A processing method for combined production of chemical raw materials by coking and hydrocracking

Abstract

The present invention discloses a method for combined production of chemical raw materials by coking and hydrocracking. The method includes: the coking raw material enters the delayed coking reaction zone for thermal cracking reaction to obtain petroleum coke and gaseous effluent; the gaseous effluent is fractionated to obtain coker dry gas, coker light distillate oil and coker wax oil; wherein, the mass content of tricyclic aromatic hydrocarbons in the coker light distillate oil is not higher than 1.0%; the coker light distillate oil is mixed with hydrogen and undergoes a normal paraffin conversion reaction in the normal paraffin conversion reaction zone, and the mass content of C7+ normal paraffins in the reaction effluent is controlled to be 0.1% - 5.0%; in the presence of hydrogen, the normal paraffin conversion reaction effluent enters the hydrocracking reaction zone for hydrocracking reaction, and then the hydrocracking reaction effluent is separated and fractionated to obtain gas fraction, light naphtha, heavy naphtha and tail oil. This method uses coker light distillate oil as the raw material for hydrocracking to produce chemical raw materials, which can significantly improve the quality and yield of chemical raw materials.

Description

Technical Field

[0001] The present invention belongs to the field of heavy oil processing, and particularly relates to a method for combined production of chemical raw materials by delayed coking and hydrocracking, especially a processing method for producing high-quality chemical raw materials from vacuum residue as the raw material. Background Art

[0002] Reducing the production of heavy fuel oil is the development trend of the world's refining industry today. Although the current capacity of catalytic cracking for single refining and blending of residue has accounted for more than 25% of the total capacity of catalytic cracking, not all residues can be processed by catalytic cracking. If the carbon residue content of the residue exceeds 10% and the metal content exceeds 100 - 150 ppm, the residue hydrotreating / catalytic cracking combined unit will also have difficulty bearing the increasingly high catalyst cost and the increasingly long shutdown time. Coupled with the growing demand for light oil products, the widening price difference between light crude oil and heavy high-sulfur crude oil, and the increasing proportion of heavy sulfur-containing / high-sulfur crude oil supply, especially since delayed coking can process cheap heavy high-sulfur high-metal residues, and coker gasoline can be used as a raw material for ethylene production by hydrogenation after cracking, delayed coking has thus become the most popular technology for residue processing and the preferred residue processing scheme for many refineries.

[0003] CN201580070326.4 discloses a method for preparing LPG and BTX, comprising: a) subjecting a mixed hydrocarbon stream to a first hydrocracking in the presence of a first hydrocracking catalyst to prepare a first hydrocracking product stream; b) separating the first hydrocracking product stream to provide at least one light hydrocarbon stream containing at least C2 and C3 hydrocarbons, an intermediate hydrocarbon stream composed of C4 and / or C5 hydrocarbons, and a heavy hydrocarbon stream containing at least C6+ hydrocarbons, and c) subjecting the heavy hydrocarbon stream to a second hydrocracking in the presence of a second hydrocracking catalyst to prepare a second hydrocracking product stream containing BTX, wherein the second hydrocracking is more severe than the first hydrocracking, d) wherein, in the presence of a C4 hydrocracking catalyst, at least part of the intermediate hydrocarbon stream is subjected to C4 hydrocracking to prepare a C4 hydrocracking product stream, and the C4 hydrocracking is optimized for converting C4 hydrocarbons into C3 hydrocarbons.

[0004] CN201480037272.7 discloses a method for producing light olefin hydrocarbon compounds from a hydrocarbon raw material, comprising the following steps: (a) feeding the hydrocarbon raw material to a reaction zone for ring opening; (b) separating the reaction product generated from the reaction zone into an overhead stream and a side stream; (c) feeding the side stream from (b) to a gasoline hydrocracker (GHC) unit; (d) separating the reaction product of the GHC in step (c) into an overhead stream containing hydrogen, methane, ethane, and liquefied petroleum gas and a stream containing aromatic hydrocarbon compounds and a small amount of hydrogen and non-aromatic hydrocarbon compounds; (e) feeding the overhead stream from the gasoline hydrocracker (GHC) unit to a steam cracker unit.

[0005] In summary, petroleum hydrocarbons have a complex composition, mainly including alkanes, cycloalkanes, and aromatics. Among them, alkanes, especially small-molecule alkanes, are high-quality ethylene raw materials, while cycloalkanes and aromatics are high-quality reforming feeds. Delayed coking causes thermal cracking of heavy raw materials, following the free radical reaction mechanism. The products have a high content of normal hydrocarbons and cyclic hydrocarbons. In the prior art, it is impossible to efficiently achieve the conversion effect of "ethylene for ethylene and aromatics for aromatics" for coker distillate oil. Therefore, it is of great significance to develop a processing method suitable for producing high-quality chemical raw materials from heavy oils such as vacuum residue. Summary of the Invention

[0006] Aiming at the problems existing in the prior art, the object of the present invention is to provide a method for combined production of chemical raw materials by coking and hydrocracking. This method uses heavy oils such as vacuum residue as raw materials, which can significantly improve the quality and yield of chemical raw materials.

[0007] The present invention provides a method for combined production of chemical raw materials by coking and hydrocracking, and the method includes:

[0008] (1) The coking raw material enters the delayed coking reaction zone for thermal cracking reaction to obtain petroleum coke and a gaseous effluent; the gaseous effluent is fractionated to obtain coker dry gas, coker light distillate oil, and coker wax oil; wherein, the mass content of tricyclic aromatics in the coker light distillate oil is not higher than 1.0%.

