Method and device for utilizing inferior aromatic-rich distillate oil

Abstract

The invention relates to the field of ethylene byproduct utilization, in particular to a method and a device for utilizing inferior aromatic-rich distillate oil. The method comprises the following steps: sequentially carrying out first-stage hydroconversion of diene and bromine number reduction, second-stage hydrorefining, first purification, third-stage hydrocracking, second purification and separation on an inferior aromatic-rich distillate oil raw material to obtain BTX and LPG. According to the method disclosed by the invention, the inferior aromatic-rich distillate oil raw material is sequentially subjected to first-stage hydroconversion of diene and reduction of bromine number, second-stage hydrorefining, first purification, third-stage hydrocracking and second purification, so that the inferior aromatic-rich distillate oil can be fully utilized, and the additional value of the inferior aromatic-rich distillate oil is increased; the quality of a refined product can be adjusted and improved, the yield of a product BTX can be improved, and the service life of a cracking catalyst can be prolonged while chemical utilization of ethylene cracking byproducts is realized and the competitiveness of preparing ethylene through steam cracking is improved.

Description

Technical Field

[0001] The invention relates to the field of ethylene by-product utilization, and in particular to a method and a device for utilizing inferior aromatic-rich fraction oil. Background Art

[0002] Inferior aromatic fractions are mainly industrial cracking byproducts, including cracking C9 and ethylene tar fractions. Cracked C9 is mainly derived from the cracked gasoline C9 fraction separated after passing through the BTX tower, with an aromatic content of more than 70%, accounting for 11%-22% of ethylene production. Most of the domestic cracked C9 is only sold as cheap primary fuel oil or after preliminary processing.

[0003] Ethylene tar, also known as cracking tar, is a high-boiling liquid product of the cracking raw material in the process of steam cracking to produce ethylene. It belongs to the diesel fraction (205-360℃) and mainly comes from the quench oil tower kettle and the heavy fuel oil gas stripping tower kettle. Ethylene tar is a heavy distillate oil rich in aromatics, mainly containing monocyclic heavy aromatics, polycyclic or condensed aromatic components. It has a complex composition, is easy to polymerize, and has high colloid, heavy metal and ash content, so it cannot be used directly. The yield of ethylene tar varies with different cracking raw materials, generally accounting for about 1 / 5 of the ethylene output. As the ethylene raw material becomes heavier, its yield tends to increase.

[0004] The main ways to utilize ethylene tar include using it as a component for blending gasoline and diesel, as a fuel, producing carbon black, extracting naphthalene and methylnaphthalene, and producing aromatic solvent oil. Among them, using it as a fuel is the main way to utilize ethylene tar in China. Not only is the economic utilization rate low, but ethylene tar is also prone to produce black smoke when it burns, causing pollution to the surrounding environment; producing carbon black is the main way to utilize ethylene tar abroad, with considerable economic benefits, but domestic technology mainly relies on imports; the process yield of recovering naphthalene developed domestically is relatively low, and the production of methylnaphthalene has shortcomings such as small scale, high energy consumption, and low product purity; hydrogenation of ethylene tar can produce high-value-added BTX aromatics, which can greatly improve the utilization rate of ethylene tar, help to increase the value of ethylene by-products, and has a good market prospect. Some foreign companies have begun to use cracking fuel oil to produce aromatic solvent oil. The main manufacturers include Exxon in the United States, Shell in the Netherlands, and Maruzen Oil Company in Japan. In general, ethylene tar is mainly used for low-value fuel purposes. As environmental regulations become increasingly stringent, the use of ethylene tar fuel will become increasingly restricted due to its untreated, sulfur- and nitrogen-rich, and highly unsaturated nature, and its disposal will become problematic.

[0005] The yield of each distillation section of ethylene tar between 205℃ and 300℃ is relatively high, accounting for about 70%, followed by colloid and asphaltene components. At the same time, ethylene tar has high sulfur content, high content of polycyclic aromatic hydrocarbons, high density, and short side chains of aromatic compounds. The main components of the distillation section between the initial distillation point and 205℃ are indene and its homologues, the distillation section between 205℃ and 225℃ is naphthalene, the distillation section between 225℃ and 245℃ is mainly methylnaphthalene, the distillation section between 245℃ and 300℃ is mainly dimethylnaphthalene, the distillation section between 300℃ and 360℃ contains a large amount of anthracene, acenaphthene, phenanthrene, etc., and the substances above 360℃ are mainly colloid and asphaltene with a very high carbon-hydrogen ratio. It can be seen that each distillation section of ethylene tar is an important raw material for chemical organic synthesis, from which many valuable chemical products can be extracted, and it has important utilization value.

[0006] In the field of heavy distillate oil hydroprocessing, catalytic cracking feedstock hydroprocessing pretreatment technology has been industrially applied since the 1970s and has been applied in many refineries processing sulfur-containing or high-sulfur crude oil. At present, there are mature catalytic cracking feedstock pretreatment technologies at home and abroad, mainly including: UOP's VGO Unionfining and APCU (partial conversion hydrocracking) technology, Haldor The company's Aroshift technology, Chevron's VGOHydrotreating technology, Exxon's VGO Hydrodesulfurization technology, IFP's T-star technology, and Mobil, AKZO, Kellogg's MAKfinging technology. Except for T-star, which uses a fluidized bed reaction process, the general FCC feedstock hydrogenation pretreatment unit uses a fixed bed process. In order to further improve product quality and conversion rate, the catalytic feedstock hydrogenation pretreatment process has gradually changed from traditional hydrodesulfurization refining (HDS) to mild hydrocracking (MHC) to improve the denitrification, residual carbon and polycyclic aromatic hydrocarbons saturation capabilities, which is more flexible to operate and has more significant economic benefits.

[0007] CN101724458A discloses a method for hydrogenating ethylene tar. The method comprises the following steps: the ethylene tar is fractionated into a light component and a heavy component after passing through a fractionating tower, the light component is sequentially passed through a hydrogenation protection catalyst, a hydrorefining catalyst, a light component hydrogenation carbon removal catalyst and a hydrocracking catalyst, and the heavy component is sequentially passed through a hydrogenation protection catalyst, a hydroremoval carbon removal catalyst and a hydroconversion catalyst, and finally a gasoline component and a diesel fraction are obtained.

[0008] CN112745949A discloses a method and system for combined processing of deoiled asphalt and aromatic fraction oil. In this method, after heavy oil is deasphalted by solvent, the deasphalted oil passes through hydrogenation reactor I and enters DCC unit to generate propylene, LCO and HCO, etc. After the deoiled asphalt passes through hydrogenation reactor II, the product is fractionated, the light component enters the hydrofining reactor to obtain gasoline and diesel components, and the heavy component enters the delayed coking unit to generate coker gasoline, coker diesel, coker wax oil, etc.

[0009] US5300212A discloses a process for hydrogenation and reforming of inferior heavy oil, in which heavy oil raw material, hydrogen and catalyst are converted in two reactors. Specifically, the raw oil and a dispersed catalyst with molybdenum phosphate as precursors first enter a first slurry bed hydrogenation reactor, and a conversion reaction is carried out under the conditions of 343-482°C and 0.345-34.5Mpa. The reaction product enters a second ebullating bed hydrogenation reactor after separation, and is converted under the conditions of 343-399°C, 5.5-27.6Mpa and a supported catalyst. The reaction product enters a distillation tower to obtain a fraction of <524°C and a fraction of >524°C, wherein the fraction of <524°C is used as a product, and the heavy fraction of >524°C is recycled to the second reactor. The process can achieve the reforming of inferior heavy oil, but the first reactor of the method adopts a dispersed catalyst and the second reactor adopts a supported catalyst. The catalyst particles brought out of the first reactor are easy to block the pores of the heavy supported catalyst of the second reactor or cover the active center of the catalyst, causing the deactivation of the catalyst, affecting the operation cycle of the whole process.

[0010] CN102134500B discloses a method for extracting naphthalene, 1-methylnaphthalene and 2-methylnaphthalene from ethylene tar. The method uses ethylene tar with a naphthalene and methylnaphthalene content of not less than 30% as raw material, and uses five distillation towers in series to sequentially separate naphthalene (>95%), tetramethylnaphthalene, 2-methylnaphthalene (98%) and 1-methylnaphthalene (98%).

[0011] CN109679011B discloses a method for producing a copolymerized petroleum resin. The method uses ethylene tar to produce C9-C 10 The fraction is mixed with high-purity C5, cresol and styrene from which cyclopentadiene has been removed, and BF3 is used as a catalyst. The upper mixed liquid is polymerized, vacuum distilled and the fraction is cooled to obtain a copolymerized petroleum resin. This method can flexibly adjust the molecular weight of the petroleum resin by adjusting the amount of cresol and styrene, so that petroleum resins with different properties can meet market demand.

[0012] CN113755211A discloses a method for producing needle coke by optimizing raw materials containing ethylene tar. The method uses aromatic hydrocarbons, asphalt and colloid in ethylene tar for hydrocracking reaction, and after the products are separated into high fraction, low fraction and gas-liquid, the oil phase enters the coking tower to generate needle coke, and the gas phase is circulated to the hydrogenation tower or used as fuel.

[0013] CN1970688B discloses a comprehensive processing technology for ethylene tar. In this process, after ethylene tar passes through a pre-fractionation tower, the light component enters a hydrogenation refining tower at 260-280°C, and the refined product enters four series-connected distillation towers to obtain solvent oil I, refined naphthalene, β-methylnaphthalene, α-methylnaphthalene, mixed methylnaphthalene, and solvent oil II.

[0014] US20080083649A1 discloses the upgraded utilization of tar. In the invention, ethylene tar is passed through a reduced pressure tubular furnace, the asphaltene fraction at the top of the tower is removed as fuel gas, the asphaltene fraction at the bottom of the tower is used to produce synthesis gas, and part of it is fed into a coking unit to produce coking gasoline and diesel.

[0015] A comprehensive analysis of the above technologies shows that the existing technologies involve the production of gasoline and diesel fractions, petroleum coke, resins, solvent oil, etc. from inferior aromatic-rich oil products, wasting precious aromatic resources in the aromatic-rich fraction oil.

