Methods and apparatus for utilizing inferior aromatic distillate oils

Abstract

This invention relates to the field of ethylene by-product utilization, specifically to a method and apparatus for utilizing low-quality aromatic-rich distillate oil. The method includes sequentially subjecting the low-quality aromatic-rich distillate oil feedstock to a first-stage hydrogenation process to convert to dienes and reduce bromine value, a second-stage hydrogenation refining process, a first purification process, a third-stage hydrocracking process, and a second purification process to separate BTX and LPG. By using this method to sequentially process the low-quality aromatic-rich distillate oil feedstock to a first-stage hydrogenation process to convert to dienes and reduce bromine value, a second-stage hydrogenation refining process, a first purification process, a third-stage hydrocracking process, and a second purification process, the low-quality aromatic-rich distillate oil can be fully utilized, increasing its added value. While achieving chemical-type utilization of ethylene cracking by-products and enhancing the competitiveness of steam cracking to produce ethylene, it can also adjust and improve the quality of refined products, increase the yield of BTX, and extend the life of cracking catalysts.

Description

Technical Field

[0001] This invention relates to the field of ethylene by-product utilization, specifically to a method and apparatus for utilizing inferior aromatic-rich distillate oil. Background Technology

[0002] Low-quality aromatic-rich distillate oils are mainly byproducts of industrial cracking, including cracked C9 and ethylene tar fractions. Cracked C9 primarily originates from the cracked gasoline C9 fraction separated after passing through a BTX tower, with an aromatic content exceeding 70%, accounting for 11%-22% of ethylene production. Domestically, the vast majority of cracked C9 is sold only as inexpensive primary fuel oil or after preliminary processing.

[0003] Ethylene tar, also known as cracked tar, is a high-boiling-point liquid product of the steam cracking process for producing ethylene. It belongs to the diesel fraction (205-360℃) and mainly originates from the bottom of quench oil towers and heavy fuel oil stripping towers. Ethylene tar is a heavy distillate oil rich in aromatics, primarily containing monocyclic heavy aromatics, polycyclic or fused-ring aromatics. It has a complex composition, is prone to polymerization, and has high levels of gum, heavy metals, and ash, making it unsuitable for direct use. The yield of ethylene tar varies depending on the cracking feedstock, generally accounting for about 1 / 5 of the ethylene production. With the increasing use of heavier feedstocks in ethylene production, its yield shows an increasing trend.

[0004] The main uses of ethylene tar include blending gasoline and diesel components, fuel, carbon black production, extraction of naphthalene and methylnaphthalene, and production of aromatic solvent oils. Among these, using it as fuel is the primary utilization of ethylene tar in China, but this method is not only economically inefficient but also produces black smoke during combustion, causing environmental pollution. Carbon black production is the main method of utilizing ethylene tar abroad, offering considerable economic benefits, but domestic technology relies heavily on imports. Domestically developed processes for recovering naphthalene have relatively low yields, and the production of methylnaphthalene suffers from limitations such as small scale, high energy consumption, and low product purity. Hydrogenation of ethylene tar can produce high-value-added BTX aromatics, which can significantly improve the utilization rate of ethylene tar and contribute to the value-added utilization of ethylene byproducts, showing promising market prospects. Foreign companies have already begun using cracked fuel oil to produce aromatic solvent oils, with major producers including ExxonMobil (USA), Shell (Netherlands), and Maruzen Oil Company (Japan). Overall, ethylene tar is mainly used for low-value fuel applications. However, with increasingly stringent environmental regulations, the use of untreated, sulfur- and nitrogen-rich, and highly unsaturated ethylene tar fuel will be increasingly restricted, and its future prospects are uncertain.

