Catalytic conversion method for producing light aromatic hydrocarbon

Abstract

The invention provides a catalytic conversion method and system for producing light aromatic hydrocarbon, catalytic cracking circulation oil is subjected to two-stage hydrotreating reaction in the presence of hydrogen and a hydrogenation catalyst, and the reaction product is separated by a bulkhead fractionating tower to obtain gas, hydrogenated light naphtha, hydrogenated heavy naphtha, hydrogenated light diesel oil and hydrogenated heavy diesel oil; wherein the hydrogenated heavy diesel oil returns to the second section of the hydrogenation device, the hydrogenated light diesel oil enters a catalytic cracking device and is subjected to a cracking reaction in the presence of a catalytic cracking catalyst, and reaction products are separated to obtain dry gas, liquefied gas, catalytic naphtha rich in light aromatic hydrocarbon, a catalytic heavy aromatic hydrocarbon fraction, catalytic cracking circulating oil and oil slurry; and the catalytic cracking circulating oil returns to the first section of the hydrogenation device for recycling. The method can significantly improve the yield of light aromatic hydrocarbons.

Description

Technical Field

[0001] The present invention relates to the field of petrochemical industry, and in particular to a catalytic conversion method 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] US4585545 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] ZL201310010219.5 discloses a method for producing aromatic compounds, wherein the diesel fraction with a distillation range of 250-450°C obtained by catalytic cracking is sent to a hydrotreating unit, where the catalytic cracking heavy cycle oil is hydrogenated and then catalytically cracked to produce benzene, toluene and xylene.

[0007] ZL201310517666.X discloses a catalytic conversion method for producing aromatics-rich gasoline, which cuts catalytic cracking light cycle oil into light and heavy fractions. The heavy fraction is hydrogenated and then fed separately to 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 catalytic conversion method for producing light aromatics based on the existing technology.

[0010] To achieve the above-mentioned object, the present invention provides a catalytic conversion method for producing light aromatics, comprising: subjecting catalytic cracking cycle oil to a two-stage hydrogenation reaction in the presence of hydrogen and a hydrogenation catalyst; and separating the reaction products through a distillation tower to obtain gas, hydrogenated light naphtha, hydrogenated heavy naphtha, hydrogenated light diesel oil, and hydrogenated heavy diesel oil; wherein the hydrogenated heavy diesel oil is returned to the second stage of the hydrogenation unit, and the hydrogenated light diesel oil is fed into a catalytic cracking unit to undergo a cracking reaction in the presence of a catalytic cracking catalyst; and the 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 oil slurry; and the catalytic cracking cycle oil is returned to the first stage of the hydrogenation unit for recycling.

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

[0012] (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;

[0013] (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;

[0014] (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;

[0015] (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.

[0016] According to the catalytic conversion method of the first aspect, wherein, in step (1), the hydrotreatment reaction conditions are:

[0017] 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

[0018] 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 .

[0019] According to the catalytic conversion method of the first aspect, wherein, in step (1), in the two-stage hydroprocessing reaction, the hydrogenation catalysts are the same or different;

[0020] 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;

[0021] 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.

[0022] According to the catalytic conversion method of the first aspect, in step (2), at least one partition wall is provided in the next-wall distillation tower, and the partition 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 hydrogenation oil and gas feed position is located in the pre-separation area.

[0023] According to the catalytic conversion method of the first aspect, wherein the top pressure of the dividing wall distillation tower is 0.1 to 1.0 MPa, preferably 0.2 to 0.7 MPa;

[0024] The tower top temperature is 50-120°C, preferably 60-100°C;

[0025] 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;

[0026] 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

[0027] 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.

[0028] According to the catalytic conversion method of the first aspect, the number of plates in the pre-separation zone in the dividing wall distillation tower is 10 to 40, preferably 20 to 35;

[0029] 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

[0030] The number of trays in the side distillation zone is 10 to 45, preferably 25 to 40.

[0031] According to the catalytic conversion method of the first aspect, wherein the side distillation zone is heavy, 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

[0032] Preferably, the pre-separation zone is provided with a gas phase extraction, the extraction position is located at the 20th to 40th plates of the pre-separation zone, preferably the 25th to 35th plates, and the return tower position is the 30th to 45th plates of the side line distillation zone, preferably the 30th to 40th plates.