[0016] Based on the technologies of selective hydrorefining and hydrocracking of aromatic-rich distillate oil, optimization and innovation are carried out to maximize the production of benzene (B), toluene (T) and xylene (X). It can make full use of inferior aromatic-rich distillate oils such as ethylene tar and cracking C9, increase their added value, realize the chemical utilization of ethylene cracking by-products, and enhance the competitiveness of steam cracking to produce ethylene. Summary of the invention

[0017] The purpose of the present invention is to overcome the problems in the prior art of using inferior aromatic-rich distillate oil as refinery fuel or producing solvent oil, gasoline and diesel blending components, etc., such as low product added value and serious environmental pollution, and to provide a method and device for utilizing inferior aromatic-rich distillate oil. The method can make full use of inferior aromatic-rich distillate oil, improve its added value, realize the chemical utilization of ethylene cracking by-products, enhance the competitiveness of steam cracking to produce ethylene, and expand a new path for producing BTX and LPG.

[0018] In order to achieve the above-mentioned purpose, the present invention provides a method for utilizing inferior aromatic-rich distillate oil on the one hand, which comprises subjecting the inferior aromatic-rich distillate oil raw material to a first-stage hydrogenation conversion of dienes and reduction of bromine value, a second-stage hydrogenation refining, a first purification, a third-stage hydrocracking, and a second purification in sequence to separate and obtain BTX and LPG.

[0019] The second aspect of the present invention provides a device for utilizing inferior aromatic-rich distillate oil, the device comprising a first stage hydrogenation device, a second stage hydrorefining device, a first purification device, a third stage hydrocracking device, and a second purification device arranged in series; wherein:

[0020] The first stage hydrogenation unit is used to remove dienes from the inferior aromatic-rich distillate oil feedstock and reduce the bromine value to obtain a first stage hydrogenation product stream;

[0021] The temperature of the product stream outlet of the two-stage hydrotreating device is higher than the temperature of the raw material inlet, and the temperature difference is not greater than 75° C., and is used to purify the first-stage hydrogenation product stream, so that the naphthalene series products are highly selectively converted into tetrahydronaphthalene series products and the sulfur and nitrogen content of the refined product is less than 1 ppm, thereby obtaining a two-stage hydrotreating product stream;

[0022] The first purification device is used to remove hydrogen sulfide from the second-stage hydrorefining product stream to obtain a monocyclic aromatic hydrocarbon mixture;

[0023] The three-stage hydrocracking unit is used to crack the monocyclic aromatic hydrocarbon mixture, so that the tetralin series is hydrocracking to generate BTX and LPG, and obtain a three-stage hydrocracking product stream;

[0024] The second purification device is used to purify and separate the three-stage hydrocracking product stream to obtain BTX and LPG.

[0025] Through the above technical solution, the present invention has the following beneficial effects:

[0026] By the method of the present invention, the inferior aromatic-rich distillate oil raw material is subjected to a first-stage hydrogenation conversion of dienes and reduction of bromine value, a second-stage hydrogenation refining, a first purification, a third-stage hydrocracking, and a second purification in sequence, so that the inferior aromatic-rich distillate oil can be fully utilized and its added value can be improved. While realizing the chemical utilization of ethylene cracking by-products and improving the competitiveness of steam cracking to produce ethylene, the quality of refined products can be adjusted and improved, the yield of product BTX can be increased, and the life of the cracking catalyst can be extended. BRIEF DESCRIPTION OF THE DRAWINGS

[0027] Figure 1 a is a TEM image of catalyst A2 in Example 1, Figure 1 b is the TEM image of A5 in Example 10;

[0028] Figure 2 This is the XRD pattern of the B1 carrier HZSM-5 / MCM-41 molecular sieve;

[0029] Figure 3 is the H2-TPR spectrum of B1;

[0030] Figure 4 This is a simplified flow chart of the utilization of inferior aromatic-rich distillate oil. DETAILED DESCRIPTION

[0031] The endpoints and any values ​​of the ranges disclosed in this article are not limited to the precise ranges or values, and these ranges or values ​​should be understood to include values ​​close to these ranges or values. For numerical ranges, the endpoint values ​​of each range, the endpoint values ​​of each range and the individual point values, and the individual point values ​​can be combined with each other to obtain one or more new numerical ranges, which should be considered as specifically disclosed in this article.

[0032] The invention provides a method for utilizing inferior aromatic-rich distillate oil, which comprises subjecting the inferior aromatic-rich distillate oil raw material to sequentially subjecting the raw material to first-stage hydrogenation conversion to dienes and reduction of bromine value, second-stage hydrogenation refining, first purification, third-stage hydrocracking, and second purification, and separating and obtaining BTX and LPG.

[0033] By the method of the present invention, the inferior aromatic-rich distillate oil raw material is subjected to a first-stage hydrogenation conversion of dienes and reduction of bromine value, a second-stage hydrogenation refining, a first purification, a third-stage hydrocracking, and a second purification in sequence, so that the inferior aromatic-rich distillate oil can be fully utilized and its added value can be improved. While realizing the chemical utilization of ethylene cracking by-products and improving the competitiveness of steam cracking to produce ethylene, the quality rate of refined products can be adjusted and improved, the yield of product BTX can be increased, and the life of the cracking catalyst can be extended.

[0034] One of the important factors for maximizing the selectivity of tetralin is the reaction temperature. Low temperature is conducive to the preparation of the desired tetralin. However, the hydrogenation of aromatic-rich distillate is a highly exothermic reaction, so it is very important to limit the increase in the temperature of the aromatic-rich distillate hydrotreating reaction.

[0035] According to a preferred embodiment of the present invention, the temperature of the product stream of the second-stage hydrofining is higher than the temperature of the feed stream before the second-stage hydrofining, and the temperature difference is not greater than 75° C. By adopting the above preferred scheme, the quality of the refined product can be further improved, the yield of the product BTX can be increased, and the life of the cracking catalyst can be extended.

[0036] In the present invention, as long as the temperature of the product stream of the aforementioned two-stage hydrorefining is higher than the temperature of the feed stream before the two-stage hydrorefining, and the temperature difference is not greater than 75°C, the purpose of the present invention can be achieved. According to a preferred embodiment of the present invention, in order to further improve the quality of the refined product, increase the yield of the product BTX, and extend the life of the cracking catalyst, the temperature difference is controlled to be 30-75°C, preferably 35-60°C.

[0037] According to a preferred embodiment of the present invention, the temperature of the hydrofining product stream is 250-310° C., preferably 260-300° C. By adopting the above preferred scheme, the quality of the refined product can be further improved, the yield of the product BTX can be increased, and the life of the cracking catalyst can be extended.

[0038] According to a preferred embodiment of the present invention, the temperature of the feedstock stream before hydrofining is 200-260° C., preferably 210-250° C. By adopting the above preferred scheme, the quality of the refined product can be further improved, the yield of the product BTX can be increased, and the life of the cracking catalyst can be extended.

[0039] In the present invention, as long as the purpose of the present invention can be achieved, there is no particular limitation on the final distillation point of the inferior aromatic-rich distillate oil feedstock. According to a preferred embodiment of the present invention, the final distillation point of the inferior aromatic-rich distillate oil feedstock is ≯300° C. By adopting the above preferred scheme, the quality of the refined product can be further improved, the yield of the product BTX can be increased, and the life of the cracking catalyst can be extended.

[0040] In the present invention, any inferior aromatic-rich distillate oil raw material can be used in the present invention. According to a preferred embodiment of the present invention, the inferior aromatic-rich distillate oil raw material is selected from at least one of a cracking C9 product stream, a cracking ethylene tar product stream, and a cracking coal tar product stream. By adopting the aforementioned preferred scheme, the quality of the refined product can be further improved, the yield of the product BTX can be increased, and the life of the cracking catalyst can be extended.

[0041] According to a preferred embodiment of the present invention, the feedstock stream before the second-stage hydrofining includes the first purified stream. By adopting the above preferred scheme, the quality of the refined product can be further improved, the yield of the product BTX can be increased, and the life of the cracking catalyst can be extended.

[0042] According to a preferred embodiment of the present invention, the mass content of the first purified stream in the feed stream before the second stage hydrofining is 0-90%, preferably 0-50%. By adopting the above preferred scheme, the quality of the refined product can be further improved and the yield of the product BTX can be increased.

[0043] According to a preferred embodiment of the present invention, the bromine value of the product stream after the first stage hydrogenation is ≯40mgBr / 100goil, and the colloid is ≯50mg / 100goil. By adopting the above preferred scheme, the quality of the refined product can be further improved, the yield of the product BTX can be increased, and the life of the cracking catalyst can be extended.

[0044] According to a preferred embodiment of the present invention, the sulfur content in the first purified product stream is less than 1 ppm, preferably less than 0.85 ppm. By adopting the above preferred scheme, the quality of the refined product can be further improved, the yield of the product BTX can be increased, and the life of the cracking catalyst can be extended.

[0045] According to a preferred embodiment of the present invention, the nitrogen content in the first purified product stream is less than 1 ppm, preferably less than 0.8 ppm. By adopting the above preferred scheme, the quality of the refined product can be further improved, the yield of the product BTX can be increased, and the life of the cracking catalyst can be extended.

[0046] In the present invention, the two-stage hydrorefining catalyst can be a conventional choice in the art. According to a preferred embodiment of the present invention, the two-stage hydrorefining catalyst includes a first active component, an active promoter and a first carrier, wherein the first active component includes a metal of Group VIB and a metal of Group VIII, preferably including at least one of Ni, Co, Fe, Pt and Pd, and Mo and / or W; the active promoter includes a lanthanide metal and / or a Group VA element, preferably La, P, Ce; after the two-stage hydrorefining catalyst is presulfurized, the number of stacking layers of the sulfide active phase in the catalyst is 3-4 layers and presents a tower-shaped structure, and preferably the stacking length of the sulfide active phase is 3-6nm. The active sites of the edges, corners, and edges of the catalyst with this feature are fully exposed, which can further improve the quality of the refined product and increase the yield of the product BTX and LPG.

[0047] In the present invention, the pre-sulfurization conditions can be conventionally selected in the art. According to a preferred embodiment of the present invention, the pre-sulfurization conditions include: pressure 0.4-0.6MPa, N2 is passed at a rate of 20-40℃ / h to 130-170℃, N2 flow rate is 500-1000ml / min, N2 is stopped, H2 and sulfurized oil (sulfurized oil contains CS2 1000-10000ppm), system pressure rises to 2-3MPa, heats up at 5-10℃ / h for 3-6h to 180-200℃, keeps constant temperature for 4-6h, heats up at 5-10℃ / h for 5-10h to 220-240℃, keeps constant temperature for 2-5h, heats up at 5-10℃ / h for 5-10h to 270-300℃, keeps constant temperature for 3-5h, heats up at 5-10℃ / h for 4-8h to 320-340℃, keeps constant temperature for 2-3h. The advantages of the present invention are illustrated by using a pressure of 0.5 MPa, N2 passing at a rate of 30°C / h to raise the temperature to 150°C, the N2 flow rate is 850 ml / min, N2 is stopped, H2 and sulfide oil (the sulfide oil contains 5000 ppm CS2), the system pressure is increased to 2.5 MPa, the temperature is increased at 5°C / h for 6 hours to 180°C, the temperature is kept constant for 5 hours, then the temperature is increased at 5°C / h for 10 hours to 230°C, the temperature is kept constant for 4 hours, then the temperature is increased at 10°C / h for 5 hours to 280°C, the temperature is kept constant for 4 hours, the temperature is increased at 10°C / h for 4 hours to 320°C at 10°C / h for 4 hours, and the temperature is kept constant for 2 hours.