[0005] Ethylene tar has a high yield (approximately 70%) in the fractions between 205℃ and 300℃, followed by gum and asphaltenes. Ethylene tar also has a high sulfur content, high polycyclic aromatic hydrocarbon content, high density, and short side chains in its aromatic compounds. The main components of the fraction from the initial boiling point to 205℃ are indene and its homologues; the fraction from 205℃ to 225℃ is naphthalene; the fraction from 225℃ to 245℃ is mainly methylnaphthalene; the fraction from 245℃ to 300℃ is mainly dimethylnaphthalene; the fraction from 300℃ to 360℃ contains large amounts of anthracene, acenaphthene, and phenanthrene; and the substances above 360℃ are mainly gums and asphaltenes with a high carbon-to-hydrogen ratio. Therefore, all fractions of ethylene tar are important raw materials for chemical organic synthesis, from which many valuable chemical products can be extracted, demonstrating significant utilization value.

[0006] In the field of heavy distillate hydrotreating, catalytic cracking feedstock hydrotreating technology has been industrially applied since the 1970s, and has been used in many refineries processing sulfur-containing or high-sulfur crude oil. Currently, mature pretreatment technologies for catalytic cracking feedstocks exist both domestically and internationally, primarily including: UOP's VGO Unionfining and APCU (partial conversion hydrocracking) technologies, and Haldor's... The company's Aroshift technology, Chevron's VGO Hydrotreating technology, Exxon's VGO Hydrodesulfurization technology, IFP's T-star technology, and Mobil, AKZO, and Kellogg's MAKfinging technology, among others, are all examples. Except for T-star, which uses a fluidized bed reaction process, most FCC feedstock hydrotreating pretreatment units employ a fixed bed process. To further improve product quality and conversion rates, catalytic feedstock hydrotreating pretreatment processes are gradually shifting from traditional hydrodesulfurization (HDS) to moderate hydrocracking (MHC) to enhance denitrification, residual carbon, and polycyclic aromatic hydrocarbon saturation capabilities. This approach offers greater operational flexibility and more significant economic benefits.

[0007] CN101724458A discloses a method for hydrogenating ethylene tar. The method involves fractionating ethylene tar into light and heavy components via a fractionation tower. The light fraction is then passed sequentially through a hydrogenation protection catalyst, a hydrogenation refining catalyst, a light fraction hydrogenation decarbonization catalyst, and a hydrocracking catalyst. The heavy components are passed sequentially through a hydrogenation protection catalyst, a hydrogenation decarbonization catalyst, and a hydroconversion catalyst, ultimately yielding gasoline and diesel fractions.

[0008] CN112745949A discloses a method and system for the combined processing of de-oiled bitumen and aromatic-rich distillate oil. In this method, heavy oil is de-asphalted using a solvent. The de-asphalted oil then passes through a hydrotreating reactor I and enters a DCC unit to produce propylene, LCO, and HCO, etc. After the de-oiled bitumen passes through a hydrotreating reactor II, the product is fractionated. The light components enter a hydrorefining reactor to obtain gasoline and diesel fuel components, while the heavy components enter a delayed coking unit to produce coking gasoline, coking diesel, and coking wax oil, etc.

[0009] US5300212A discloses a process for hydrorefining inferior heavy oil. This method involves the conversion of heavy oil feedstock, hydrogen, and catalyst in two reactors. Specifically, the feedstock and a dispersed catalyst with molybdenum phosphate as a precursor first enter a first slurry-bed hydrorefining reactor, where a conversion reaction occurs at 343–482°C and 0.345–34.5 MPa. The reaction products, after separation, enter a second fluidized bed hydrorefining reactor, where conversion occurs at 343–399°C and 5.5–27.6 MPa under the action of a supported catalyst. The reaction products then enter a distillation column, yielding a <524°C fraction and a >524°C fraction. The <524°C fraction is used as the product, while the >524°C heavy fraction is recycled back to the second reactor. This process can refining inferior heavy oil. However, because the first reactor uses a dispersed catalyst and the second reactor uses a supported catalyst, catalyst particles carried out from the first reactor can easily clog the pores of the heavily supported catalyst in the second reactor or cover the active sites of the catalyst, causing catalyst deactivation and affecting the overall operating cycle.