[0009] (2) The coker light distillate oil obtained in step (1) is mixed with hydrogen and passes through the normal hydrocarbon conversion reaction zone for normal hydrocarbon conversion reaction, and the mass content of C7 + normal paraffins in the reaction effluent is 0.1% - 5.0%.

[0010] (3) In the presence of hydrogen, the normal hydrocarbon conversion reaction effluent from step (2) enters the hydrocracking reaction zone for hydrocracking reaction, and then the hydrocracking reaction effluent is separated and fractionated to obtain a gas fraction, light naphtha, heavy naphtha, and tail oil.

[0011] According to the present invention, the coking raw material oil in step (1) is a heavy oil with an initial boiling point greater than 350°C, and can be selected from one or more of atmospheric residue, vacuum residue, visbreaking residue, heavy deasphalted oil, catalytic cracking slurry, heavy oil, topped crude oil, shale oil, and coal liquefied oil, preferably vacuum residue. The initial boiling point of the vacuum residue is 420°C - 620°C, preferably 450°C - 550°C; the sulfur mass content is 2% - 10%, preferably 4% - 8%; the nitrogen mass content is 0.2% - 1.0%, preferably 0.3% - 0.5%; the metal content is 100 mg / kg - 500 mg / kg, preferably 200 mg / kg - 300 mg / kg.

[0012] According to the present invention, in step (1), the temperature of the delayed coking reaction zone is 480 to 520 °C, preferably 490 to 505 °C; the operating pressure is 0.1 MPa to 2.0 MPa, preferably 0.2 MPa to 0.5 MPa, and it can be a constant pressure operation or a variable pressure operation; the recycle ratio is 0.1 to 0.5, preferably 0.2 to 0.4, and the recycle ratio is the mass ratio of the coker gas oil fraction to the feedstock. The initial boiling point of the coker gas oil is 280 °C to 340 °C, preferably 290 °C to 320 °C; the final boiling point is 450 °C to 580 °C, preferably 480 °C to 520 °C.

[0013] According to the present invention, in step (1), the initial boiling point of the light coker distillate oil is 50 °C to 80 °C, preferably 60 °C to 70 °C; the final boiling point is 280 °C to 340 °C, preferably 290 °C to 320 °C.

[0014] According to the present invention, preferably, in step (1), the mass content of tricyclic aromatic hydrocarbons in the light coker distillate oil is 0.2% to 0.6%.

[0015] According to the present invention, preferably, in step (2), the mass content of C7 + n-alkanes in the reaction effluent is 1.0% to 3.0%.

[0016] According to the present invention, the chemical raw materials mainly include ethane, propane, butane, light naphtha, and may also include heavy naphtha. Among them, heavy naphtha is used as a reforming raw material to produce BTX, and ethane, propane, butane, and light naphtha are used as raw materials for producing lower olefins. For example, they are used as steam cracking raw materials to produce ethylene, and propane and butane can also be directly dehydrogenated to produce propylene and butene. Among them, lower olefins refer to olefins with four or fewer carbon atoms, especially ethylene, propylene, and butadiene.

[0017] According to the present invention, in step (2), the n-paraffin conversion reaction zone is filled with catalysts; specifically, a hydrofining catalyst and a first hydrocracking catalyst are filled in sequence along the material flow direction. The first hydrocracking catalyst can be one or more catalysts. The hydrofining catalyst can be one or more catalysts. The filling ratio of the hydrofining catalyst to the first hydrocracking catalyst is 0.5 to 5:1 by volume.

[0018] According to the present invention, the hydrofining catalyst described in step (2) can adopt a conventional hydrofining catalyst, which is mainly used for hydrodesulfurization, denitrification and other impurities. The hydrofining catalyst includes a carrier and a hydrogenation active metal. The carrier is an inorganic refractory oxide, generally selected from one or more of alumina, amorphous silica-alumina, silica or titanium oxide, etc.; the hydrogenation active metal includes Group VIB and / or Group VIII metal components. Further preferably, in the hydrofining catalyst, Group VIB is preferably selected from tungsten and / or molybdenum, and its content in the catalyst is 5 wt% to 30 wt% based on the mass of the oxide, preferably 10 wt% to 20 wt%; Group VIII is preferably selected from nickel and / or cobalt, and its content in the catalyst is 1 wt% to 6 wt% based on the mass of the oxide, preferably 1.5 wt% to 5 wt%. The content of the carrier in the catalyst is 64 wt% to 94 wt%, preferably 75 wt% to 88.5 wt%.

[0019] According to the present invention, in step (2), the first hydrocracking catalyst includes an active metal component and a carrier; the carrier includes a molecular sieve having a selective cracking effect on normal paraffins, preferably selected from one or more of ZSM-5 molecular sieve, ZSM-11, ZSM-12, ZSM-22, ZSM-23, ZSM-35 and ZSM-38 molecular sieves, preferably ZSM-5 molecular sieve. The SiO2 / Al2O3 molar ratio of the ZSM-5 molecular sieve is 20 to 60. The carrier may also include a binder. Preferably, the binder is alumina. The active metal component includes at least one of metals in Group VIB and Group VIII. The Group VIB metal is preferably molybdenum and / or tungsten, and the Group VIII metal is preferably cobalt and / or nickel.