[0033] According to the catalytic conversion method of the first aspect, wherein, in step (2), the hydrogenated light naphtha is separated from the common distillation section of the dividing wall distillation tower, with an initial distillation point of 20 to 40° C. and a final distillation point of 55 to 65° C.;

[0034] 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;

[0035] 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

[0036] The hydrogenated heavy diesel oil is separated from the lower part of the common stripping section of the distillation tower, and the initial distillation point is 230-280°C.

[0037] According to the catalytic conversion method of the first aspect, wherein, in step (3), the initial boiling point of the catalytic naphtha is 20-40°C and the final boiling point is 140-160°C;

[0038] 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

[0039] The initial boiling point of the catalytic cracking cycle oil is 210-250°C, and the final boiling point is 320-360°C.

[0040] According to the catalytic conversion method of the first aspect, wherein, in step (3), the conditions of the catalytic cracking reaction include:

[0041] The weight hourly space velocity of oil and gas is 4 to 20h -1 , preferably 8 to 16 hours -1 ;

[0042] The catalytic cracking reaction temperature is 520-650°C, preferably 560-640°C; and / or

[0043] The weight ratio of agent to oil is 6 to 26, preferably 8 to 20.

[0044] According to the catalytic conversion method of the first aspect, wherein, 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;

[0045] 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;

[0046] The binder is selected from silicon oxide binder and / or aluminum oxide binder;

[0047] The carrier is selected from one or more of silicon dioxide, kaolin, montmorillonite, diatomaceous earth, halloysite, saponite, rectorite, sepiolite, attapulgite, hydrotalcite and bentonite.

[0048] A second aspect of the present invention provides a catalytic conversion system for producing light aromatics, comprising:

[0049] 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;

[0050] 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;

[0051] A catalytic cracking unit is 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 circulating oil outlet and an oil slurry outlet. The catalytic heavy aromatics outlet is connected to a catalytic heavy aromatics fractionation pipeline, and the catalytic cracking circulating oil outlet is connected to a catalytic cracking circulating oil inlet on the first section of the hydrogenation unit. Optionally, the catalytic heavy aromatics fractionation pipeline is connected to the hydrogenated light diesel inlet.

[0052] According to the catalytic conversion system of the second aspect, wherein the catalytic cracking unit is a single reactor or a dual reactor;

[0053] 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.

[0054] Through the above technical solution, the present invention utilizes a bulkhead separation tower to efficiently separate hydrogenated light diesel oil with a high yield of light aromatics from hydrogenated catalytic cracking cycle oil as a catalytic cracking feedstock, thereby increasing the yield of light aromatics.

[0055] Other features and advantages of the present invention will be described in detail in the following detailed description. BRIEF DESCRIPTION OF THE DRAWINGS

[0056] 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:

[0057] Figure 1 This is a schematic diagram of the system structure in a preferred embodiment of the present invention.

[0058] Figure 2 Schematic diagram of the comparative example system structure of the present invention.

[0059] Description of Reference Numerals

[0060] 100. Hydrogenation unit (including hot and cold high and low separation sections); 200. Dividing wall distillation tower; 300. Catalytic cracking unit; 400. Light naphtha separation tower; 500. Heavy naphtha separation tower;

[0061] 101. Hydrogen inlet; 102. Catalytic cracking cycle oil inlet; 103. Hydrogenation oil and gas pipeline; 104. Hydrogenation unit first section; 105. Hydrogenation unit second section;

[0062] 201. Pre-separation zone; 202. Public distillation section; 203. Public stripping section; 204. Side distillation zone; 205. Dividing wall; 206. Hydrogenated light diesel pipeline; 207. Hydrogenated heavy diesel pipeline;

[0063] 301. Catalytic heavy aromatics fraction pipeline; 302. Catalytic cracking circulating oil pipeline;

[0064] 501. Hydrogenated diesel pipeline;

[0065] A01, A01′. Gas; A02, A02′. Hydrogenated light naphtha; B, B′. Hydrogenated heavy naphtha; C, C′. Dry gas; D, D′. Liquefied gas; E, E′. Catalytic naphtha; F, F′. Catalytic slurry oil, G. Catalytic heavy aromatics. DETAILED DESCRIPTION

[0066] 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.

[0067] 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.

[0068] 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.

[0069] 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.

[0070] 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.

[0071] 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.

[0072] 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.