[0048] According to a preferred embodiment of the present invention, the first carrier in the two-stage hydrorefining catalyst is selected from at least one of Al2O3, Al2O3-TiO2 (e.g., Al2O3 90-99%, TiO2 1-10%), Al2O3-TiO2-SiO2 (e.g., Al2O3 60-98%, TiO2 1-10%, SiO2 1-30%), and Al2O3-SiO2 (e.g., Al2O3 70-99%, SiO2 1-30%), and is further preferably selected from one or two of Al2O3, Al2O3-TiO2, and Al2O3-TiO2-SiO2; by adopting the aforementioned preferred scheme, the quality of the refined product can be further improved, the yield of the product BTX can be increased, and the life of the cracking catalyst can be extended.

[0049] According to a preferred embodiment of the present invention, the average pore size of the first carrier in the two-stage hydrorefining catalyst is 7-20 nm. By adopting the above preferred scheme, the quality of the refined product can be further improved, the yield of the product BTX can be increased, and the life of the cracking catalyst can be extended.

[0050] According to a preferred embodiment of the present invention, the pore volume of the first carrier in the two-stage hydrorefining catalyst is 0.5-0.98 cm 3 By adopting the above preferred scheme, the quality of the refined product can be further improved, the yield of the product BTX can be increased, and the life of the cracking catalyst can be extended.

[0051] According to a preferred embodiment of the present invention, the specific surface area of ​​the first carrier in the two-stage hydrotreating catalyst is 190-380m 2 / g; By adopting the above preferred scheme, the quality of the refined product can be further improved, the yield of the product BTX can be increased, and the life of the cracking catalyst can be extended.

[0052] In the present invention, the amount of each component of the two-stage hydrorefining catalyst is not particularly limited. According to a preferred embodiment of the present invention, in the two-stage hydrorefining catalyst, the first active component is 50-300 g / L in terms of oxide, preferably 100-300 g / L; the active promoter is 5-50 g / L in terms of oxide, preferably 5-40 g / L. By adopting the above preferred scheme, the quality of the refined product can be further improved, the yield of the product BTX can be increased, and the life of the cracking catalyst can be extended.

[0053] In the present invention, as long as the above-mentioned hydrorefining catalyst of the present invention can be prepared, the preparation method can be a conventional choice in the art. According to a preferred embodiment of the present invention, the preparation method of the two-stage hydrorefining catalyst includes: the preparation method of the two-stage hydrorefining catalyst includes: mixing the first active component source, the active auxiliary agent source, the first carrier source and the additive to obtain a mixed liquid, drying and roasting; wherein the concentration of the Group VIB metal in the mixed liquid is 150-850g / L in terms of oxide; the concentration of the Group VIII metal in terms of oxide is 50-300g / L; the additive is a mixture of an amino compound, an organic acid metal complexing agent and an alcohol. By adopting the above-mentioned preferred scheme, the quality of the refined product can be further improved, the yield of the product BTX can be increased, and the life of the cracking catalyst can be extended.

[0054] According to a preferred embodiment of the present invention, the additive is selected from at least one of urea, acetamide, ethylenediamine, ethanolamine, diethanolamine, triethanolamine, and at least one of 1,2-cyclohexanediaminetetraacetic acid, citric acid, tartaric acid, ethylenediaminetetraacetic acid, and ethylene glycol and / or ethanol. By adopting the above preferred embodiment, the quality of the refined product can be further improved, the yield of the product BTX can be increased, and the life of the cracking catalyst can be extended.

[0055] According to a preferred embodiment of the present invention, the mass fraction of the additive in the mixed solution is 0.01-10%, more preferably 1.5-7.5%. By adopting the above preferred solution, the quality of the refined product can be further improved, the yield of the product BTX can be increased, and the life of the cracking catalyst can be extended.

[0056] According to a preferred embodiment of the present invention, the mass ratio of the amino compound, the organic acid metal complexing agent and the alcohol in the additive is 1-2:1-2:1-2.

[0057] In the present invention, the three-stage hydrocracking catalyst can be a conventional choice in the art. According to a preferred embodiment of the present invention, the three-stage hydrocracking catalyst includes a second active component and a second carrier, the second active component includes a Group VIII metal, preferably Ni; the second carrier includes a layered HZSM-5 / MCM-41 molecular sieve and optional Ce oxide and Zr oxide. By adopting the above preferred scheme, the yield of the product BTX is increased and the life of the cracking catalyst is extended.

[0058] According to a preferred embodiment of the present invention, the interlayer thickness of the layered ZSM-5 / MCM-41 molecular sieve is 5-10 nm, preferably 5-7 nm. By adopting the above preferred solution, the yield of the product BTX is increased and the life of the cracking catalyst is extended.

[0059] According to a preferred embodiment of the present invention, the pore size of the layered ZSM-5 / MCM-41 molecular sieve is 3-18 nm, preferably 3-11 nm. By adopting the above preferred scheme, the yield of the product BTX is increased and the life of the cracking catalyst is extended.

[0060] According to a preferred embodiment of the present invention, the medium-strong acid content / total acid content of the layered ZSM-5 / MCM-41 molecular sieve is 0.55-0.85, preferably 0.55-0.65. By adopting the above preferred scheme, the yield of the product BTX is increased and the life of the cracking catalyst is extended.

[0061] According to a preferred embodiment of the present invention, the specific surface area of ​​the layered ZSM-5 / MCM-41 molecular sieve is ≮400m 2 ·g -1 , preferably 400-550m 2 ·g -1 By adopting the above preferred scheme, the yield of the product BTX is increased and the life of the cracking catalyst is extended.

[0062] According to a preferred embodiment of the present invention, the SiO2 / Al2O3 molar ratio of the layered ZSM-5 / MCM-41 molecular sieve is 20-200, preferably 20-60. By adopting the above preferred scheme, the yield of the product BTX is increased and the life of the cracking catalyst is extended.

[0063] In the present invention, the amount of each component of the three-stage hydrocracking catalyst is not particularly limited. According to a preferred embodiment of the present invention, in the three-stage hydrocracking catalyst, the content of the second active component in terms of oxide is 50-260 g / L, preferably 50-150 g / L; the content of Ce oxide and Zr oxide in terms of oxide is 0.1-100 g / L, preferably 0.1-50 g / L. By adopting the above preferred scheme, the yield of the product BTX is increased and the life of the cracking catalyst is extended.

[0064] According to a preferred embodiment of the present invention, the TPR hydrogen atmosphere reduction temperature of the three-stage hydrocracking catalyst is lower than 400° C., and more preferably lower than 390° C. By adopting the above preferred scheme, the yield of the product BTX is increased and the life of the cracking catalyst is extended.

[0065] According to a preferred embodiment of the present invention, the dispersion of the second active component element in the three-stage hydrocracking catalyst is greater than 8%, preferably 8-15%. By adopting the above preferred scheme, the yield of the product BTX is increased and the life of the cracking catalyst is extended.

[0066] In the present invention, as long as the three-stage hydrocracking catalyst of the present invention can be prepared, the preparation method can be a conventional choice in the art. According to a preferred embodiment of the present invention, the preparation method of the three-stage hydrocracking catalyst includes: mixing, drying and roasting the second active component source, the second carrier source and the chelating surfactant. By adopting the above preferred scheme, the yield of the product BTX is improved and the life of the cracking catalyst is extended.

[0067] According to a preferred embodiment of the present invention, the chelating surfactant is selected from at least one of 1,2-cyclohexanediaminetetraacetic acid, citric acid, tartaric acid, and ethylenediaminetetraacetic acid, preferably one of 1,2-cyclohexanediaminetetraacetic acid, citric acid, and ethylenediaminetetraacetic acid. By adopting the above preferred scheme, the yield of the product BTX is increased and the life of the cracking catalyst is extended.

[0068] According to a preferred embodiment of the present invention, the amount of the chelating surfactant in the mixed material is 0.01-10%, more preferably 1.5-5%. By adopting the above preferred solution, the yield of the product BTX is increased and the life of the cracking catalyst is extended.

[0069] According to a preferred embodiment of the present invention, the conditions for the second-stage hydrofining include: a temperature of 200-270°C, preferably 210-250°C.

[0070] According to a preferred embodiment of the present invention, the conditions of the second-stage hydrofining include: a pressure of 2.0-8.0 MPa, preferably 2.2-6.0 MPa.

[0071] According to a preferred embodiment of the present invention, the conditions of the second-stage hydrotreating include: a hydrogen-to-oil volume ratio of 400-3000, preferably 600-2000.

[0072] According to a preferred embodiment of the present invention, the conditions of the second-stage hydrofining include: the liquid phase volume space velocity is 0.2-3h -1 , preferably 0.6-1.2h -1 .

[0073] According to a preferred embodiment of the present invention, the conditions of the three-stage hydrocracking include: a pressure of 2-8 MPa, preferably 3.0-6.5 MPa.

[0074] According to a preferred embodiment of the present invention, the conditions of the three-stage hydrocracking include: a space velocity of 0.2-3h -1 , preferably 0.6-1.2h -1 .

[0075] According to a preferred embodiment of the present invention, the conditions of the three-stage hydrocracking include: a temperature of 260-500°C, preferably 260-480°C.

[0076] According to a preferred embodiment of the present invention, the conditions of the three-stage hydrocracking include: a hydrogen-to-oil volume ratio of 400-3000, preferably 600-2000.

[0077] In the present invention, there is no special requirement for the two-stage hydrofining method. According to a preferred embodiment of the present invention, the two-stage hydrofining method is that the material passes through at least two layers of catalyst along the flow direction, and the pore size of the first carrier in the catalyst decreases successively. By adopting the above preferred scheme, the quality of the refined product can be further improved, the yield of the product BTX can be increased, and the life of the cracking catalyst can be extended.