[0010] CN102134500B discloses a method for extracting naphthalene, 1-methylnaphthalene, and 2-methylnaphthalene from ethylene tar. The method uses ethylene tar containing at least 30% naphthalene and methylnaphthalene as raw material, employing five distillation columns connected in series to sequentially separate naphthalene (>95%), tetramethylnaphthalene, 2-methylnaphthalene (98%), and 1-methylnaphthalene (98%).

[0011] CN109679011B discloses a method for producing copolymerized petroleum resin. This method uses ethylene tar through vacuum distillation and rectification to remove naphthalene, yielding C9-C6 resin. 10 The distillate is mixed with high-purity C5 resin, cresol, and styrene (after removing cyclopentadiene), and polymerized using BF3 as a catalyst. The mixture is then subjected to vacuum distillation, and the distillate is cooled to obtain a copolymerized petroleum resin. This method allows for flexible adjustment of the molecular weight of the petroleum resin by varying the amounts of cresol and styrene, thus enabling the production of petroleum resins with different properties to meet market demands.

[0012] CN113755211A discloses a method for producing needle coke using optimized ethylene tar feedstock. The method involves hydrocracking aromatics, pitch, and gums from ethylene tar. After high-phase, low-phase, and gas-liquid separation, the oil phase is fed into a coking tower to generate needle coke, while the gas phase is recycled to a hydrotreating tower or used as fuel.

[0013] CN1970688B discloses a comprehensive processing technology for ethylene tar. In this process, after the ethylene tar passes through a pre-fractionation tower, the light components are fed into a hydrorefining tower at 260-280°C. The refined products are then fed into four distillation towers connected in series to obtain solvent oil I, refined naphthalene, β-methylnaphthalene, α-methylnaphthalene, mixed methylnaphthalene, and solvent oil II, respectively.

[0014] US20080083649A1 discloses an upgraded utilization of tar. In this invention, ethylene tar is processed in a vacuum tube furnace. The top of the furnace is deasphalted to produce fuel gas, while the bottom of the furnace contains asphaltenes to produce syngas. A portion of the syngas is also fed into a coking unit to produce coking gasoline and diesel.

[0015] A comprehensive analysis of the above technologies reveals that existing technologies involve the production of gasoline and diesel fractions, petroleum coke, resins, and solvent oils from inferior aromatic oils, thus wasting valuable aromatic resources in the aromatic fraction oils.

[0016] Based on technologies such as selective hydrorefining and hydrocracking of aromatic distillate oils, optimization and innovation are carried out to maximize the production of benzene (B), toluene (T) and xylene (X). This can make full use of inferior aromatic distillate oils such as ethylene tar and cracked 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 this invention is to overcome the problems of low added value and serious environmental pollution caused by using inferior aromatic distillate oil as refinery fuel or in the production of solvent oil, gasoline and diesel blending components, etc. in the existing technology. This invention provides a method and apparatus for utilizing inferior aromatic distillate oil. This method can make full use of inferior aromatic distillate oil, increase its added value, realize the chemical utilization of ethylene cracking by-products, enhance the competitiveness of steam cracking to produce ethylene, and expand new pathways for the production of BTX and LPG.

[0018] To achieve the above objectives, the present invention provides a method for utilizing inferior aromatic-rich distillate oil, which includes sequentially subjecting the inferior aromatic-rich distillate oil feedstock to a first-stage hydrogenation to convert dienes and reduce bromine value, a second-stage hydrogenation refining, a first purification, a third-stage hydrocracking, and a second purification to separate BTX and LPG.

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

[0020] The first-stage hydrogenation unit is used to remove dienes from inferior aromatic-rich distillate 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 hydrorefining unit is higher than that of the raw material inlet, and the temperature difference is not greater than 75°C. It is used to refine the product stream of the first-stage hydrorefining unit, so that the naphthalene series compounds are selectively generated into tetrahydronaphthalene series compounds and the sulfur and nitrogen of the refined product are less than 1 ppm, thus obtaining the product stream of the two-stage hydrorefining unit.