[0020] According to the present invention, in step (2), preferably, for the first hydrocracking catalyst, based on the weight of the catalyst, the content of the Group VIB metal (calculated as the oxide) is 5.0% to 15.0%, the content of the Group VIII metal (calculated as the oxide) is 2.0% to 5.0%, and the content of the carrier is 80.0% to 93.0%.

[0021] According to the present invention, in step (2), preferably, in the carrier of the first hydrocracking catalyst, based on the weight of the carrier, the content of the binder is 8% to 60%, and the content of the molecular sieve is 40% to 92%.

[0022] According to the present invention, in step (2), preferably, the specific surface area of the first hydrocracking catalyst is 200 to 400 m 2 2 / g, and the pore volume is 0.25 to 0.45 mL / g.

[0023] According to the present invention, the preparation method of the first hydrocracking catalyst in step (2) can be prepared according to the conventional methods in the art. The preparation method includes the preparation of a support and the loading of active metal components. The process of preparing the support is as follows: The shape-selective cracking molecular sieve and the binder are mechanically mixed, formed, and then dried and calcined to make the catalyst support. The drying and calcination of the support can adopt conventional conditions. The conditions for drying are: drying at 100°C to 150°C for 1 to 12 hours. The conditions for calcination are: calcining at 450°C to 550°C for 2.5 to 6.0 hours.

[0024] According to the present invention, in step (2), in the preparation method of the first hydrocracking catalyst, the method for loading active metal components is a conventional method, such as kneading method, impregnation method, etc., and the impregnation method is preferred. The impregnation method can be saturated impregnation method, excess impregnation method or complex impregnation method, that is, the catalyst support is impregnated with a solution containing the required active components, and then dried and calcined to obtain the first hydrocracking catalyst. The conditions for drying are: drying at 100°C to 150°C for 1 to 12 hours. The conditions for calcination are: calcining at 450°C to 550°C for 2.5 to 6.0 hours.

[0025] According to the present invention, in step (2), the nitrogen content in the reaction effluent is below 50 mg / kg, and more preferably below 20 mg / kg.

[0026] According to the present invention, in step (2), the reaction conditions in the n-paraffin conversion reaction zone are as follows: the reaction pressure is 5 to 9 MPa, preferably 6 to 8 MPa.

[0027] According to the present invention, in step (2), the reaction conditions in the n-paraffin conversion reaction zone are as follows: the average reaction temperature is 250 to 450°C, preferably 300 to 400°C; the liquid hourly space velocity is 0.1 to 15.0 h -1 , preferably 1.0 to 5.0 h -1 ; the hydrogen-oil volume ratio is 100:1 to 2500:1, preferably 400:1 to 2000:1.

[0028] According to the present invention, in step (3), a second hydrocracking catalyst is loaded in the hydrocracking reaction zone. The second hydrocracking catalyst can be one or more catalysts.

[0029] According to the present invention, the hydrocracking reaction effluent obtained in step (3) is a hydrocracking product containing monocyclic cyclic hydrocarbons.

[0030] According to the present invention, in the hydrocracking reaction effluent of step (3), the mass ratio of C6-C8 monocyclic cyclic hydrocarbons to the total cyclic hydrocarbon mass in the coker light distillate oil feedstock is 0.35 to 0.55, preferably 0.41 to 0.50.

[0031] According to the present invention, the tail oil obtained after the hydrocracking reaction effluent in step (3) is subjected to distillation separation can be recycled to the hydrocracking reaction zone in step (3).

[0032] According to the present invention, in step (3), the second hydrocracking catalyst has the functions of ring opening and cracking of polycyclic cyclic hydrocarbons, selectively cracking the side chains of isoparaffins or cyclic hydrocarbons and retaining monocyclic cyclic hydrocarbons. The second hydrocracking catalyst comprises a cracking component, a hydrogenation component and a binder. The second hydrocracking catalyst can be a commercially available product or prepared according to the prior art. The hydrogenation component is at least one of a metal, a metal oxide, and a metal sulfide of an active metal component; the active metal component comprises a metal of Group VIB and / or Group VIII; the active metal component is more preferably at least one of iron, chromium, molybdenum, tungsten, cobalt, and nickel. In the second hydrocracking catalyst, the binder is alumina and / or silica; the cracking component comprises an acidic molecular sieve, preferably at least one of Beta zeolite and Y zeolite.

[0033] According to the present invention, preferably, in step (3), the hydrocracking reaction zone is filled with a catalyst with Y zeolite as the cracking component and a catalyst with Beta zeolite as the cracking component in sequence along the material flow direction; preferably, the volume ratio of the catalyst with Y zeolite as the cracking component to the catalyst with Beta zeolite as the cracking component is 1:1 to 1:5, preferably 1:2 to 1:4.

[0034] According to the present invention, in the second hydrocracking catalyst in step (3), based on the weight of the second hydrocracking catalyst, the content of the hydrogenation component in terms of oxide is 5 wt% to 40 wt%, preferably 10 wt% to 30 wt%; the content of the cracking component is 10 wt% to 80 wt%, preferably 20 wt% to 60 wt%; the content of the binder is 5 wt% to 85 wt%, preferably 10 wt% to 50 wt%.