[0078] The second aspect of the present invention provides a device for utilizing inferior aromatic fraction oil for the method of the present invention, the device comprising a first stage hydrogenation device, a second stage hydrorefining device, a first purification device, a third stage hydrocracking device, and a second purification device arranged in series; wherein:

[0079] The first stage hydrogenation unit is used to convert dienes in the inferior aromatic-rich distillate oil feedstock and reduce the bromine value to obtain a first stage hydrogenation product stream;

[0080] The temperature of the product stream outlet of the two-stage hydrotreating device is higher than the temperature of the raw material inlet, and the temperature difference is not greater than 75° C., and is used to purify the first-stage hydrogenation product stream, so that the naphthalene series products are highly selectively converted into tetrahydronaphthalene series products and the sulfur and nitrogen content of the refined product is less than 1 ppm, thereby obtaining a two-stage hydrotreating product stream;

[0081] The first purification device is used to remove sulfur and nitrogen from the second-stage hydrorefining product stream to obtain a monocyclic aromatic hydrocarbon mixture;

[0082] The three-stage hydrocracking unit is used to crack the monocyclic aromatic hydrocarbon mixture, so that the tetralin series is hydrocracking to generate BTX and LPG, and obtain a three-stage hydrocracking product stream;

[0083] The second purification device is used to purify and separate the three-stage hydrocracking product stream to obtain BTX and LPG.

[0084] By utilizing the inferior aromatic-rich distillate oil raw material through the device of the present invention, the quality of the refined product can be further adjusted and improved, and the yield of the product BTX and LPG can be increased.

[0085] According to a preferred embodiment of the present invention, the monocyclic aromatic hydrocarbon mixture outlet of the first purification device is provided with a circulation pipeline connected to the raw material inlet of the second-stage hydrotreating device. By adopting the above preferred scheme, the quality of the refined product can be further adjusted and improved, and the yield of the product BTX and LPG can be increased.

[0086] According to a preferred embodiment of the present invention, the first purification device and the second purification device are each independently selected from at least one of a high separation tank, an oil-water separation tank, a stripping tower, a low separation tank and a distillation tower.

[0087] According to a preferred embodiment of the present invention, the second-stage hydrogenation device is loaded with a second-stage hydrorefining catalyst, the second-stage hydrorefining catalyst comprises a first active component, an active promoter and a first carrier, the first active component comprises a Group VIB metal and a Group VIII metal, preferably comprises at least one of Ni, Co, Fe, Pt and Pd, and Mo and / or W; the active promoter comprises a lanthanide metal and / or a Group VA element, preferably La, P, Ce; the second-stage hydrorefining catalyst is presulfurized, The number of stacking layers of the sulfide active phase in the catalyst is 3-4 and has a tower-shaped structure. Preferably, the stacking length of the sulfide active phase is 3-6 nm, and the active sites of the edges, corners and sides are fully exposed. The three-stage hydrocracking unit is loaded with a three-stage hydrocracking catalyst, and the three-stage hydrocracking catalyst includes a second active component and a second carrier, and the second active component contains a Group VIII metal, preferably Ni; the second carrier contains a layered HZSM-5 / MCM-41 molecular sieve and optional Ce oxide and Zr oxide.

[0088] like Figure 4 As shown: The present invention exemplarily provides an inferior aromatic fraction oil utilization device, comprising: ethylene tar is subjected to weight removal, first-stage hydrogenation to remove dienes, and the product of reduced bromine value is heated by a second-stage heater and then enters a second-stage hydrogenation refining reactor; the product passes through a second-stage high-fraction tank, and then the hydrogen enters a second-stage compressor, and the liquid phase enters a second-stage oil-water separation tank, and then enters a stripping tower to remove hydrogen sulfide gas; the second-stage purified product enters a second-stage hydrogenation refining reactor through a second-stage circulation pump, and enters a third-stage heater and then enters a third-stage hydrocracking reactor; the three-stage products are successively passed through three-stage high-fraction and low-fraction and then enter a distillation tower to obtain the final products BTX and LPG.

[0089] The present invention will be described in detail below through examples.

[0090] In the following examples, the method for evaluating the hydrogenation of aromatic-rich cracking distillate oil includes:

[0091] Aromatic fraction oil hydrorefining raw materials: distillation range 165-300°C, sulfur content = 500ppm, nitrogen content = 40ppm; raw material composition: naphthalene series ≮40%.

[0092] Evaluation conditions:

[0093] Hydrofining: liquid phase volume space velocity V = 0.8 pressure; P = 4.0 MPa; H2 / Oil (v / v) = 1000; and / or the mass ratio of the second stage refined product circulation to the second stage liquid phase feed is 0 to 2;

[0094] Hydrocracking: temperature T = 400°C; liquid phase volume space velocity V = 0.8; H2 / Oil (v / v) = 600; pressure P = 4.0 MPa.

[0095] In the present invention, the dispersion test method of the active group Ni is the hydroxyl titration method.

[0096] R=[Ni] / [Ni] 总 =2 / 3×V0×N A ×1 / 22400W / P / (N A ×1 / M)

[0097] Where:

[0098] R-----Ni dispersion;

[0099] [Ni]-----Number of nickel atoms on the surface;

[0100] [Ni] 总 -----Total number of nickel atoms;

[0101] V0-----titration of hydrogen, mL;

[0102] N A -----Avogadro constant (6.02) × 10 23 ;

[0103] W-----sample mass, g;

[0104] P-----mass fraction of nickel in the sample, %;

[0105] M-----The atomic weight of nickel is 58.7.

[0106] In the present invention, the test method for the reduction of the hydrocracking catalyst TPR in hydrogen atmosphere is the hydrogen-oxygen titration method.

[0107] In the following implementation, a German Bruker D8 Advance X-ray diffractometer was used to perform phase analysis of the catalyst sample, with an operating voltage of 30 kV, a current of 30 mA, and a scanning range of 0° to 80°.

[0108] The H2-TPR analysis was performed using an AutoChem 2920 dynamic adsorption instrument from Micromeritics Instruments, USA. The reducing gas was a 10% H2-Ar mixed gas by volume, the sample mass was 50 mg, the gas flow rate was 50 ml / min, and the temperature was raised from room temperature to 800°C at a heating rate of 10°C / min.

[0109] NH3-TPD was measured using the Altamira AMI-3300 chemical adsorption instrument from Micromeritics Instruments, USA;

[0110] The pore structure parameters of the carrier were tested using a Tristar 3000 surface area analyzer from Micrometrics Company at a test temperature of -196°C. The sample was vacuum activated at 300°C for 6 hours before the test.

[0111] The structural characteristics of the catalyst and the size of the metal nanoparticles were tested using a G2F30 transmission electron microscope (TEM) from FEI Company of the United States.

[0112] In the following embodiments, unless otherwise specified, "%" refers to mass percentage.

[0113] Example 1

[0114] 1. Preparation of Hydrorefining Catalyst

[0115] 1L of carrier Al2O3 (pore size: 16nm) and carrier Al2O3-TiO2-SiO2 (pore size: 11nm, Al2O3 67%, TiO2 5%, SiO2 28%) were taken respectively and mixed with 0.75L of a mixed solution of nickel acetate, ammonium metatungstate, lanthanum nitrate, ammonium molybdate, 1,2-cyclohexanediaminetetraacetic acid, urea, ethanol and phosphoric acid (containing NiO 40g, MoO3 60g, La2O3 5g, WO3 165g, 1,2-cyclohexanediaminetetraacetic acid, urea, ethanol in total 129g, P2O5 10g), wherein the total amount of 1,2-cyclohexanediaminetetraacetic acid, urea and ethanol (mass ratio 1:1:1) was 7.5% of the mass of the mixed solution. The mixture was dried at 110℃ for 6 hours and calcined at 450℃ for 4 hours to prepare catalysts A1 and A2 respectively. The TEM image of A2 after vulcanization is shown in Figure 1 a. It can be seen from the figure that the active phase of the sulfided catalyst presents a tower-shaped structure.

[0116] A1 and A2 catalysts contain NiO 40g / L, MoO3 60g / L, La2O3 5g / L, WO3 165g / L, and P2O5 10g / L.

[0117] 1L of carrier Al2O3-SiO2 (pore size: 10nm, Al2O3 95%, SiO2 5%) was mixed with 0.75L of a mixed solution of ammonium molybdate, cobalt nitrate, cobalt acetate, 1,2-cyclohexanediaminetetraacetic acid and phosphoric acid (containing 96g MoO3, 25g CoO, 5g La2O3, 52g 1,2-cyclohexanediaminetetraacetic acid, urea, and ethanol in total, and 5g P2O5), wherein the total amount of 1,2-cyclohexanediaminetetraacetic acid, urea, and ethanol was 3.0% of the mass of the mixed solution. The mixture was dried at 110°C for 6 hours and calcined at 450°C for 4 hours to prepare catalyst A3.

[0118] A3 catalyst contains MoO3 96g / L, La2O3 5g / L, CoO 25g / L, and P2O5 5g / L.

[0119] 2. Preparation of hydrocracking catalyst

[0120] Take 1L of the formed hydrogen-type layered HZSM-5 / MCM-41 molecular sieve (SiO2 / Al2O3 molar ratio 30, interlayer thickness 6nm, pore size 7nm, medium-strong acid content / total acid content 0.78) and mix it with 0.4L of a mixed solution of nickel acetate, nickel nitrate, basic nickel carbonate, cerium nitrate, zirconium nitrate, 1,2-cyclohexanediaminetetraacetic acid, etc. (containing NiO 108g, CeO2 8g, ZrO2 6g, 1,2-cyclohexanediaminetetraacetic acid 48g), wherein the amount of 1,2-cyclohexanediaminetetraacetic acid is 7.5% of the mass of the mixed solution. Dry at 110℃ for 6 hours and calcine at 450℃ for 4 hours to prepare catalyst B1. The TPR hydrogen atmosphere reduction temperature of catalyst B1 is measured to be 386℃. Figure 3 shown.

[0121] The XRD spectrum of HZSM-5 / MCM-41 molecular sieve is shown in Figure 2 This indicates that HZSM-5 / MCM-41 has both mesoporous and microporous structures, and the ordered mesoporous channels and strong acidity of ZSM-5 make it a suitable carrier for hydrocracking catalysts.

[0122] Catalyst B1 contains NiO 108 g / L, CeO2 8 g / L, and ZrO2 6 g / L.

[0123] 3. Hydrorefining catalyst loading

[0124] The volume ratio of A1, A2 and A3 catalysts is 3:5:2, and they are sequentially loaded into the adiabatic bed reactor (two-stage hydrotreating reactor, the same below).