[0073] Before describing the technical solution of the present invention, the terms used herein are defined as follows:

[0074] The term "initial distillation point" refers to the temperature at the starting point of the distillation range.

[0075] The term "end point" refers to the temperature at the end of the distillation range.

[0076] The light aromatic hydrocarbons mentioned in the present invention refer to C6 to C8 aromatic hydrocarbons.

[0077] The present invention provides a catalytic conversion method for producing light aromatics, the method comprising:

[0078] (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;

[0079] (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;

[0080] (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;

[0081] (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.

[0082] The inventors of the present invention have found that light diesel produced by hydrogenation of catalytic cracking cycle oil is more suitable as a catalytic cracking feedstock for producing light aromatics than full-fraction diesel, and can significantly increase the yield of light aromatics.

[0083] Optionally, the heavy aromatic hydrocarbon fraction of the catalytic cracking product is returned to the catalytic cracking unit for zone conversion to further increase the yield of light aromatic hydrocarbons.

[0084] In one embodiment, in step (1), the hydrotreating reaction conditions are:

[0085] 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

[0086] 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 .

[0087] In one embodiment, in step (1), in the two-stage hydroprocessing reaction, the hydrogenation catalysts are the same or different;

[0088] 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;

[0089] 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.

[0090] In one embodiment, in step (2), at least one partition wall is provided in the distillation tower, and the partition wall divides the distillation tower into four areas: a pre-separation area, a common distillation section, a common stripping section and a side distillation area, and the hydrogenated oil and gas feed position is located in the pre-separation area.

[0091] In one embodiment, in the dividing wall distillation tower, the top pressure is 0.1 to 1.0 MPa, preferably 0.2 to 0.7 MPa;

[0092] The tower top temperature is 50-120°C, preferably 60-100°C;

[0093] 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;

[0094] 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

[0095] 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.

[0096] In one embodiment, the number of trays in the pre-separation zone in the dividing wall distillation tower is 10 to 40, preferably 20 to 35;

[0097] 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

[0098] The number of trays in the side distillation zone is 10 to 45, preferably 25 to 40.

[0099] In one embodiment, the side distillation zone is heavy, and the extraction and return positions are the 20th to 45th trays, preferably the 30th to 40th trays, of the side distillation zone. Preferably, the side distillation zone is further provided with a mid-stage circulation heat extraction; and / or

[0100] Preferably, the pre-separation zone is provided with a gas phase extraction, the extraction position is located at the 20th to 40th plates of the pre-separation zone, preferably the 25th to 35th plates, and the return tower position is the 30th to 45th plates of the side line distillation zone, preferably the 30th to 40th plates.

[0101] In one embodiment, in step (2), the light naphtha is separated from the common distillation section of the dividing wall distillation tower, with an initial distillation point of 20 to 40° C. and a final distillation point of 55 to 65° C.;

[0102] The heavy naphtha is produced from the 5th to 30th plates, preferably the 10th to 20th plates, of the side distillation zone of the dividing wall distillation tower, with an initial distillation point of 55 to 65° C. and a final distillation point of 140 to 170° C.

[0103] 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

[0104] The hydrogenated heavy diesel oil is separated from the lower part of the common stripping section of the distillation tower, and the initial distillation point is 230-280°C.

[0105] In one embodiment, in step (3), the initial boiling point of the catalytic naphtha is 20-40°C and the final boiling point is 140-160°C;

[0106] 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

[0107] The initial boiling point of the catalytic cracking cycle oil is 210-250°C, and the final boiling point is 320-360°C.

[0108] In one embodiment, in step (3), the conditions of the catalytic cracking reaction include:

[0109] The weight hourly space velocity of oil and gas is 4 to 20h -1 , preferably 8 to 16 hours -1 ;

[0110] The catalytic cracking reaction temperature is 520-650°C, preferably 560-640°C; and / or

[0111] The weight ratio of agent to oil is 6 to 26, preferably 8 to 20.

[0112] The present invention adopts a dividing wall distillation tower to separate hydrogenated light diesel oil from hydrogenated heavy diesel oil. The circulation of hydrogenated heavy diesel oil can further improve the yield of light aromatic hydrocarbons from diesel hydrogenation. Hydrogenated light diesel oil is the best raw material for catalytic cracking to produce light aromatic hydrocarbons, and has a similar distillation range to the heavy aromatic hydrocarbon fraction in the catalytic product. The hydrogenated light diesel oil and the recycled catalytic heavy aromatic hydrocarbons are converted together as catalytic cracking raw materials to improve the utilization efficiency of aromatic hydrocarbons. The method realizes the maximum production of light aromatic hydrocarbons through hydrogenation and catalytic cracking.

