / or The agent-oil weight ratio is 6 to 26, preferably 8 to 12; Furthermore, the conditions for catalytic conversion in the second reaction zone satisfy at least one of the following conditions: The weight hourly space velocity of oil and gas is 4 to 20h -1 , preferably 8 to 12 hours -1 ; The reaction time is 1 to 10 seconds, preferably 2 to 6 seconds; and / or The weight ratio of agent to oil is 4 to 14, preferably 6 to 10.

12. The method according to any one of claims 1 to 11, characterized in that The second distillation tower device in step (4) satisfies at least one of the following conditions: The top pressure of the distillation tower device is 0.15-0.4 MPa, preferably 0.18-0.35 MPa, and the top temperature is 100-160°C, preferably 115-145°C. The first fraction oil and gas has an initial boiling point of 0-40°C and a final boiling point of 220-250°C, and the first fraction oil and gas is degassed after being combined with the hydrogenated oil and gas and then enters the first distillation tower device of step (2) for separation; The second distillate oil has an initial boiling point of 220-250° C. and a final boiling point of 340-360° C.; and / or The initial boiling point of the oil slurry is 340-360°C.

13. A system for producing light aromatics, comprising: A hydrogenation unit, wherein the hydrogenation unit is a two-stage hydrogenation unit, provided with a hydrogen inlet and a hydrogenation product outlet, a catalytic cracking cycle oil inlet provided on the first stage of the hydrogenation unit, and an optional second recycled fraction inlet provided on the second stage of the hydrogenation unit; A catalytic cracking unit comprising a catalytic cracking reactor, wherein the catalytic cracking reactor is provided with a first reaction zone and an optional second reaction zone, wherein the first reaction zone is provided with a first recycled fraction inlet, and the optional second reaction zone is provided with a second recycled fraction inlet; a second distillation tower device, wherein the product distillation tower device is provided with a catalytic cracking reaction product inlet, a first distillate oil and gas outlet, a second distillate oil outlet and an oil slurry outlet, wherein the second distillate oil outlet is connected to the catalytic cracking cycle oil inlet of the hydrogenation unit; a degassing and separation device, wherein the degassing and separation device is provided with an oil and gas inlet, a rich gas outlet, and a distillate oil outlet, the oil and gas inlet of the degassing and separation device being connected to the first distillate oil and gas outlet of the second rectifying tower device and the hydrogenated product outlet of the hydrogenation device, so that the first distillate oil and gas from the second rectifying tower device and the hydrogenated oil and gas from the hydrogenation device are degassed and separated into rich gas and distillate oil; a first distillation tower device, the distillation tower device being provided with a distillate inlet, an outlet for a light fraction at the top of the tower, an outlet for a fraction rich in light aromatics, a first recycled fraction outlet, and a second recycled fraction outlet, the distillate inlet being connected to the distillate outlet of the degassing separation device, the first recycled fraction outlet being connected to the first recycled fraction inlet of the catalytic cracking device, the second recycled fraction outlet being connected to a second recycled fraction inlet provided in an optional second reaction zone of the catalytic cracking device, or the second recycled fraction outlet being connected to a second recycled fraction inlet optionally provided in the second section of the hydrogenation device; The separation device is provided with an inlet for a light aromatic hydrocarbon-rich fraction and an outlet for a light aromatic hydrocarbon-rich fraction, and the inlet for the light aromatic hydrocarbon-rich fraction is connected to the outlet for the light aromatic hydrocarbon-rich fraction of the first distillation tower device.

14. The system according to claim 13, wherein: The first distillation tower device is a dividing wall distillation tower, and at least one dividing wall is provided in the dividing wall distillation tower, and the dividing wall divides the dividing wall distillation tower into four areas: a common stripping section provided at the bottom of the tower, a common distillation section provided at the top of the tower, a pre-separation zone provided in the middle of the tower, and a side distillation zone; the pre-separation zone is provided with a distillate oil inlet, and the distillate oil inlet is connected to the distillate oil outlet of the degassing separation device; the common distillation section is provided with a light distillate outlet at the top of the tower, the side distillation zone is provided with a light aromatics-rich distillate outlet and a first recycle distillate outlet, and the bottom of the common stripping section is provided with a second recycle distillate outlet. At this time, the first recycle fraction outlet of the side distillation zone of the distillation tower is connected to the first recycle fraction inlet of the catalytic cracking reactor. The second recycle fraction outlet at the bottom of the common stripping section of the dividing wall distillation tower is connected to the second recycle fraction inlet optionally provided on the second reaction zone of the first distillation tower device, or is connected to the second recycle fraction inlet optionally provided on the second section of the hydrogenation device.

15. The system according to claim 13 or 14, characterized in that The system further includes a high-low separation device, which is provided with an inlet for hydrogenated products, a hydrogen outlet, and a hydrogenated oil and gas outlet, and the high-low separation device separates the hydrogenated oil and gas from the hydrogenated products from the hydrogenation device; At this time, the hydrogenation product inlet is connected to the hydrogenation product outlet of the hydrogenation device, and the oil and gas inlet of the degassing separation device is connected to the first fraction oil and gas outlet of the second distillation tower device and the hydrogenation oil and gas outlet of the high and low fraction device.

16. The system according to any one of claims 13 to 15, characterized in that The first reaction zone and the second reaction zone of the catalytic cracking reactor are arranged in the same or different reactors; The reactor is selected from one or more of the following: a constant diameter riser, a constant linear speed riser, a variable diameter riser, a variable linear speed riser, a fluidized bed, and a composite reactor. The composite reactor is composed of a constant diameter or variable diameter riser and a fluidized bed. Preferably, the reactor is a riser reactor. More preferably, the first reaction zone and the second reaction zone of the catalytic cracking reactor are arranged on a double-riser reactor, wherein the main riser is the first reaction zone and the secondary riser is the second reaction zone.

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