[0035] According to the present invention, the preparation method of the second hydrocracking catalyst in step (3) can be prepared according to the conventional methods in the art. The preparation method includes the preparation of a carrier and the loading of a hydrogenation component, wherein the preparation process of the carrier is as follows: the cracking component and the binder are mechanically mixed, formed, and then dried and calcined to form a catalyst carrier. The drying and calcination of the carrier can adopt conventional conditions. The conditions for drying are: drying at 100 °C to 150 °C for 1 to 12 hours. The conditions for calcination are: calcining at 450 °C to 550 °C for 2.5 to 6.0 hours.

[0036] According to the present invention, in step (3), in the preparation method of the second hydrocracking catalyst, the method for loading the hydrogenation component is a conventional method, such as kneading method, impregnation method, etc., and the impregnation method is preferred. The impregnation method can be saturated impregnation method, excess impregnation method or complex impregnation method, that is, impregnating the catalyst carrier with a solution containing the required hydrogenation component, and then drying and calcining to obtain the second hydrocracking catalyst. The conditions for the drying are: drying at 100°C to 150°C for 1 to 12 hours. The conditions for the calcining are: calcining at 450°C to 550°C for 2.5 to 6.0 hours.

[0037] According to the present invention, in step (3), the reaction conditions in the hydrocracking reaction zone are as follows: the reaction pressure is 5 to 9 MPa, preferably 6 to 8 MPa.

[0038] According to the present invention, in step (3), the reaction conditions in the hydrocracking reaction zone are as follows: the average reaction temperature is 250 to 450°C, preferably 300 to 400°C; the liquid hourly space velocity is 0.1 to 15.0 h -1 , preferably 1.0 to 5.0 h -1 ; the hydrogen-oil volume ratio is 100:1 to 2500:1, preferably 400:1 to 2000:1.

[0039] According to the present invention, preferably, the normal paraffin conversion reaction zone and the hydrocracking reaction zone adopt the same pressure.

[0040] According to the present invention, preferably, the effluent from the hydrocracking reaction in step (3) is subjected to supplementary hydrofining. The supplementary hydrofining can be to load a hydrofining catalyst at the bottom of the hydrocracking reaction zone, or it can enter a separate hydrofining reaction zone.

[0041] According to the present invention, preferably, the effluent from the hydrocracking reaction in step (3) can also first enter the fractionation system, and the separated heavy naphtha component is subjected to supplementary hydrofining.

[0042] According to the present invention, the tail oil obtained in step (3) can be recycled to the hydrocracking reaction zone.

[0043] Petroleum hydrocarbon compositions are complex and mainly contain paraffins, naphthenes and aromatics, while high-quality ethylene feedstocks are small-molecule normal paraffins, and reforming feedstocks are monocyclic naphthenes and aromatics. The inventors have found through research that the technical solution of the present invention can highly selectively generate small-molecule normal paraffins, thereby realizing the efficient enrichment of small-molecule normal paraffins in ethylene feedstocks, and at the same time retaining monocyclic hydrocarbons in heavy naphtha as much as possible to achieve the efficient enrichment of high-quality reforming feedstocks, so as to achieve the purpose of greatly improving the yields of chemical feedstocks (i.e., ethylene feedstocks and reforming feedstocks) and the quality of ethylene feedstocks and reforming feedstocks, and thus complete the present invention.

[0044] Compared with the prior art, the present invention has the following beneficial technical effects:

[0045] (1) In the prior art, when hydrocracking coker diesel as a raw material to produce chemical raw materials, the final boiling point of the coker diesel fractionated from the coker distillate is generally 350 - 380 °C. This also results in a relatively high content of tricyclic aromatic hydrocarbons in the coker diesel fraction, and a relatively high reaction pressure is required to achieve the hydrocracking of tricyclic aromatic hydrocarbons. However, a relatively high reaction pressure will also cause some monocyclic aromatic hydrocarbons to open-ring crack during the hydrocracking process, resulting in the loss of aromatic hydrocarbons. After the light coker distillate oil of the present invention enters the fractionating tower, by controlling the content of tricyclic aromatic hydrocarbons in the top oil of the fractionating tower to be no higher than 1%, and selecting an appropriate reaction pressure, the mixed processing of monocyclic aromatic hydrocarbons and bicyclic aromatic hydrocarbons is realized, reducing the loss of aromatic hydrocarbons in the raw material during the hydrogenation process and increasing the aromatic hydrocarbon content in the hydrogenation product.

[0046] (2) The light coker distillate oil first enters the n-paraffin conversion reaction zone and contacts the shape-selective cracking catalyst to carry out the chain-breaking reaction of straight-chain hydrocarbons. By controlling the n-paraffin content in the effluent of the shape-selective cracking reaction within a certain range, it is ensured that the straight-chain hydrocarbons are enriched in the light naphtha and gas products after the cracking reaction. Then, through hydrocracking, the polycyclic cyclic hydrocarbons are ring-opened and cracked while retaining the monocyclic cyclic hydrocarbons, and further selective chain-breaking of the isomeric hydrocarbons is carried out to achieve efficient separation of the paraffins and cyclic hydrocarbons in the coker naphtha. + The n-paraffin content in the effluent of the shape-selective cracking reaction is within a certain range, ensuring that the straight-chain hydrocarbons are enriched in the light naphtha and gas products after the cracking reaction. Then, through hydrocracking, the polycyclic cyclic hydrocarbons are ring-opened and cracked while retaining the monocyclic cyclic hydrocarbons, and further selective chain-breaking of the isomeric hydrocarbons is carried out to achieve efficient separation of the paraffins and cyclic hydrocarbons in the coker naphtha.