[0113] In one embodiment, in step (3), the catalytic cracking catalyst contains 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;

[0114] 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;

[0115] The binder is selected from silicon oxide binder and / or aluminum oxide binder;

[0116] The carrier is selected from one or more of silicon dioxide, kaolin, montmorillonite, diatomaceous earth, halloysite, saponite, rectorite, sepiolite, attapulgite, hydrotalcite and bentonite.

[0117] The present invention also provides a catalytic conversion system for producing light aromatics, comprising:

[0118] 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;

[0119] 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;

[0120] A catalytic cracking unit is 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 circulating oil outlet and an oil slurry outlet. The catalytic heavy aromatics outlet is connected to a catalytic heavy aromatics fractionation pipeline, and the catalytic cracking circulating oil outlet is connected to a catalytic cracking circulating oil inlet on the first section of the hydrogenation unit. Optionally, the catalytic heavy aromatics fractionation pipeline is connected to the hydrogenated light diesel inlet.

[0121] In one embodiment, the catalytic cracking unit is a single reactor or a dual reactor;

[0122] 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.

[0123] refer to Figure 1As a preferred embodiment of the present invention, the catalytic cracking circulating oil with a boiling range of 220 to 350°C enters the first section 104 of the hydrogenation unit through the catalytic cracking circulating oil inlet 102, and contacts with the hydrogen and hydrogenation catalyst entering through the hydrogen inlet 101 to perform a two-stage hydrogenation reaction. After the reaction oil and gas are separated into circulating hydrogen and low-fraction gas through hot and cold high and low separation, the liquid phase enters the pre-separation zone 201 of the next-door distillation tower 200 through the hydrogenation oil and gas pipeline 103. The pre-separated oil and gas enter the side line distillation zone 204 through the pre-separation zone oil and gas pipeline 208. The top public distillation section 202 separates the fraction below 60°C, and then separates the gas A01 and the hydrogenated light naphtha fraction A02. The bottom public stripping section 203 separates the hydrogenated heavy diesel oil with a temperature above 240°C. Pipeline 207 returns to the second section 105 of the hydrogenation unit. Hydrogenated heavy naphtha B with a distillation range of 60-160°C is separated from the upper portion of the sideline fractionation zone 204. Hydrogenated light diesel with a distillation range of 160-240°C is separated from the lower portion and enters the catalytic cracking unit 300 via the hydrogenated light diesel pipeline 206 as the feedstock for the first reaction zone, where it contacts and reacts with the catalytic cracking catalyst. The reaction products are dry gas C, liquefied gas D, catalytic naphtha E rich in benzene, toluene, and xylene, and slurry oil F. The catalytic heavy aromatics fraction with a distillation range of 150-240°C is returned to the catalytic cracking unit 300 via the catalytic heavy aromatics fraction pipeline 301 for zone-by-zone conversion. The catalytic cracking circulating oil with a distillation range of 240-350°C is returned to the first section 104 of the hydrogenation unit via the catalytic cracking circulating oil pipeline 302 for recycling. The catalytic cracking unit utilizes a single riser reactor.

[0124] 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.

[0125] The hydrogenation catalysts used in the examples and comparative examples were purchased from brands FZC and FC-70, and the catalytic cracking catalyst was purchased from brands SLA-1.

[0126] The properties of the catalytic cracking cycle oil used in the Examples and Comparative Examples are shown in Table 1.

[0127] Table 1

[0128] Raw oil name Catalytic cracking cycle oil <![CDATA[Density (20 °C), kg / m 3 > 934.3 Hydrocarbon mass composition, weight % 100 Alkanes 13.2 Cycloalkanes 5.7 Aromatics 81.1 Monocyclic aromatic hydrocarbons 31.9 Bicyclic aromatic hydrocarbons 46.1 Other aromatics 3.1 Distillation range, ℃ 223～350