[0125] 4. Sulfidation of Hydrorefining Catalyst

[0126] Under the system pressure of 0.5MPa, N2 was passed and the temperature was raised to 150℃ at a rate of 30℃ / h, the N2 flow rate was 850ml / min, N2 was stopped, H2 and sulfide oil (sulfide oil contains CS2 5000ppm) were passed, the system pressure was raised to 2.5MPa, the temperature was raised to 180℃ at 5℃ / h for 6h, kept constant for 5h, then raised to 230℃ at 5℃ / h for 10h, kept constant for 4h, then raised to 280℃ at 10℃ / h for 5h, kept constant for 4h, then raised to 320℃ at 10℃ / h for 4h, kept constant for 2h, and the catalyst sulfidation was completed. When sulfide oil was passed, the bed inlet temperature was reduced to the temperature required for the reaction.

[0127] The TEM of A2 catalyst is shown in 1a, which shows that the active phase of the hydrogenation catalyst after sulfurization presents a tower-shaped structure, with 3-5 active phase stacking layers and a length of 3-6nm, and the active sites such as edges, corners and sides are fully exposed. Studies have shown that hydrogenation catalysts with this type of active phase structure have high hydrogenation activity and stronger desulfurization and denitrification activity.

[0128] 5. Hydrocracking catalyst loading and reduction

[0129] B1 was loaded into an adiabatic bed reactor (three-stage hydrocracking reactor, the same below), the room temperature was heated to 500°C at a rate of 20°C / h, and the temperature was maintained for 3 hours in a hydrogen atmosphere with a hydrogen gas flow rate of 500 ml / min.

[0130] 6. Catalyst evaluation

[0131] (The inferior aromatic-rich distillate oil raw material comes from the product stream of ethylene tar, and the final distillation point is 300°C) the first-stage hydrogenation product (bromine value is 40mgBr / 100g oil, and gum is 48mg / 100g oil) is mixed with new hydrogen and second-stage circulating hydrogen, and heated in the second-stage heater and then enters the second-stage hydrotreating reactor. The hydrotreating product enters the second-stage high-separation tank, the hydrogen enters the second-stage compressor for circulation, the liquid phase enters the second-stage oil-water separation tank, the oil phase enters the second-stage low-separation tank and is removed from the oil phase by a gas stripping tower. The mass ratio of the second-stage purified product circulation volume to the second-stage liquid phase feed is 0.5. The second-stage purified product is heated by a third-stage heater and then enters a third-stage hydrocracking reactor. The cracking product enters the third-stage high-separation tank, the hydrogen enters the third-stage compressor for circulation, and the liquid phase enters the distillation tower after the third-stage low-separation tank for fractionation to obtain BTX and LPG.

[0132] The temperature of the hydrotreating feedstock stream is 225°C, and the temperature of the hydrotreating product stream is 275°C.

[0133] The evaluation results are listed in Table 1 (purification) and Table 2 (cracking).

[0134] Example 2

[0135] Same as Example 1, except that A1 is replaced by A4. Preparation of Hydrofining Catalyst A4:

[0136] 1L of carrier Al2O3 (pore size: 16nm) was mixed with 0.75L of a mixed solution of nickel acetate, nickel nitrate, ammonium molybdate, lanthanum nitrate, 1,2-cyclohexanediaminetetraacetic acid, urea, ethanol and phosphoric acid (containing 40g NiO, 108g MoO3, 5g La2O3, 36g 1,2-cyclohexanediaminetetraacetic acid, urea and ethanol, and 10g P2O5), wherein the total amount of 1,2-cyclohexanediaminetetraacetic acid, urea and ethanol was 1.5% of the mass of the mixed solution. The mixture was dried at 110°C for 6 hours and calcined at 450°C for 4 hours to prepare catalyst A4.

[0137] A4 catalyst contains NiO 28g / L, MoO3 108g / L, La2O3 5g / L, and P2O5 10g / L.

[0138] The evaluation results are listed in Table 1 (purification) and Table 2 (cracking).

[0139] Example 3

[0140] 1. Hydrorefining catalyst loading

[0141] The volume ratio of A2 and A3 catalysts is 8:2, and they are loaded into the adiabatic bed reactor in sequence.

[0142] 2. Sulfidation of Hydrorefining Catalyst

[0143] Under the system pressure of 0.5MPa, N2 was passed and the temperature was raised to 150℃ at a rate of 30℃ / h, the N2 flow rate was 850ml / min, N2 was stopped, H2 and sulfide oil (sulfide oil contains CS2 5000ppm) were passed, the system pressure was raised to 2.5MPa, the temperature was raised to 180℃ at 5℃ / h for 6h, kept constant for 5h, then raised to 230℃ at 5℃ / h for 10h, kept constant for 4h, then raised to 280℃ at 10℃ / h for 5h, kept constant for 4h, then raised to 320℃ at 10℃ / h for 4h, kept constant for 2h, and the catalyst sulfidation was completed. When sulfide oil was passed, the bed inlet temperature was reduced to the temperature required for the reaction.

[0144] 3. Hydrocracking catalyst loading and reduction

[0145] B1 was loaded into an adiabatic bed reactor, and the room temperature was heated to 500°C at a rate of 20°C / h, and the temperature was kept constant for 3 hours under hydrogen atmosphere at a hydrogen gas rate of 500 ml / min.

[0146] 4. Catalyst evaluation

[0147] (The inferior aromatic-rich distillate oil raw material comes from the product stream of ethylene tar, and the final distillation point is 300°C) the first-stage hydrogenation product (bromine value is 40mgBr / 100g oil, and gum is 48mg / 100g oil) is mixed with new hydrogen and second-stage circulating hydrogen, and heated in the second-stage heater and then enters the second-stage hydrotreating reactor. The hydrotreating product enters the second-stage high-separation tank, the hydrogen enters the second-stage compressor for circulation, the liquid phase enters the second-stage oil-water separation tank, the oil phase enters the second-stage low-separation tank and is removed from the oil phase by a gas stripping tower. The mass ratio of the second-stage purified product circulation volume to the second-stage liquid phase feed is 0.5. The second-stage purified product is heated by a third-stage heater and then enters a third-stage hydrocracking reactor. The cracking product enters the third-stage high-separation tank, the hydrogen enters the third-stage compressor for circulation, and the liquid phase enters the distillation tower after the third-stage low-separation tank for fractionation to obtain BTX and LPG.

[0148] The temperature of the hydrotreating feedstock stream is 225°C, and the temperature of the hydrotreating product stream is 290°C.

[0149] The evaluation results are listed in Table 1 (purification) and Table 2 (cracking).

[0150] Example 4

[0151] 1. Hydrorefining catalyst loading

[0152] Catalysts A4, A2 and A3 are sequentially loaded into the adiabatic bed reactor in a volume ratio of 3:5:2.

[0153] 2. Sulfidation of Hydrorefining Catalyst

[0154] Under the system pressure of 0.5MPa, N2 was passed and the temperature was raised to 150℃ at a rate of 30℃ / h, the N2 flow rate was 850ml / min, N2 was stopped, H2 and sulfide oil (sulfide oil contains CS2 5000ppm) were passed, the system pressure was raised to 2.5MPa, the temperature was raised to 180℃ at 5℃ / h for 6h, kept constant for 5h, then raised to 230℃ at 5℃ / h for 10h, kept constant for 4h, then raised to 280℃ at 10℃ / h for 5h, kept constant for 4h, then raised to 320℃ at 10℃ / h for 4h, kept constant for 2h, and the catalyst sulfidation was completed. When sulfide oil was passed, the bed inlet temperature was reduced to the temperature required for the reaction.

[0155] 3. Hydrocracking catalyst reduction

[0156] B1 was loaded into an adiabatic bed reactor, and the temperature was raised from room temperature to 500°C at a rate of 20°C / h, and kept constant for 3 hours under a hydrogen atmosphere at a hydrogen gas rate of 500 ml / min.

[0157] 4. Catalyst evaluation

[0158] (The inferior aromatic-rich distillate oil raw material comes from the product stream of ethylene tar, and the final distillation point is 300°C) the first-stage hydrogenation product (bromine value is 40mgBr / 100g oil, and gum is 48mg / 100g oil) is mixed with new hydrogen and second-stage circulating hydrogen, and heated in the second-stage heater and then enters the second-stage hydrotreating reactor. The hydrotreating product enters the second-stage high-separation tank, the hydrogen enters the second-stage compressor for circulation, the liquid phase enters the second-stage oil-water separation tank, the oil phase enters the second-stage low-separation tank and then passes through the gas stripping tower to remove H2S in the oil phase. The second-stage purified product is not circulated to the inlet of the hydrotreating reactor. The second-stage purified product is heated by the third-stage heater and enters the third-stage hydrocracking reactor. The cracking product enters the third-stage high-separation tank, the hydrogen enters the third-stage compressor for circulation, the liquid phase enters the third-stage low-separation tank and then enters the distillation tower for distillation to obtain BTX and LPG.

[0159] The temperature of the hydrotreating feedstock stream is 225°C, and the temperature of the hydrotreating product stream is 300°C.

[0160] The evaluation results are listed in Table 1 (purification) and Table 2 (cracking).

[0161] Example 5

[0162] 1. Hydrorefining catalyst loading

[0163] Catalysts A4, A2 and A3 are sequentially loaded into the adiabatic bed reactor in a volume ratio of 3:5:2.

[0164] 2. Sulfidation of Hydrorefining Catalyst

[0165] Under the system pressure of 0.5MPa, N2 was passed and the temperature was raised to 150℃ at a rate of 30℃ / h, the N2 flow rate was 850ml / min, N2 was stopped, H2 and sulfide oil (sulfide oil contains CS2 5000ppm) were passed, the system pressure was raised to 2.5MPa, the temperature was raised to 180℃ at 5℃ / h for 6h, kept constant for 5h, then raised to 230℃ at 5℃ / h for 10h, kept constant for 4h, then raised to 280℃ at 10℃ / h for 5h, kept constant for 4h, then raised to 320℃ at 10℃ / h for 4h, kept constant for 2h, and the catalyst sulfidation was completed. When sulfide oil was passed, the bed inlet temperature was reduced to the temperature required for the reaction.

[0166] 3. Hydrocracking catalyst loading and reduction

[0167] B1 was loaded into an adiabatic bed reactor, and the room temperature was heated to 500°C at a rate of 20°C / h, and the temperature was kept constant for 3 hours under hydrogen atmosphere at a hydrogen gas rate of 500 ml / min.