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

[0023] The three-stage hydrocracking unit is used to crack a mixture of monocyclic aromatic hydrocarbons, so that tetrahydronaphthalene compounds are hydrocracking to generate BTX and LPG, and a three-stage hydrocracking product stream is obtained.

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

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

[0026] The method of this invention involves sequentially processing inferior aromatic-rich distillate feedstock through a first-stage hydrogenation to convert it into dienes and reduce its bromine value, a second-stage hydrogenation refining, a first purification, a third-stage hydrocracking, and a second purification. This fully utilizes the inferior aromatic-rich distillate feedstock, increases its added value, realizes the chemical utilization of ethylene cracking by-products, enhances the competitiveness of steam cracking to produce ethylene, while also regulating and improving the quality of refined products, increasing the yield of BTX, and extending the life of cracking catalysts. Attached Figure Description

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

[0028] Figure 2 The image shows the XRD pattern of the B1-supported HZSM-5 / MCM-41 molecular sieve.

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

[0030] Figure 4 This is a simplified flowchart of the utilization process of inferior aromatic distillate oil. Detailed Implementation

[0031] The endpoints and any values ​​of the ranges disclosed herein 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 the various ranges, the endpoint values ​​of the various ranges and individual point values, and individual point values ​​can be combined with each other to obtain one or more new numerical ranges, which should be considered as specifically disclosed herein.

[0032] This invention provides a method for utilizing inferior aromatic-rich distillate oil. The method includes sequentially subjecting the inferior aromatic-rich distillate oil feedstock to a first-stage hydrogenation to convert dienes and reduce bromine value, a second-stage hydrogenation refining, a first purification, a third-stage hydrocracking, and a second purification to separate BTX and LPG.

[0033] The method of this invention involves sequentially processing inferior aromatic-rich distillate feedstock through a first-stage hydrogenation to convert it into dienes and reduce its bromine value, a second-stage hydrogenation refining, a first purification, a third-stage hydrocracking, and a second purification. This fully utilizes the inferior aromatic-rich distillate feedstock, increases its added value, realizes the chemical utilization of ethylene cracking by-products, enhances the competitiveness of steam cracking to produce ethylene, and at the same time, can adjust and improve the quality rate of refined products, increase the yield of BTX, and extend the life of cracking catalysts.

[0034] One of the key factors for maximizing the selectivity of tetrahydronaphthalene derivatives is the reaction temperature; lower temperatures are more favorable for preparing the desired tetrahydronaphthalene derivatives. However, the hydrogenation of aromatic-rich distillate oils is a strongly exothermic reaction, therefore limiting the increase in the reaction temperature for the hydrogenation refining of aromatic-rich distillate oils is crucial.

[0035] According to a preferred embodiment of the present invention, the temperature of the product stream from the second-stage hydrorefining is higher than the temperature of the feed stream before the second-stage hydrorefining, and the temperature difference is no greater than 75°C. By adopting the aforementioned preferred solution, the quality of the refined product can be further improved, the yield of BTX product can be increased, and the lifespan of the cracking catalyst can be extended.

[0036] In this invention, as long as the temperature of the product stream from the 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 objective of this invention can be achieved. According to a preferred embodiment of this invention, in order to further improve the quality of the refined product, increase the yield of 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 hydrorefined product stream is 250-310°C, preferably 260-300°C. By adopting the aforementioned preferred embodiment, the quality of the refined product can be further improved, the yield of BTX product can be increased, and the lifespan of the cracking catalyst can be extended.

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

[0039] In this invention, as long as the objective of the invention can be achieved, there is no particular limitation on the final boiling point of the inferior aromatic-rich distillate feedstock. According to a preferred embodiment of the invention, the final boiling point of the inferior aromatic-rich distillate feedstock is ≤300℃. By adopting the aforementioned preferred solution, the quality of the refined product can be further improved, the yield of BTX product can be increased, and the life of the cracking catalyst can be extended.