[0129] Example 1

[0130] refer to Figure 1The hydrogenation unit 100 utilizes a fixed-bed reactor and comprises hot and cold high-low sections, including a first hydrogenation unit section 104 and a second hydrogenation unit section 105. FZC and FC-70 hydrogenation catalysts are placed in the upper and lower sections, respectively. Hydrogen enters the first hydrogenation unit section 104 through a hydrogen inlet 101, and catalytic cracking cycle oil enters the first hydrogenation unit section 104 through a catalytic cracking cycle oil inlet 102. A dividing wall 205 is installed within the bulkhead distillation tower 200, radially separating the tower. The number of stages in the pre-separation zone 201, common distillation section 202, common stripping section 203, and side distillation zone 204 are 34, 6, 4, and 36, respectively. The top reflux ratio is 6. Hydrogenated heavy naphtha B is produced from the 15th stage from the top of the side distillation zone 204, and hydrogenated light diesel is produced from the 32nd stage from the top of the side distillation zone 204. The top pressure of the bulkhead distillation tower 200 is 0.4 MPa, and the top temperature is 67°C. The side-stream distillation section is equipped with a mid-stage heat extraction system, with the side-stream withdrawal and return points located at trays 31 and 30 of the side-stream distillation section. The pre-separation section of the bulkhead distillation tower is equipped with a gas phase withdrawal point located at tray 25 of the pre-separation section, and the return point is tray 30 of the side-stream distillation section. The liquid mass flow ratio from the common distillation section to the pre-separation section and the side-stream distillation section is 2.64; the gas mass flow ratio from the common stripping section to the pre-separation section and the side-stream distillation section is 4.22; and the total gas mass flow rate at the pre-separation section withdrawal point to the withdrawal gas mass flow ratio is 69.36. The catalytic cracking unit 300 utilizes a single riser reactor.

[0131] Catalytic cracking cycle oil with a distillation range of 220-350℃ and hydrogen are sequentially passed through the upper and lower sections of the hydrogenation unit and contacted with two hydrogenation catalysts for reaction. The hydrogen partial pressure is 8.0MPa, the upper and lower reaction temperatures are 360℃ and 380℃ respectively, and the volume space velocity is 2.0h -1 and 1.0h -1 The volume ratio of hydrogen to oil is 800 and 1300 respectively. The reaction oil and gas enter the distillation tower 200 through the hydrogenation oil and gas pipeline 103 for separation, and three fractions with low to high boiling points are separated. The light fraction with a boiling point of less than 60°C is extracted from the top of the tower, which can be further divided into A01 gas and A02 hydrogenation light naphtha fraction (25-60°C). The common stripping section 203 at the bottom of the tower separates hydrogenation heavy diesel with a boiling point greater than 240°C and passes through the hydrogenation heavy diesel pipeline. 207 returns to the second section 105 of the hydrogenation unit, and the upper part of the side distillation zone 204 (the 15th plate from the top) extracts the hydrogenated heavy naphtha B at 57-150°C, and the lower part (the 32nd plate from the top) extracts the hydrogenated light diesel at 150-240°C. The hydrogenated light diesel enters the catalytic cracking unit 300 through the hydrogenated light diesel pipeline 206. The reaction temperature is 580°C, the catalyst-oil mass ratio is 10, and the weight hourly space velocity is 10h -1The reaction occurs under the following conditions: contact with the catalyst. After the reaction, the oil and gas are separated to yield dry gas C, liquefied gas D, catalytic naphtha E rich in benzene, toluene, and xylene with a distillation range of 25-150°C, catalytic heavy aromatics with a distillation range of 150-240°C, catalytic cracking cycle oil and slurry oil F with a distillation range of 240-350°C. The catalytic heavy aromatics fraction with a distillation range of 150-240°C is returned to the catalytic cracking unit 300 for recycling and conversion via the heavy aromatics fraction pipeline 301, and the catalytic cracking cycle oil with a distillation range of 240-350°C is returned to the hydrogenation unit via the catalytic cracking cycle oil pipeline 302. The product distribution is shown in Table 2.