[0047] (3) The cyclic hydrocarbons in the heavy naphtha of the hydrocracking product are highly enriched and the paraffin content is low. Using this as the catalytic reforming raw material, the catalytic reforming unit only needs to adopt two reaction zones: six-membered ring dehydrogenation and five-membered ring isomerization dehydrogenation, canceling the paraffin cyclization reaction zone, and significantly reducing the energy consumption of the catalytic reforming unit. BRIEF DESCRIPTION OF THE DRAWINGS

[0048] Figure 1 is a schematic process flow diagram of Embodiments 1 - 4 of the present invention;

[0049] Main reference numerals description:

[0050] 1 - vacuum residue, 2 - delayed coking reaction zone, 3 - petroleum coke, 4 - delayed coking gas effluent, 5 - fractionating tower, 6 - coking gas fraction, 7 - light coker distillate oil, 8 - coker wax oil, 9 - hydrogen, 10 - n-paraffin conversion reaction zone, 11 - n-paraffin conversion reaction effluent, 12 - hydrocracking reaction zone, 13 - hydrocracking reaction effluent, 14 - separator, 15 - gas-phase stream rich in hydrogen gas, 16 - liquid-phase stream, 17 - fractionating tower, 18 - gas fraction, 19 - light naphtha, 20 - heavy naphtha, 21 - tail oil. DETAILED DESCRIPTION OF THE INVENTION

[0051] The functions and effects of the present invention will be further illustrated by the following examples, but the following examples do not limit the method of the present invention.

[0052] In the present invention, % is the mass fraction unless otherwise specified.

[0053] The overall volumetric space velocity in the examples and comparative examples is the ratio of the fresh feed volume to the total volume of the catalyst.

[0054] The method of the present invention, as Figure 1 shown, includes: vacuum residue 1 enters the delayed coking reaction zone 2 to obtain petroleum coke 3 and the delayed coking gas effluent 4. The delayed coking gas effluent 4 enters the fractionating tower 5 to be separated into a coking gas fraction 6, a coking light distillate oil 7, and a coking wax oil 8. The coking light distillate oil 8 is mixed with hydrogen 9 and enters the n-paraffin conversion reaction zone 10 for n-paraffin conversion reaction. The n-paraffin conversion reaction effluent 11 enters the hydrocracking reaction zone 12 for hydrocracking reaction. The hydrocracking reaction effluent 13 enters the separator 14. The separated gas-phase stream, the hydrogen-rich gas 15, is recycled for use. The liquid-phase stream 16 enters the fractionating tower 17 to be fractionated into a gas fraction 18, light naphtha 19, heavy naphtha 20, and tail oil 21. The tail oil 21 is recycled to the inlet of the hydrocracking reaction zone 12.

[0055] In the present invention, the first hydrocracking catalyst in each example is represented by Cat-A plus a number, such as Cat-A1, Cat-A2, Cat-A3, Cat-A4. The first hydrocracking catalyst is prepared by a conventional active metal saturation impregnation method, and the physical and chemical properties of the obtained catalyst are shown in Table 1.

[0056] In the present invention, the second hydrocracking catalyst in each example is represented by Cat-B plus a number, such as Cat-B1, Cat-B2. The physical and chemical properties of the catalyst are shown in Table 2. The second hydrocracking catalyst in each example is prepared by a conventional active metal saturation impregnation method. Among them, the properties of the Beta zeolite used in Cat-B1 are as follows: the SiO2 / Al2O3 molar ratio is 30, the specific surface area is 350 m 2 / g, and the pore volume is 0.32 cm 3 / g. The properties of the Y zeolite used in Cat-B2 are as follows: the SiO2 / Al2O3 molar ratio is 15, the specific surface area is 400 m 2 / g, and the pore volume is 0.30 cm 3 / g.

[0057] In the present invention, in Examples 1 to 3, Cat-B2 and Cat-B1 are successively filled in the hydrocracking reaction zone along the material flow direction.

[0058] In the present invention, the hydrofining catalyst in each example is represented by Cat-J. The properties of the hydrofining catalyst are shown in Table 3.

[0059] In the present invention, the feedstock oil in each example is vacuum residue, and its main properties are shown in Table 4.

[0060] In the present invention, the loading ratio of the hydrofining catalyst and the first hydrocracking catalyst in each example is 1:2.

[0061] In the present invention, the nitrogen content in the reaction effluent in step (2) in each example is below 20 mg / kg.

[0062] In the present invention, the ethylene feedstock in each example refers to ethane, propane, butane and light naphtha obtained in step (3), and ethane, propane, butane and light naphtha can be directly used as the feedstock for steam cracking to produce ethylene.

[0063] In the present invention, the distillation range of light naphtha is the liquid component with a boiling range less than 60 °C, the distillation range of heavy naphtha is 60 - 175 °C, and the distillation range of the tail oil is the component with a boiling range > 175 °C.

[0064] In the present invention, the yield of the ethylene feedstock refers to the mass ratio of ethane, propane, butane and light naphtha in the hydrocracking product to the fresh hydrocracking feedstock (coker light distillate oil), and the yield of heavy naphtha refers to the mass ratio of heavy naphtha in the hydrocracking product to the fresh hydrocracking feedstock (coker light distillate oil).