[0040] In this invention, any low-quality aromatic-rich distillate feedstock can be used. According to a preferred embodiment of the invention, the low-quality aromatic-rich distillate feedstock is selected from at least one of the following: cracked C9 product stream, cracked ethylene tar product stream, and cracked coal tar product stream. By adopting the aforementioned preferred embodiment, the quality of the refined product can be further improved, the yield of BTX product can be increased, and the cracking catalyst lifetime can be extended.

[0041] According to a preferred embodiment of the present invention, the feed stream before the two-stage hydrorefining includes a first purified feed stream. By adopting the aforementioned preferred scheme, the quality of the refined product can be further improved, the yield of BTX product can be increased, and the cracking catalyst lifetime 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 two-stage hydrogenation refining is 0-90%, preferably 0-50%. By adopting the aforementioned preferred embodiment, 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 of hydrogenation is ≤40 mgBr / 100g oil, and the gum content is ≤50 mg / 100g oil. By adopting the aforementioned preferred scheme, the quality of the refined product can be further improved, the yield of product BTX can be increased, and the cracking catalyst life 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 aforementioned preferred embodiment, the quality of the refined product can be further improved, the yield of product BTX can be increased, and the lifespan 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 aforementioned preferred embodiment, the quality of the purified product can be further improved, the yield of product BTX can be increased, and the lifespan of the cracking catalyst can be extended.

[0046] In this invention, the catalyst for two-stage hydrorefining can be a conventional choice in the art. According to a preferred embodiment of the invention, the catalyst for two-stage hydrorefining comprises a first active component, an active promoter, and a first support. The first active component comprises Group VIB and Group VIII metals, preferably at least one of Ni, Co, Fe, Pt, and Pd, and Mo and / or W. The active promoter comprises lanthanide metals and / or Group VA elements, preferably La, P, and Ce. After pre-sulfurization, the catalyst for two-stage hydrorefining has 3-4 layers of sulfide active phase stacked in a tower-like structure, preferably with a stack length of 3-6 nm. Catalysts with this characteristic have fully exposed active sites such as edges, corners, and sides, which can further improve the quality of the refined product and increase the yield of BTX and LPG.

[0047] In this invention, the pre-vulcanization conditions can be conventionally chosen in the art. According to a preferred embodiment of this invention, the pre-vulcanization conditions include: a pressure of 0.4-0.6 MPa, heating to 130-170°C at a rate of 20-40°C / h with N2, an N2 flow rate of 500-1000 ml / min, stopping the N2 flow, and introducing H2 and vulcanized oil (containing CS2). (1000-10000ppm), the system pressure is increased to 2-3MPa, the temperature is increased at 5-10℃ / h for 3-6h to 180-200℃, held at the temperature for 4-6h, then increased at 5-10℃ / h for 5-10h to 220-240℃, held at the temperature for 2-5h, then increased at 5-10℃ / h for 5-10h to 270-300℃, held at the temperature for 3-5h, then increased at 5-10℃ / h for 4-8h to 320-340℃, and held at the temperature for 2-3h. The advantages of this invention are illustrated by the following example: The system is heated to 150°C at a rate of 30°C / h under a pressure of 0.5 MPa, with N2 flow rate of 850 ml / min. The N2 flow is then stopped, and H2 and sulfurized oil (containing 5000 ppm CS2) are introduced. The system pressure is increased to 2.5 MPa, and the temperature is raised to 180°C at a rate of 5°C / h for 6 hours. After holding this temperature for 5 hours, the temperature is raised to 230°C at a rate of 5°C / h for 10 hours. After holding this temperature for 4 hours, the temperature is raised to 280°C at a rate of 10°C / h for 5 hours. After holding this temperature for 4 hours, the temperature is raised to 320°C at a rate of 10°C / h for 4 hours. This is followed by holding this temperature for 2 hours.