[0168] 4. Catalyst evaluation

[0169] (The inferior aromatic-rich distillate oil raw material comes from the product stream of cracking carbon nine, and the final distillation point is 275°C) The first-stage hydrogenation product (bromine value is 35mgBr / 100g oil, and gum is 40mg / 100g oil) is mixed with new hydrogen and second-stage circulating hydrogen, and is heated in the second-stage heater and then enters the second-stage hydrotreating reactor. The hydrotreating product enters the second-stage high-separation tank, the hydrogen enters the second-stage compressor for circulation, and the liquid phase enters the second-stage oil-water separation tank. The oil phase enters the second-stage low-separation tank and is removed from the oil phase by the gas stripping tower. The mass ratio of the second-stage purified product circulation volume to the second-stage liquid phase feed is 0.5. The second-stage purified product is heated by the third-stage heater and then enters the third-stage hydrocracking reactor. The cracking product enters the third-stage high-separation tank, the hydrogen enters the third-stage compressor for circulation, and the liquid phase enters the third-stage low-separation tank and then enters the distillation tower for distillation to obtain BTX and LPG.

[0170] The temperature of the hydrotreating feedstock stream is 225°C, and the temperature of the hydrotreating product stream is 270°C.

[0171] The evaluation results are listed in Table 1 (purification) and Table 2 (cracking).

[0172] Example 6

[0173] 1. Hydrorefining catalyst loading

[0174] The volume ratio of A4 and A3 catalysts is 8:2, and they are loaded into the adiabatic bed reactor in sequence.

[0175] 2. Sulfidation of Hydrorefining Catalyst

[0176] Under the system pressure of 0.5MPa, N2 was passed and the temperature was raised to 150℃ at a rate of 30℃ / h, the N2 flow rate was 850ml / min, N2 was stopped, H2 and sulfide oil (sulfide oil contains CS2 5000ppm) were passed, the system pressure was raised to 2.5MPa, the temperature was raised to 180℃ at 5℃ / h for 6h, kept constant for 5h, then raised to 230℃ at 5℃ / h for 10h, kept constant for 4h, then raised to 280℃ at 10℃ / h for 5h, kept constant for 4h, then raised to 320℃ at 10℃ / h for 4h, kept constant for 2h, and the catalyst sulfidation was completed. When sulfide oil was passed, the bed inlet temperature was reduced to the temperature required for the reaction.

[0177] 3. Hydrocracking catalyst loading and reduction

[0178] B1 was loaded into an adiabatic bed reactor, and the room temperature was heated to 500°C at a rate of 20°C / h, and the temperature was kept constant for 3 hours under hydrogen atmosphere at a hydrogen gas rate of 500 ml / min.

[0179] 4. Catalyst evaluation

[0180] (The inferior aromatic-rich distillate oil raw material comes from the product stream of cracking carbon nine, and the final distillation point is 275°C) The first-stage hydrogenation product (bromine value is 35mgBr / 100g oil, and gum is 40mg / 100g oil) is mixed with new hydrogen and second-stage circulating hydrogen, and is heated in the second-stage heater and then enters the second-stage hydrotreating reactor. The hydrotreating product enters the second-stage high-separation tank, the hydrogen enters the second-stage compressor for circulation, and the liquid phase enters the second-stage oil-water separation tank. The oil phase enters the second-stage low-separation tank and is removed from the oil phase by the gas stripping tower. The mass ratio of the second-stage purified product circulation volume to the second-stage liquid phase feed is 0.5. The second-stage purified product is heated by the third-stage heater and then enters the third-stage hydrocracking reactor. The cracking product enters the third-stage high-separation tank, the hydrogen enters the third-stage compressor for circulation, and the liquid phase enters the third-stage low-separation tank and then enters the distillation tower for distillation to obtain BTX and LPG.

[0181] The temperature of the hydrotreating feedstock stream is 225°C, and the temperature of the hydrotreating product stream is 258°C.

[0182] The evaluation results are listed in Table 1 (purification) and Table 2 (cracking).

[0183] Example 7

[0184] 1. Hydrorefining catalyst loading

[0185] Take commercially available NiMo catalyst C1, NiMoW catalyst C2, and CoMo catalyst C3, with the volume ratio of C1, C2, and C3 being 3:5:2, and load them into an adiabatic bed reactor in sequence.

[0186] 2. Sulfidation of Hydrorefining Catalyst

[0187] Under the system pressure of 0.5MPa, N2 was passed and the temperature was raised to 150℃ at a rate of 30℃ / h, the N2 flow rate was 850ml / min, N2 was stopped, H2 and sulfide oil (sulfide oil contains CS2 5000ppm) were passed, the system pressure was raised to 2.5MPa, the temperature was raised to 180℃ at 5℃ / h for 6h, kept constant for 5h, then raised to 230℃ at 5℃ / h for 10h, kept constant for 4h, then raised to 280℃ at 10℃ / h for 5h, kept constant for 4h, then raised to 320℃ at 10℃ / h for 4h, kept constant for 2h, and the catalyst sulfidation was completed. When sulfide oil was passed, the bed inlet temperature was reduced to the temperature required for the reaction.

[0188] 3. Hydrocracking catalyst loading and reduction

[0189] B1 was loaded into an adiabatic bed reactor, and the room temperature was heated to 500°C at a rate of 20°C / h, and the temperature was kept constant for 3 hours under hydrogen atmosphere at a hydrogen gas rate of 500 ml / min.

[0190] 4. Catalyst evaluation

[0191] (The inferior aromatic-rich distillate oil raw material comes from the product stream of ethylene tar, and the final distillation point is 300°C) the first-stage hydrogenation product (bromine value is 40mgBr / 100g oil, and gum is 48mg / 100g oil) is mixed with new hydrogen and second-stage circulating hydrogen, and heated in the second-stage heater and then enters the second-stage hydrotreating reactor. The hydrotreating product enters the second-stage high-separation tank, the hydrogen enters the second-stage compressor for circulation, the liquid phase enters the second-stage oil-water separation tank, the oil phase enters the second-stage low-separation tank and is removed from the oil phase by a gas stripping tower. The mass ratio of the second-stage purified product circulation volume to the second-stage liquid phase feed is 0.5. The second-stage purified product is heated by a third-stage heater and then enters a third-stage hydrocracking reactor. The cracking product enters the third-stage high-separation tank, the hydrogen enters the third-stage compressor for circulation, and the liquid phase enters the distillation tower after the third-stage low-separation tank for fractionation to obtain BTX and LPG.

[0192] The temperature of the hydrotreating feedstock stream is 225°C, and the temperature of the hydrotreating product stream is 278°C.

[0193] The evaluation results are listed in Table 1 (purification) and Table 2 (cracking).

[0194] Example 8

[0195] 1. Hydrorefining catalyst loading

[0196] Catalysts A4, A2 and A3 are sequentially loaded into the adiabatic bed reactor in a volume ratio of 3:5:2.

[0197] 2. Sulfidation of Hydrorefining Catalyst

[0198] Under the system pressure of 0.5MPa, N2 was passed and the temperature was raised to 150℃ at a rate of 30℃ / h, the N2 flow rate was 850ml / min, N2 was stopped, H2 and sulfide oil (sulfide oil contains CS2 5000ppm) were passed, the system pressure was raised to 2.5MPa, the temperature was raised to 180℃ at 5℃ / h for 6h, kept constant for 5h, then raised to 230℃ at 5℃ / h for 10h, kept constant for 4h, then raised to 280℃ at 10℃ / h for 5h, kept constant for 4h, then raised to 320℃ at 10℃ / h for 4h, kept constant for 2h, and the catalyst sulfidation was completed. When sulfide oil was passed, the bed inlet temperature was reduced to the temperature required for the reaction.

[0199] 3. Hydrocracking catalyst loading and reduction

[0200] B1 was loaded into an adiabatic bed reactor, and the room temperature was heated to 500°C at a rate of 20°C / h, and the temperature was kept constant for 3 hours under hydrogen atmosphere at a hydrogen gas rate of 500 ml / min.

[0201] 4. Catalyst evaluation

[0202] (The inferior aromatic-rich distillate oil raw material comes from the product stream of ethylene tar, and the final distillation point is 300°C) the first-stage hydrogenation product (bromine value is 40mgBr / 100g oil, and gum is 48mg / 100g oil) is mixed with new hydrogen and second-stage circulating hydrogen, and heated in the second-stage heater and then enters the second-stage hydrotreating reactor. The hydrotreating product enters the second-stage high-separation tank, the hydrogen enters the second-stage compressor for circulation, the liquid phase enters the second-stage oil-water separation tank, the oil phase enters the second-stage low-separation tank and is removed from the oil phase by a gas stripping tower. The mass ratio of the second-stage purified product circulation volume to the second-stage liquid phase feed is 2.1. The second-stage purified product is heated by a third-stage heater and then enters a third-stage hydrocracking reactor. The cracking product enters the third-stage high-separation tank, the hydrogen enters the third-stage compressor for circulation, and the liquid phase enters the distillation tower after the third-stage low-separation tank for distillation to obtain BTX and LPG.

[0203] The temperature of the hydrotreating feedstock stream is 225°C, and the temperature of the hydrotreating product stream is 247°C.

[0204] The evaluation results are listed in Table 1 (purification) and Table 2 (cracking).

[0205] Example 9

[0206] 1. Hydrorefining catalyst loading

[0207] Catalysts A4, A2 and A3 are sequentially loaded into the adiabatic bed reactor in a volume ratio of 3:5:2.

[0208] 2. Sulfidation of Hydrorefining Catalyst

[0209] Under the system pressure of 0.5MPa, N2 was passed and the temperature was raised to 150℃ at a rate of 30℃ / h, the N2 flow rate was 850ml / min, N2 was stopped, H2 and sulfide oil (sulfide oil contains CS2 5000ppm) were passed, the system pressure was raised to 2.5MPa, the temperature was raised to 180℃ at 5℃ / h for 6h, kept constant for 5h, then raised to 230℃ at 5℃ / h for 10h, kept constant for 4h, then raised to 280℃ at 10℃ / h for 5h, kept constant for 4h, then raised to 320℃ at 10℃ / h for 4h, kept constant for 2h, and the catalyst sulfidation was completed. When sulfide oil was passed, the bed inlet temperature was reduced to the temperature required for the reaction.

[0210] 3. Hydrocracking catalyst loading and reduction

[0211] B1 was loaded into an adiabatic bed reactor, and the room temperature was heated to 500°C at a rate of 20°C / h, and the temperature was kept constant for 3 hours under hydrogen atmosphere at a hydrogen gas rate of 500 ml / min.