[0132] Example 2

[0133] The hydrogenation unit 100 utilizes a fixed-bed reactor, comprising a first section 104 and a second section 105. Two hydrogenation catalysts, RN-411 and RHC-100, are placed in the upper and lower sections, respectively. Hydrogen enters the first section 104 through a hydrogen inlet 101, and catalytic cracking cycle oil enters the first section 104 through a catalytic cracking cycle oil inlet 102. A dividing wall 205 is installed within the bulkhead distillation tower 200, radially separating the tower. The pre-separation zone 201, common distillation section 202, common stripping section 203, and side distillation zone 204 have 34, 6, 4, and 36 stages, respectively. The top reflux ratio is 6. Hydrogenated heavy naphtha B is produced from the 15th stage from the top of the side distillation zone 201, and hydrogenated light diesel is produced from the 32nd stage at the top of the side distillation zone 204. The top pressure of the bulkhead distillation tower 200 is 0.4 MPa, and the top temperature is 67°C. The side-stream distillation section is equipped with a mid-stage heat extraction system, with the side-stream withdrawal and return points located at trays 31 and 30 of the side-stream distillation section. The pre-separation section of the bulkhead distillation tower is equipped with a gas phase withdrawal point located at tray 25 of the pre-separation section, and the return point is tray 30 of the side-stream distillation section. The liquid mass flow ratio from the common distillation section to the pre-separation section and the side-stream distillation section is 2.64; the gas mass flow ratio from the common stripping section to the pre-separation section and the side-stream distillation section is 4.22; and the total gas mass flow rate to the withdrawal gas mass flow ratio at the pre-separation section withdrawal point is 69.36. The catalytic cracking unit 300 utilizes a single riser reactor, with the catalyst passing through the second reaction zone and the first reaction zone in sequence from bottom to top.

[0134] Catalytic cracking cycle oil with a distillation range of 220-350℃ and hydrogen are sequentially passed through the upper and lower sections of the hydrogenation unit 100 and contacted with two hydrogenation catalysts for reaction. The hydrogen partial pressure is 8.0MPa, the upper and lower reaction temperatures are 360℃ and 380℃ respectively, and the volume space velocity is 2.0h -1 and 1.0h -1The volume ratios of hydrogen and oil are 800 and 1300 respectively. The reaction oil and gas enter the distillation tower 200 next door for separation through the hydrogenation oil and gas pipeline 103, and three fractions with boiling points from low to high are separated. The light fraction with a boiling point of less than 60°C is extracted from the top of the tower, which can be further divided into A01 gas and A02 hydrogenated light naphtha fraction (25-60°C). The common stripping section 203 at the bottom of the tower separates hydrogenated heavy diesel with a temperature greater than 240°C and returns to the second section 105 of the hydrogenation unit. The upper part of the side distillation zone 204 (the 15th plate from top to bottom) extracts hydrogenated heavy naphtha B with a temperature of 57-150°C, and the lower part (the 32nd plate from top to bottom) extracts hydrogenated light diesel with a temperature of 150-240°C. The hydrogenated light diesel enters the catalytic cracking unit 300 through the hydrogenation light diesel pipeline 206. When the reaction temperature is 580°C, the catalyst-oil mass ratio is 10, and the weight hourly space velocity is 10h -1 Under these conditions, the reaction occurs with the catalyst. After the reaction, the oil and gas are separated to produce dry gas C, liquefied gas D, catalytic naphtha E with a distillation range of 25-150°C, catalytic heavy aromatics with a distillation range of 150-240°C, catalytic cracking cycle oil with a distillation range of 240-350°C, and slurry oil F. The catalytic cracking cycle oil with a distillation range of 240-350°C is returned to the hydrogenation unit 100 via catalytic cracking cycle oil pipeline 302, while the catalytic heavy aromatics fraction is not returned to the catalytic cracking unit. The product distribution is shown in Table 2.

[0135] Comparative Example

[0136] like Figure 2 As shown, this comparative example differs from the example in that after the catalytic cracking circulating oil is hydrogenated in the hydrogenation unit 100, the reaction oil and gas are separated into circulating hydrogen and low-fraction gas through cold and hot high and low fractionation. The liquid phase then enters the light naphtha separation tower 400, which has 30 stages and a reflux ratio of 0.8. Gas A01 and hydrogenated light naphtha A02 fractions are separated at the top of the tower. Hydrogenated heavy naphtha and hydrogenated diesel fractions are separated at the bottom of the tower and fed to the heavy naphtha separation tower 500, which has 40 stages and a reflux ratio of 1.8. Hydrogenated heavy naphtha fraction B is produced from the top of the tower, and the hydrogenated diesel fraction at the bottom of the tower enters the catalytic cracking unit 300 via the hydrogenation full-range diesel pipeline 501. The operating conditions of the catalytic cracking unit 300 are the same as those of Example 2. The product distribution is shown in Table 2.