[0048] According to a preferred embodiment of the present invention, the first support 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 more 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 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 support in the two-stage hydrorefining catalyst is 7-20 nm. By adopting the aforementioned preferred embodiment, the quality of the refined product can be further improved, the yield of BTX product can be increased, and the lifespan of the cracking catalyst can be extended.

[0050] According to a preferred embodiment of the present invention, the pore volume of the first support in the two-stage hydrorefining catalyst is 0.5-0.98 cm³. 3 / g. By adopting the aforementioned preferred scheme, the quality of the refined product can be further improved, the yield of BTX product 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 support in the two-stage hydrorefining catalyst is 190-380 m². 2 / g; By adopting the aforementioned preferred scheme, the quality of refined products can be further improved, the yield of product BTX can be increased, and the life of cracking catalyst can be extended.

[0052] In this invention, there are no particular limitations on the amount of each component in the catalyst for two-stage hydrorefining. According to a preferred embodiment of the invention, in the catalyst for two-stage hydrorefining, the first active component, calculated as oxide, is 50-300 g / L, preferably 100-300 g / L; the active auxiliary agent, calculated as oxide, is 5-50 g / L, preferably 5-40 g / L. By adopting the aforementioned preferred scheme, the quality of the refined product can be further improved, the yield of BTX product can be increased, and the life of the cracking catalyst can be extended.

[0053] In this invention, any method that can prepare the aforementioned hydrorefining catalyst is acceptable, and the preparation method can be a conventional choice in the art. According to a preferred embodiment of this invention, the preparation method of the two-stage hydrorefining catalyst includes: mixing a first active component source, an active auxiliary agent source, a first support source, and an additive to obtain a mixed solution, drying, and calcining; wherein the concentration of Group VIB metals in the mixed solution, calculated as oxides, is 150-850 g / L; the concentration of Group VIII metals, calculated as oxides, is 50-300 g / L; and the additive is a mixture of amino compounds, organic acid metal complexing agents, and alcohols. By adopting the aforementioned preferred scheme, the quality of the refined product can be further improved, the yield of BTX product can be increased, and the cracking catalyst lifetime 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, and triethanolamine, and at least one of 1,2-cyclohexanediaminetetraacetic acid, citric acid, tartaric acid, and ethylenediaminetetraacetic acid, as well as ethylene glycol and / or ethanol. By adopting the aforementioned preferred embodiment, the quality of the refined product can be further improved, the yield of BTX product can be increased, and the cracking catalyst lifetime can be extended.

[0055] According to a preferred embodiment of the present invention, the mass fraction of the additive in the mixture is 0.01-10%, more preferably 1.5-7.5%. By adopting the aforementioned preferred embodiment, the quality of the refined product can be further improved, the yield of BTX product 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 this invention, the catalyst for three-stage hydrocracking can be a conventional choice in the art. According to a preferred embodiment of the invention, the catalyst for three-stage hydrocracking includes a second active component and a second support. The second active component comprises a Group VIII metal, preferably Ni; the second support comprises layered HZSM-5 / MCM-41 molecular sieve and optionally Ce oxide and Zr oxide. By adopting the aforementioned preferred scheme, the yield of product BTX is improved and the lifetime 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 aforementioned preferred embodiment, the yield of product BTX is improved and the lifetime 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 aforementioned preferred embodiment, the yield of product BTX is improved and the lifetime of the cracking catalyst is extended.

[0060] According to a preferred embodiment of the present invention, the ratio of medium-strong acid content to 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 aforementioned preferred embodiment, the yield of product BTX is improved and the lifetime 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 ≥400 m². 2 ·g -1 Preferably 400-550m 2 ·g -1 By adopting the aforementioned preferred scheme, the yield of product BTX is improved and the lifespan 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 aforementioned preferred embodiment, the yield of product BTX is improved and the lifetime of the cracking catalyst is extended.