[0212] 4. Catalyst evaluation

[0213] (The inferior aromatic-rich distillate oil raw material comes from the product stream of ethylene tar, and the final distillation point is 300°C) the first-stage hydrogenation product (bromine value is 40mgBr / 100g oil, and gum is 48mg / 100g oil) is mixed with new hydrogen and second-stage circulating hydrogen, and heated in the second-stage heater and then enters the second-stage hydrotreating reactor. The hydrotreating product enters the second-stage high-separation tank, the hydrogen enters the second-stage compressor for circulation, the liquid phase enters the second-stage oil-water separation tank, the oil phase enters the second-stage low-separation tank and is removed from the oil phase by a gas stripping tower. The mass ratio of the second-stage purified product circulation volume to the second-stage liquid phase feed is 0.1. The second-stage purified product is heated by a third-stage heater and then enters a third-stage hydrocracking reactor. The cracking product enters the third-stage high-separation tank, the hydrogen enters the third-stage compressor for circulation, and the liquid phase enters the distillation tower after the third-stage low-separation tank for fractionation to obtain BTX and LPG.

[0214] The temperature of the hydrotreating feedstock stream is 225°C, and the temperature of the hydrotreating product stream is 305°C.

[0215] The evaluation results are listed in Table 1 (purification) and Table 2 (cracking).

[0216] Example 10

[0217] 1. Preparation of Hydrorefining Catalyst

[0218] 1L of Al2O3-TiO2-SiO2 (pore size: 11nm, Al2O3 67%, TiO2 5%, SiO2 28%) was mixed with 0.75L of a mixed solution of nickel acetate, ammonium metatungstate, lanthanum nitrate, ammonium molybdate and phosphoric acid (containing NiO 40g, MoO3 60g, La2O3 5g, WO3165g, P2O5 10g), dried at 110°C for 6 hours, and calcined at 450°C for 4 hours to prepare catalyst A5. The TEM image was obtained after sulfurization. Figure 1 b. It can be seen from the figure that the sulfided catalyst is partially agglomerated and the chip length is 5-15nm.

[0219] A5 contains NiO 40g / L, MoO3 60g / L, La2O3 5g / L, WO3 165g / L, and P2O5 10g / L.

[0220] 2. Hydrorefining catalyst loading

[0221] Catalysts A4, A5 and A3 are sequentially loaded into the adiabatic bed reactor in a volume ratio of 3:5:2.

[0222] 3. Sulfidation of Hydrorefining Catalyst

[0223] Under the system pressure of 0.5MPa, N2 was passed and the temperature was raised to 150℃ at a rate of 30℃ / h, the N2 flow rate was 850ml / min, N2 was stopped, H2 and sulfide oil (sulfide oil contains CS2 5000ppm) were passed, the system pressure was raised to 2.5MPa, the temperature was raised to 180℃ at 5℃ / h for 6h, kept constant for 5h, then raised to 230℃ at 5℃ / h for 10h, kept constant for 4h, then raised to 280℃ at 10℃ / h for 5h, kept constant for 4h, then raised to 320℃ at 10℃ / h for 4h, kept constant for 2h, and the catalyst sulfidation was completed. When sulfide oil was passed, the bed inlet temperature was reduced to the temperature required for the reaction.

[0224] 4. Hydrocracking catalyst reduction

[0225] B1 was loaded into an adiabatic bed reactor, and the room temperature was heated to 500°C at a rate of 20°C / h, and the temperature was kept constant for 3 hours under hydrogen atmosphere at a hydrogen gas rate of 500 ml / min.

[0226] 5. Catalyst evaluation

[0227] (The inferior aromatic-rich distillate oil raw material comes from the product stream of ethylene tar, and the final distillation point is 300°C) the first-stage hydrogenation product (bromine value is 40mgBr / 100g oil, and gum is 48mg / 100g oil) is mixed with new hydrogen and second-stage circulating hydrogen, and heated in the second-stage heater and then enters the second-stage hydrotreating reactor. The hydrotreating product enters the second-stage high-separation tank, the hydrogen enters the second-stage compressor for circulation, the liquid phase enters the second-stage oil-water separation tank, the oil phase enters the second-stage low-separation tank and is removed from the oil phase by a gas stripping tower. The mass ratio of the second-stage purified product circulation volume to the second-stage liquid phase feed is 0.5. The second-stage purified product is heated by a third-stage heater and then enters a third-stage hydrocracking reactor. The cracking product enters the third-stage high-separation tank, the hydrogen enters the third-stage compressor for circulation, and the liquid phase enters the distillation tower after the third-stage low-separation tank for fractionation to obtain BTX and LPG.

[0228] The temperature of the hydrotreating feedstock stream is 225°C, and the temperature of the hydrotreating product stream is 275°C.

[0229] The evaluation results are listed in Table 1 (purification) and Table 2 (cracking).

[0230] Embodiment 11

[0231] 1. Hydrorefining catalyst loading

[0232] Take the A1 refined catalyst and load it into the adiabatic bed reactor.

[0233] 2. Sulfidation of Hydrorefining Catalyst

[0234] Under the system pressure of 0.5MPa, N2 was passed and the temperature was raised to 150℃ at a rate of 30℃ / h, the N2 flow rate was 850ml / min, N2 was stopped, H2 and sulfide oil (sulfide oil contains CS2 5000ppm) were passed, the system pressure was raised to 2.5MPa, the temperature was raised to 180℃ at 5℃ / h for 6h, kept constant for 5h, then raised to 230℃ at 5℃ / h for 10h, kept constant for 4h, then raised to 280℃ at 10℃ / h for 5h, kept constant for 4h, then raised to 320℃ at 10℃ / h for 4h, kept constant for 2h, and the catalyst sulfidation was completed. When sulfide oil was passed, the bed inlet temperature was reduced to the temperature required for the reaction.

[0235] 3. Hydrocracking catalyst reduction

[0236] B1 was loaded into an adiabatic bed reactor, and the room temperature was heated to 500°C at a rate of 20°C / h, and the temperature was kept constant for 3 hours under hydrogen atmosphere at a hydrogen gas rate of 500 ml / min.

[0237] 4. Catalyst evaluation

[0238] (The inferior aromatic-rich distillate oil raw material comes from the product stream of ethylene tar, and the final distillation point is 300°C) the first-stage hydrogenation product (bromine value is 40mgBr / 100g oil, and gum is 48mg / 100g oil) is mixed with new hydrogen and second-stage circulating hydrogen, and heated in the second-stage heater and then enters the second-stage hydrotreating reactor. The hydrotreating product enters the second-stage high-separation tank, the hydrogen enters the second-stage compressor for circulation, the liquid phase enters the second-stage oil-water separation tank, the oil phase enters the second-stage low-separation tank and is removed from the oil phase by a gas stripping tower. The mass ratio of the second-stage purified product circulation volume to the second-stage liquid phase feed is 0.5. The second-stage purified product is heated by a third-stage heater and then enters a third-stage hydrocracking reactor. The cracking product enters the third-stage high-separation tank, the hydrogen enters the third-stage compressor for circulation, and the liquid phase enters the distillation tower after the third-stage low-separation tank for fractionation to obtain BTX and LPG.

[0239] The temperature of the hydrotreating feedstock stream is 225°C, and the temperature of the hydrotreating product stream is 260°C.

[0240] The evaluation results are listed in Table 1 (purification) and Table 2 (cracking).

[0241] Example 12

[0242] 1. Hydrorefining catalyst loading

[0243] Catalysts A4, A2 and A3 are sequentially loaded into the adiabatic bed reactor in a volume ratio of 3:5:2.

[0244] 2. Sulfidation of Hydrorefining Catalyst

[0245] Under the system pressure of 0.5MPa, N2 was passed and the temperature was raised to 150℃ at a rate of 30℃ / h, the N2 flow rate was 850ml / min, N2 was stopped, H2 and sulfide oil (sulfide oil contains CS2 5000ppm) were passed, the system pressure was raised to 2.5MPa, the temperature was raised to 180℃ at 5℃ / h for 6h, kept constant for 5h, then raised to 230℃ at 5℃ / h for 10h, kept constant for 4h, then raised to 280℃ at 10℃ / h for 5h, kept constant for 4h, then raised to 320℃ at 10℃ / h for 4h, kept constant for 2h, and the catalyst sulfidation was completed. When sulfide oil was passed, the bed inlet temperature was reduced to the temperature required for the reaction.

[0246] 3. Hydrocracking catalyst loading and reduction

[0247] B1 was loaded into an adiabatic bed reactor, and the room temperature was heated to 500°C at a rate of 20°C / h, and the temperature was kept constant for 3 hours under hydrogen atmosphere at a hydrogen gas rate of 500 ml / min.

[0248] 4. Catalyst evaluation

[0249] (The inferior aromatic fraction oil raw material comes from the product stream of ethylene tar, and the final distillation point is 300°C) After the raw material is deweighted (the bromine value is 150mgBr / 100g oil, and the gum is 100mg / 100g oil), it is mixed with new hydrogen and second-stage circulating hydrogen, heated in the second-stage heater, and then entered into the second-stage hydrotreating reactor. The hydrotreating product enters the second-stage high-separation tank, the hydrogen enters the second-stage compressor for circulation, and the liquid phase enters the second-stage oil-water separation tank. The oil phase enters the second-stage low-separation tank and is removed from the oil phase by the gas stripping tower. The mass ratio of the second-stage purified product circulation volume to the second-stage liquid phase feed is 0.5. The second-stage purified product is heated by the third-stage heater and then enters the third-stage hydrocracking reactor. The cracking product enters the third-stage high-separation tank, the hydrogen enters the third-stage compressor for circulation, and the liquid phase enters the distillation tower after the third-stage low-separation tank for fractionation to obtain BTX and LPG.

[0250] The temperature of the hydrotreating feedstock stream is 225°C, and the temperature of the hydrotreating product stream is 317°C.

[0251] The evaluation results are listed in Table 1 (purification) and Table 2 (cracking).

[0252] Table 1

[0253]

[0254]

[0255] Table 2

[0256] Benzene / wt% Toluene / wt% Xylene / wt% BTX / wt% Ni dispersion / % Catalyst TPR / ℃ Example 1 12.9 28.8 18.7 60.4 14.8 386 Example 2 13.8 30.3 19.3 63.4 14.8 386 Example 3 13.6 27.8 18.9 60.3 14.8 386 Example 4 15.2 28.1 19.4 62.7 14.8 386 Example 5 15.1 26.3 15.9 57.3 14.8 386 Example 6 15.3 28.0 18.5 61.8 14.8 386 Example 7 12.1 25.6 16.5 54.2 14.8 386 Example 8 10.1 23.8 15.2 49.1 14.8 386 Example 9 11.2 22.8 16.1 50.1 14.8 386 Example 10 14.1 18.6 12.8 45.5 3.1 386 Embodiment 11 9.4 24.6 11.3 45.3 14.8 386 Example 12 11.8 20.2 13.9 45.9 14.8 386

[0257] The preferred embodiments of the present invention are described in detail above, but the present invention is not limited thereto. Within the technical concept of the present invention, the technical solution of the present invention can be subjected to a variety of simple modifications, including the combination of various technical features in any other suitable manner, and these simple modifications and combinations should also be regarded as the contents disclosed by the present invention and belong to the protection scope of the present invention.