[0065] Examples 1 - 4

[0066] The processing method for combined production of chemical raw materials by coking and hydrocracking adopts the process as Figure 1 shown in the figure. The method specifically includes:

[0067] (1) The vacuum residue enters the delayed coking reaction zone for thermal cracking reaction to obtain petroleum coke and a gas-phase effluent; the gas-phase effluent enters the fractionating tower to obtain coker dry gas, coker light distillate oil and coker wax oil;

[0068] (2) The coker light distillate oil obtained in step (1) is mixed with hydrogen and undergoes a normal paraffin conversion reaction in the normal paraffin conversion reaction zone, and the mass content of C7 + normal paraffins in the reaction effluent is controlled; the normal paraffin conversion reaction zone is filled with a catalyst; along the direction of the material flow, the catalyst is filled with a hydrofining catalyst and a first hydrocracking catalyst in sequence.

[0069] (3) The reaction effluent from the normal paraffin conversion reaction in step (1) is separated into a gas-phase material flow and a liquid-phase material flow by gas-liquid separation. The gas-phase material flow is recycled, and the liquid-phase material flow enters the fractionating tower, where it is fractionated to obtain a gas fraction, light naphtha, heavy naphtha and tail oil. The tail oil is recycled to the inlet of the hydrocracking reaction zone.

[0070] The process conditions and hydrotreating effects of each example are shown in Table 6.

[0071] Comparative Example 1

[0072] The difference from Example 1 is that the coker light distillate oil directly enters the hydrocracking reaction zone after hydrofining.

[0073] The process conditions and hydrofining effect in this example are shown in Table 6.

[0074] Comparative Example 2

[0075] The difference from Example 1 is that in step (1), the content of C7 normal paraffins in the effluent of the normal paraffin conversion reaction is controlled to be 6%. + The content of normal paraffins is 6%.

[0076] The process conditions and hydrofining effect in this example are shown in Table 6.

[0077] Comparative Example 3

[0078] The difference from Example 1 is that the catalyst loading order in the hydrocracking reaction zone is different from that in Example 1. In this example, the order of catalysts Cat-B2 and Cat-B1 is exchanged. Specifically, in this example, catalysts Cat-B1 and catalyst Cat-B2 are loaded successively along the material flow direction in the hydrocracking reaction zone.

[0079] The process conditions and hydrofining effect in this example are shown in Table 6.

[0080] Comparative Example 4

[0081] The difference from Example 1 is that the content of tricyclic mass aromatics in the coker light distillate oil is 3.1%.

[0082] The process conditions and hydrofining effect in this example are shown in Table 6.

[0083] Table 1 Physicochemical properties of the first hydrocracking catalyst

[0084]

[0085]

[0086] Table 2 Physicochemical properties of the second hydrocracking catalyst

[0087] Catalyst properties Cat-B1 Cat-B2 <![CDATA[Pore volume, cm 3 / g]]> 0.35 0.35 <![CDATA[Specific surface area, m 2 / g]]> 300 300 Catalyst composition and content Beta, wt% 50 - Y, wt% - 30 <![CDATA[MoO3, wt%]]> 10 20 NiO, wt% 5 5 Aluminum oxide, wt% 35 45

[0088] Table 3 Hydrofining catalyst

[0089] Catalyst properties Cat-J <![CDATA[Pore volume, cm 3 / g]]> 0.35 <![CDATA[Specific surface area, m 2 / g]]> 200 Catalyst composition and content <![CDATA[MoO3, wt%]]> 22 NiO, wt% 4.5 Aluminum oxide, wt% 73.5

[0090] Table 4 Main properties of the feedstock oil

[0091] Name of feedstock oil Vacuum residue <![CDATA[Density (20 °C) / kg·m -3 > 1.032 Distillation range / ℃ IBP / 10% 464 / 558 30％ / 50％ 606 / 646 70％ / 90％ 749 / 973 95% / EBP 1025 / 1055 Sulfur content, wt% 5.61 Nitrogen content, wt% 0.38 Ni+V, mg / kg 200

[0092] Table 5 Delayed coking process conditions and main properties of the coker light distillate oil

[0093] Item Example 1 Example 2 Example 3 Example 4 Reaction temperature / ℃ 498 500 502 504 Reaction pressure / MPa 0.3 0.3 0.3 0.3 Recycle ratio 0.3 0.2 0.3 0.4 Distillation range of coker light distillate oil IBP 61 61 61 61 EBP 290 300 310 320 <![CDATA[C7 + Normal paraffin, wt%]]> 24.0 23.8 23.5 23.2 Cyclic hydrocarbons, wt% 58.0 58.6 59.7 61.4 Nitrogen content, mg / kg 1200 1260 1300 1350 Amount of tricyclic aromatic hydrocarbons in coker light distillate oil, % 0.2 0.3 0.4 0.6

[0094] Continued Table 5

[0095] Item Comparative Example 1 Comparative Example 2 Comparative Example 3 Comparative Example 4 Reaction temperature / ℃ 498 498 498 505 Reaction pressure / MPa 0.3 0.3 0.3 0.3 Recycle ratio 0.3 0.3 0.3 0.3 Distillation range of coker light distillate oil IBP 61 61 61 61 EBP 290 290 290 360 <![CDATA[C7 + n-alkane, wt%]]> 24.0 24.0 24.0 22.4 Cyclic hydrocarbons, wt% 58.0 58.0 58.0 65.0 Nitrogen content, mg / kg 1200 1200 1200 1500 Amount of tricyclic aromatic hydrocarbons in coker naphtha, % 0.2 0.2 0.2 3.1

[0096] Table 6 Process Conditions and Hydrogenation Effects of Each Example

[0097]

[0098]

[0099] Continued Table 6

[0100]

[0101] Note: Catalyst volume ratio *, in Comparative Examples 1 and 2, it is the volume ratio of Cat-B2 and Cat-B1; in Comparative Example 3, it is the volume ratio of Cat-B1 and Cat-B2.