[0137] Table 2

[0138]

[0139] As can be seen from Table 2, the total C6-C8 light aromatics yields of Example 1 and Example 2 are 11.68 and 3.96 percentage points higher than that of the comparative example, respectively.

[0140] Based on Aspen, the separation process accuracy and energy consumption of the embodiment and comparative example were calculated, and the comparison is shown in Table 3.

[0141] Table 3

[0142]

[0143]

[0144] Compared with the comparative example, the separation accuracy of the embodiment for heavy naphtha is comparable under the working condition of producing an additional stream of hydrogenated light diesel oil, and the total energy consumption remains basically unchanged, while saving the equipment investment of a separation tower.

[0145] 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 catalytic conversion method for producing light aromatic hydrocarbons, characterized in that: The method comprises: (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; (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; (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; (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.

2. The catalytic conversion method according to claim 1, characterized in that: In step (1), the hydrotreatment reaction conditions are: 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 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 .

3. The catalytic conversion method according to claim 1, characterized in that: In step (1), in the two-stage hydroprocessing reaction, the hydrogenation catalysts are the same or different; 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; 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 catalytic conversion method according to claim 1, characterized in that: In step (2), at least one partition wall is provided in the distillation tower, and the partition wall divides the distillation tower into four areas: a pre-separation area, a common distillation section, a common stripping section and a side distillation area. The hydrogenated oil and gas feed position is located in the pre-separation area.

5. The method according to claim 4, characterized in that In the dividing wall distillation tower, the top pressure is 0.1-1.0 MPa, preferably 0.2-0.7 MPa; The tower top temperature is 50-120°C, preferably 60-100°C; 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 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.

6. The catalytic conversion method according to claim 4, characterized in that: 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 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 The number of trays in the side distillation zone is 10 to 45, preferably 25 to 40.

7. The catalytic conversion method according to claim 6, characterized in that: The side distillation zone is heavy, and the extraction and return positions are the 20th to 45th trays, preferably the 30th to 40th trays, of the side distillation zone. Preferably, the side distillation zone is further provided with a mid-stage circulation heat extraction; and / or Preferably, the pre-separation zone is provided with a gas phase extraction, the extraction position is located at the 20th to 40th plates of the pre-separation zone, preferably the 25th to 35th plates, and the return tower position is the 30th to 45th plates of the side line distillation zone, preferably the 30th to 40th plates.

8. The catalytic conversion method according to any one of claims 4 to 7, characterized in that: In step (2), the hydrogenated light naphtha 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 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; 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 The hydrogenated heavy diesel oil is separated from the lower part of the common stripping section of the distillation tower, and the initial distillation point is 230-280°C.

9. The catalytic conversion method according to claim 1, characterized in that: In step (3), the initial boiling point of the catalytic naphtha is 20-40° C., and the final boiling point is 140-160° C.; 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 The initial boiling point of the catalytic cracking cycle oil is 210-250°C, and the final boiling point is 320-360°C.

10. The catalytic conversion method according to claim 1, characterized in that: In step (3), the conditions for the catalytic cracking reaction include: The weight hourly space velocity of oil and gas is 4 to 20h -1 , preferably 8 to 16 hours -1 ; The catalytic cracking reaction temperature is 520-650°C, preferably 560-640°C; and / or The weight ratio of agent to oil is 6 to 26, preferably 8 to 20.

11. The catalytic conversion method according to claim 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; 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; The binder is selected from silicon oxide binder and / or aluminum oxide binder; The carrier is selected from one or more of silicon dioxide, kaolin, montmorillonite, diatomaceous earth, halloysite, saponite, rectorite, sepiolite, attapulgite, hydrotalcite and bentonite.

12. A catalytic conversion 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 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; 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; A catalytic cracking unit is 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 circulating oil outlet and an oil slurry outlet. The catalytic heavy aromatics outlet is connected to a catalytic heavy aromatics fractionation pipeline, and the catalytic cracking circulating oil outlet is connected to a catalytic cracking circulating oil inlet on the first section of the hydrogenation unit. Optionally, the catalytic heavy aromatics fractionation pipeline is connected to the hydrogenated light diesel inlet.

13. The catalytic conversion system according to claim 12, characterized in that: The catalytic cracking unit is a single reactor or a dual reactor; 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.

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