Claims

1. A method for utilizing inferior aromatic-rich distillate oil, characterized in that: The method comprises the following steps: subjecting inferior aromatic-rich fraction oil raw materials to first-stage hydrogenation conversion of dienes and reduction of bromine value, second-stage hydrogenation refining, first purification, third-stage hydrocracking, and second purification, and separating and obtaining BTX and LPG.

2. The method according to claim 1, wherein: The temperature of the product stream of the second-stage hydrotreating is higher than the temperature of the feed stream before the second-stage hydrotreating, and the temperature difference is not greater than 75°C, preferably 30-75°C, more preferably 35-60°C.

3. The method according to claim 1 or 2, wherein: The temperature of the hydrofinished product stream is 250-310°C, preferably 260-300°C; and / or The temperature of the feedstock stream before hydrofining is 200-260°C, preferably 210-250°C.

4. The method according to any one of claims 1 to 3, wherein: The final distillation point of the inferior aromatic-rich distillate oil raw material is ≯300°C; Preferably, the inferior aromatic-rich fraction oil feedstock is selected from at least one of a cracking C9 product stream, a cracking ethylene tar product stream and a cracking coal tar product stream.

5. The method according to any one of claims 1 to 4, wherein: The feedstock stream before the second-stage hydrofining comprises a first purified stream; Preferably, the mass content of the first purified stream in the feed stream before the second-stage hydrofining is 0-90%, preferably 0-50%.

6. The method according to any one of claims 1 to 5, wherein: The bromine value of the product stream after the first stage hydrogenation is ≯40 mgBr / 100 g oil, and the colloid is ≯50 mg / 100 g oil; and / or The first purified product stream has a sulfur content of less than 1 ppm, preferably less than 0.85 ppm; and / or a nitrogen content of less than 1 ppm, preferably less than 0.8 ppm.

7. The method according to any one of claims 1 to 6, wherein: The two-stage hydrotreating catalyst comprises a first active component, an active promoter and a first carrier, wherein the first active component comprises a Group VIB metal and a Group VIII metal, preferably at least one of Ni, Co, Fe, Pt and Pd, and Mo and / or W; the active promoter comprises a lanthanide metal and / or a Group VA element, preferably La, P, Ce; After the two-stage hydrofining catalyst is presulfurized, the number of stacking layers of the sulfide active phase in the catalyst is 3-4 layers and the structure is tower-shaped, and preferably the stacking length of the sulfide active phase is 3-6 nm; Preferably, The first carrier in the two-stage hydrotreating catalyst is selected from at least one of Al2O3, Al2O3-TiO2, Al2O3-TiO2-SiO2, and Al2O3-SiO2, and is further preferably selected from one or two of Al2O3, Al2O3-TiO2, and Al2O3-TiO2-SiO2; and / or The average pore diameter of the first carrier in the two-stage hydrorefining catalyst is 7-20 nm; and / or the pore volume is 0.5-0.98 cm 3 / g; and / or specific surface area of ​​190-380m 2 / g; More preferably, In the two-stage hydrorefining catalyst, the density of the first active component in terms of oxide is 50-300 g / L, preferably 100-300 g / L; the density of the active promoter in terms of oxide is 5-50 g / L, preferably 5-40 g / L.

8. The method according to any one of claims 1 to 7, wherein: The preparation method of the catalyst for two-stage hydrofining comprises: mixing a first active component source, an active auxiliary agent source, a first carrier source and an additive to obtain a mixed liquid, drying and calcining; The concentration of the Group VIB metal in the mixed solution as oxide is 150-850 g / L; the concentration of the Group VIII metal as oxide is 50-300 g / L; The additive is a mixture of an amino compound, an organic acid metal complexing agent and an alcohol; Preferably, The additive is selected from at least one of urea, acetamide, ethylenediamine, ethanolamine, diethanolamine, triethanolamine, and at least one of 1,2-cyclohexanediaminetetraacetic acid, citric acid, tartaric acid, ethylenediaminetetraacetic acid, and ethylene glycol and / or ethanol; and / or The mass fraction of the additive in the mixed solution is 0.01-10%, more preferably 1.5-7.5%; More preferably, The mass ratio of the amino compound, the organic acid metal complexing agent and the alcohol in the additive is 1-2:1-2:1-2.

9. The method according to any one of claims 1 to 8, wherein: The three-stage hydrocracking catalyst comprises a second active component and a second carrier, wherein the second active component comprises a Group VIII metal, preferably Ni; the second carrier comprises a layered HZSM-5 / MCM-41 molecular sieve and optionally Ce oxide and Zr oxide; Preferably, the interlayer thickness of the layered ZSM-5 / MCM-41 molecular sieve is 5-10 nm, preferably 5-7 nm; and / or The pore size of the layered ZSM-5 / MCM-41 molecular sieve is 3-18 nm, preferably 3-11 nm; and / or The medium-strong acid content / total acid content of the layered ZSM-5 / MCM-41 molecular sieve is 0.55-0.85, preferably 0.55-0.65; and / or The specific surface area of ​​layered ZSM-5 / MCM-41 molecular sieve is ≮400m 2 ·g -1 , preferably 400-550m 2 ·g -1 ; and / or The SiO2 / Al2O3 molar ratio of the layered ZSM-5 / MCM-41 molecular sieve is 20-200, preferably 20-60; More preferably, In the three-stage hydrocracking catalyst, the content of the second active component in terms of oxide is 50-260 g / L, preferably 50-150 g / L; the content of Ce oxide and Zr oxide in terms of oxide is 0.1-100 g / L, preferably 0.1-50 g / L; More preferably, The TPR hydrogen atmosphere reduction temperature of the three-stage hydrocracking catalyst is lower than 400° C., and more preferably lower than 390° C.; and / or The dispersion degree of the second active component element in the three-stage hydrocracking catalyst is greater than 8%, preferably 8-15%.

10. The method according to any one of claims 1 to 9, wherein: The preparation method of the three-stage hydrocracking catalyst comprises: mixing, drying and calcining a second active component source, a second carrier source and a chelating surfactant; Preferably, the chelating surfactant is selected from at least one of 1,2-cyclohexanediaminetetraacetic acid, citric acid, tartaric acid, and ethylenediaminetetraacetic acid, preferably one of 1,2-cyclohexanediaminetetraacetic acid, citric acid, and ethylenediaminetetraacetic acid; more preferably, the amount of the chelating surfactant in the mixed material is 0.01-10%, and further preferably 1.5-5%.

11. The method according to any one of claims 1 to 10, wherein: The conditions of the two-stage hydrofining include: temperature of 200-270°C; and / or pressure of 2.0-8.0 MPa; and / or hydrogen-to-oil volume ratio of 400-3000; and / or liquid phase volume space velocity of 0.2-3h -1 ; and / or The conditions of the three-stage hydrocracking include: pressure of 2-8 MPa; and / or space velocity of 0.2-3 h -1 ; and / or temperature is 260-500°C; and / or hydrogen to oil volume ratio is 400-3000. Preferably, The conditions of the two-stage hydrofining include: temperature of 210-250°C; and / or pressure of 2.2-6.0 MPa; and / or hydrogen-oil volume ratio of 600-2000; and / or liquid phase volume space velocity of 0.6-1.2 h -1 ; The conditions of the three-stage hydrocracking include: pressure of 3.0-6.5 MPa; and / or space velocity of 0.6-1.2 h -1 ; and / or temperature is 260-480°C; and / or hydrogen to oil volume ratio is 600-2000.

12. The method according to any one of claims 1 to 11, wherein: The two-stage hydrofining method is that the material passes through at least two layers of catalyst along the flow direction, and the pore size of the first carrier in the catalyst decreases successively.

13. A device for utilizing low-quality aromatic-rich distillate oil used in the method according to any one of claims 1 to 12, characterized in that: The device comprises a first stage hydrogenation device, a second stage hydrofining device, a first purification device, a third stage hydrocracking device, and a second purification device arranged in series; wherein, The first stage hydrogenation unit is used to convert dienes in the inferior aromatic-rich distillate oil feedstock and reduce the bromine value to obtain a first stage hydrogenation product stream; The temperature of the product stream outlet of the two-stage hydrotreating device is higher than the temperature of the raw material inlet, and the temperature difference is not greater than 75° C., and is used to purify the first-stage hydrogenation product stream, so that the naphthalene series products are highly selectively converted into tetrahydronaphthalene series products and the sulfur and nitrogen content of the refined product is less than 1 ppm, thereby obtaining a two-stage hydrotreating product stream; The first purification device is used to remove hydrogen sulfide from the second-stage hydrorefining product stream to obtain a monocyclic aromatic hydrocarbon mixture; The three-stage hydrocracking unit is used to crack the monocyclic aromatic hydrocarbon mixture, so that the tetralin series is hydrocracking to generate BTX and LPG, and obtain a three-stage hydrocracking product stream; The second purification device is used to purify and separate the three-stage hydrocracking product stream to obtain BTX and LPG; Preferably, The monocyclic aromatic hydrocarbon mixture discharge port of the first purification device is provided with a circulation pipeline connected to the raw material inlet of the second-stage hydrotreating device.

14. The device according to claim 13, wherein: The first purification device and the second purification device are independently selected from at least one of a high separation tank, an oil-water separation tank, a gas stripping tower, a low separation tank and a distillation tower; and / or The two-stage hydrogenation device is loaded with a two-stage hydrogenation refining catalyst, the two-stage hydrogenation refining catalyst comprises a first active component, an active promoter and a first carrier, the first active component comprises a Group VIB metal and a Group VIII metal, preferably comprises at least one of Ni, Co, Fe, Pt and Pd, and Mo and / or W; the active promoter comprises a lanthanide metal and / or a Group VA element, preferably La, P, Ce; after the two-stage hydrogenation refining catalyst is presulfurized, the number of stacking layers of the sulfide active phase in the catalyst is 3-4 layers and presents a tower-shaped structure, preferably the stacking length of the sulfide active phase is 3-6nm, and the active sites of the edges, corners and sides are fully exposed; The three-stage hydrocracking unit is loaded with a three-stage hydrocracking catalyst, which includes a second active component and a second carrier, wherein the second active component contains a Group VIII metal, preferably Ni; and the second carrier contains a layered HZSM-5 / MCM-41 molecular sieve and optional Ce oxide and Zr oxide.

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