[0102] The specific embodiments of the present invention have been described in detail above. However, the present invention is not limited thereto. Within the technical concept of the present invention, various simple modifications can be made to the technical solutions of the present invention, including any other suitable combination of each technical feature. These simple modifications and combinations should also be regarded as the content disclosed by the present invention and fall within the protection scope of the present invention.

Claims

1. A method for combined production of chemical raw materials by coking and hydrocracking, characterized in that, The method includes: (1) The coking feedstock enters the delayed coking reaction zone for thermal cracking reaction to obtain petroleum coke and gaseous effluent; the gaseous effluent is fractionated to obtain coker dry gas, coker light distillate oil and coker wax oil; wherein, the mass content of tricyclic aromatic hydrocarbons in the coker light distillate oil is not higher than 1.0%; (2) The coking light distillate oil obtained in step (1) is mixed with hydrogen and subjected to a normal paraffin conversion reaction in a normal paraffin conversion reaction zone, and the mass content of C7 + normal paraffins in the reaction effluent is 0.1% to 5.0%; (3) In the presence of hydrogen, the effluent from the n-paraffin conversion reaction in step (2) enters the hydrocracking reaction zone for hydrocracking reaction, and then the hydrocracking reaction effluent is separated and fractionated to obtain gas fraction, light naphtha, heavy naphtha and tail oil; In step (2), the n-paraffin conversion reaction zone is filled with catalysts, and a hydrofining catalyst and a first hydrocracking catalyst are filled in sequence along the material flow direction; The first hydrocracking catalyst includes an active metal component and a carrier; the carrier includes a molecular sieve with selective cracking of n-alkanes, and the molecular sieve is selected from one or more of ZSM-5 molecular sieve, ZSM-11, ZSM-12, ZSM-22, ZSM-23, ZSM-35 and ZSM-38 molecular sieves; the active metal component includes at least one of Group VIB metals and Group VIII metals; For the first hydrocracking catalyst, based on the weight of the catalyst, the content of Group VIB metal in terms of oxide is 5.0% - 15.0%, the content of Group VIII metal in terms of oxide is 2.0% - 5.0%, and the content of the carrier is 80.0% - 93.0%; in the carrier of the first hydrocracking catalyst, based on the weight of the carrier, the content of the binder is 8% - 60%, and the content of the molecular sieve is 40% - 92%; In step (3), the hydrocracking reaction zone is filled with a second hydrocracking catalyst; the second hydrocracking catalyst includes a cracking component, a hydrogenation component and a binder; in step (3), the hydrocracking reaction zone is filled with a catalyst with Y molecular sieve as the cracking component and a catalyst with Beta molecular sieve as the cracking component in sequence along the material flow direction; For the second hydrocracking catalyst, based on the weight of the second hydrocracking catalyst, the content of the hydrogenation component in terms of oxide is 5wt% - 40wt%; the content of the cracking component is 10wt% - 80wt%; the content of the binder is 5wt% - 85wt%; The chemical raw materials include ethane, propane, butane, light naphtha and heavy naphtha. Among them, heavy naphtha is used as a reforming raw material to produce BTX, and ethane, propane, butane and light naphtha are used as raw materials for producing light olefins.

2. The method according to claim 1, wherein The filling volume ratio of the hydrofining catalyst to the first hydrocracking catalyst is 0.5 to 5:1; and / or, the nitrogen content in the reaction effluent of step (2) is below 50 mg / kg; and / or, the mass content of C7 + n-alkanes is 1.0% to 3.0%.

3. The method according to claim 1, characterized in that, The nitrogen content in the reaction effluent of step (2) is below 20 mg / kg.

4. The method according to claim 1, wherein The hydrofining catalyst includes a carrier and a hydrogenation active metal; wherein: The carrier is an inorganic refractory oxide, and the carrier is selected from one or several of alumina, amorphous silica-alumina, silica and titanium oxide; and / or, the hydrogenation active metal includes Group VIB metal and / or Group VIII metal components.

5. The method according to claim 4, wherein In the hydrofining catalyst, the Group VIB metal is selected from tungsten and / or molybdenum, and its content in the catalyst based on the mass of the oxide is 5 wt% to 30 wt%; and / or, the Group VIII metal is selected from nickel and / or cobalt, and its content in the catalyst based on the mass of the oxide is 1 wt% to 6 wt%; and / or, the content of the carrier in the catalyst is 64 wt% to 94 wt%.

6. The method according to claim 4, wherein In the hydrofining catalyst, the Group VIB metal, based on the mass of the oxide, has a content in the catalyst of 10 wt% to 20 wt%; and / or, the Group VIII metal, based on the mass of the oxide, has a content in the catalyst of 1.5 wt% to 5 wt%; and / or, the content of the carrier in the catalyst is 75 wt% to 88.5 wt%.

7. The method according to claim 1, wherein In the first hydrocracking catalyst, the molecular sieve is ZSM-5 molecular sieve; The Group VIB metal is molybdenum and / or tungsten, and the Group VIII metal is cobalt and / or nickel.

8. The method according to claim 1, wherein In step (1), the coking feedstock is a heavy oil with an initial boiling point greater than 350 °C, and is selected from one or more of atmospheric residue, vacuum residue, visbreaking residue, heavy deasphalted oil, catalytic cracking slurry, heavy oil, topped crude oil, shale oil, and coal liquefied oil.

9. The method according to claim 8, wherein In step (1), the coking feedstock is vacuum residue; the initial boiling point of the vacuum residue is 420 °C to 620 °C; the sulfur mass content is 2% to 10%; the nitrogen mass content is 0.2% to 1.0%; the metal content is 100 mg / kg to 500 mg / kg.

10. The method according to claim 9, wherein The initial boiling point of the vacuum residue is 450 °C to 550 °C; the sulfur mass content is 4% to 8%; the nitrogen mass content is 0.3% to 0.5%; the metal content is 200 mg / kg to 300 mg / kg.

11. The method according to claim 1, wherein In step (1), the temperature of the delayed coking reaction zone is 480 to 520 °C; the operating pressure is 0.1 MPa to 2.0 MPa; the recycle ratio is 0.1 to 0.5, and the recycle ratio is the mass ratio of the coker gas oil fraction to the coking feedstock.

12. The method according to claim 1, wherein In step (1), the temperature of the delayed coking reaction zone is 490 to 505 °C; the operating pressure is 0.2 MPa to 0.5 MPa; the recycle ratio is 0.2 to 0.4, and the recycle ratio is the mass ratio of the coker gas oil fraction to the coking feedstock.

13. The method according to claim 1, wherein The initial boiling point of the coker light distillate oil is 50 °C to 80 °C; the final boiling point is 280 °C to 340 °C; and / or, the mass content of tricyclic aromatics in the coker light distillate oil is 0.2% to 0.6%.

14. The method according to claim 1, characterized in that, The initial boiling point of the coker light distillate oil is 60 °C to 70 °C; the final boiling point is 290 °C to 320 °C.

15. The method according to claim 1, wherein The reaction conditions in the normal paraffin conversion reaction zone in step (2) are as follows: the reaction pressure is 5 to 9 MPa; the average reaction temperature is 250 to 450 °C; and / or, the liquid hourly space velocity is 0.1 to 15.0 h -1 ; and / or, the hydrogen-oil volume ratio is 100:1 to 2500:

1.

16. The method according to claim 1, wherein The reaction conditions in the normal paraffin conversion reaction zone in step (2) are as follows: the reaction pressure is 6 to 8 MPa; the average reaction temperature is 300 to 400 °C; and / or, the liquid hourly space velocity is 1.0 to 5.0 h -1 ; and / or, the hydrogen-to-oil volume ratio is 400:1 to 2000:

1.

17. The method according to claim 1, wherein In the second hydrocracking catalyst, based on the weight of the second hydrocracking catalyst, the content of the hydrogenation component in terms of oxide is 5 wt% to 40 wt%; the content of the cracking component is 20 wt% to 60 wt%; the content of the binder is 10 wt% to 50 wt%; and / or, in the second hydrocracking catalyst, the hydrogenation component is at least one of the metal, metal oxide, and metal sulfide of the active metal component; the active metal component includes Group VIB and / or Group VIII metals.

18. The method according to claim 17, wherein In the second hydrocracking catalyst, the active metal component is at least one of iron, chromium, molybdenum, tungsten, cobalt, and nickel; And / or, in the second hydrocracking catalyst, the binder is alumina and / or silica.

19. The method according to claim 1, characterized in that, In the second hydrocracking catalyst, the volume ratio of the catalyst with Y zeolite as the cracking component to the catalyst with Beta zeolite as the cracking component is 1:1 to 1:

5.

20. The method according to claim 1, wherein In the second hydrocracking catalyst, the volume ratio of the catalyst with Y zeolite as the cracking component to the catalyst with Beta zeolite as the cracking component is 1:2 to 1:

4.

21. The method according to claim 1, wherein The reaction conditions in the hydrocracking reaction zone in step (2) are as follows: the reaction pressure is 5 to 9 MPa; the average reaction temperature is 250 to 450 °C; and / or, the liquid hourly space velocity is 0.1 to 15.0 h -1 ; and / or, the hydrogen-to-oil volume ratio is 100:1 to 2500:

1.

22. The method according to claim 1, wherein The reaction conditions in the hydrocracking reaction zone in step (2) are as follows: the reaction pressure is 6 - 8 MPa; the average reaction temperature is 300 - 400 °C; and / or, the liquid hourly space velocity is 1.0 - 5.0 h -1 ; and / or, the hydrogen - to - oil volume ratio is 400:1 - 2000:

1.

23. The method according to claim 1, wherein The hydrocracking reaction effluent obtained in step (3) is a hydrocracking product containing monocyclic cyclic hydrocarbons.

24. The method according to claim 23, wherein In the hydrocracking reaction effluent of step (3), the mass ratio of C6-C8 monocyclic cyclic hydrocarbons to the total cyclic hydrocarbon mass in the coker light distillate oil feedstock is 0.35 to 0.

55.

25. The method according to claim 23, wherein In the hydrocracking reaction effluent of step (3), the mass ratio of C6-C8 monocyclic cyclic hydrocarbons to the total cyclic hydrocarbon mass in the coker light distillate oil feedstock is 0.41 to 0.

50.

26. The method according to claim 1, wherein The normal paraffin conversion reaction zone in step (2) and the hydrocracking reaction zone in step (3) adopt the same pressure.

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