[0063] In this invention, the amount of each component in the three-stage hydrocracking catalyst is not particularly limited. According to a preferred embodiment of the invention, the content of the second active component (calculated as oxide) in the three-stage hydrocracking catalyst is 50-260 g / L, preferably 50-150 g / L; the content of Ce oxide and Zr oxide (calculated as oxide) is 0.1-100 g / L, preferably 0.1-50 g / L. By adopting the aforementioned preferred embodiment, the yield of product BTX is improved and the lifetime 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 below 400°C, and more preferably below 390°C. By adopting the aforementioned preferred embodiment, the yield of product BTX is improved and the lifetime 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 aforementioned preferred embodiment, the yield of product BTX is improved and the lifespan of the cracking catalyst is extended.

[0066] In this invention, any method that can prepare the aforementioned three-stage hydrocracking catalyst is acceptable. The preparation method can be a conventional choice in the art. According to a preferred embodiment of this invention, the preparation method of the three-stage hydrocracking catalyst includes: mixing, drying, and calcining a second active component source, a second support source, and a chelating surfactant. By adopting the aforementioned preferred scheme, the yield of product BTX is improved, and the lifetime 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 aforementioned preferred embodiment, the yield of product BTX is improved and the cracking catalyst lifetime is extended.

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

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

[0070] According to a preferred embodiment of the present invention, the conditions for the two-stage hydrorefining 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 for the two-stage hydrorefining 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 for the two-stage hydrorefining include: a liquid hourly space velocity (LHSV) of 0.2-3 h⁻¹. -1 Preferably 0.6-1.2h -1 .

[0073] According to a preferred embodiment of the present invention, the conditions for 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 for the three-stage hydrocracking include: space velocity of 0.2-3 h⁻¹. -1 Preferably 0.6-1.2h -1 .

[0075] According to a preferred embodiment of the present invention, the conditions for 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 for the three-stage hydrocracking include: a hydrogen-to-oil volume ratio of 400-3000, preferably 600-2000.

[0077] In this invention, there are no particular requirements for the two-stage hydrorefining method. According to a preferred embodiment of the invention, the two-stage hydrorefining method involves the material passing through at least two layers of catalyst along the flow direction, with the pore size of the first support in the catalyst decreasing sequentially. By adopting the aforementioned preferred scheme, the quality of the refined product can be further improved, the yield of BTX product can be increased, and the lifespan of the cracking catalyst can be extended.

[0078] A second aspect of the present invention provides an apparatus for utilizing inferior aromatic-rich distillate oil in the method described herein, the apparatus comprising a first-stage hydrogenation unit, a second-stage hydrorefining unit, a first purification unit, a third-stage hydrocracking unit, and a second purification unit arranged in series; wherein,

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

[0080] The temperature of the product stream outlet of the two-stage hydrorefining unit is higher than that of the raw material inlet, and the temperature difference is not greater than 75°C. It is used to refine the product stream of the first-stage hydrorefining unit, so that the naphthalene series compounds are selectively generated into tetrahydronaphthalene series compounds and the sulfur and nitrogen of the refined product are less than 1 ppm, thus obtaining the product stream of the two-stage hydrorefining unit.

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

[0082] The three-stage hydrocracking unit is used to crack a mixture of monocyclic aromatic hydrocarbons, so that tetrahydronaphthalene compounds are hydrocracking to generate BTX and LPG, and a three-stage hydrocracking product stream is obtained.

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

[0084] By utilizing the apparatus of this invention to process inferior aromatic distillate feedstocks, the quality of refined products can be further adjusted and improved, and the yields of BTX and LPG can be increased.

[0085] According to a preferred embodiment of the present invention, the outlet of the monocyclic aromatic hydrocarbon mixture of the first purification device is connected to a circulation pipeline that communicates with the feed inlet of the second-stage hydrorefining device. By adopting the aforementioned preferred embodiment, the quality of the refined product can be further adjusted and improved, and the yields of 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-precision tank, an oil-water separator, a stripping tower, a low-precision tank, and a distillation tower.

[0087] According to a preferred embodiment of the present invention, the two-stage hydrogenation unit is loaded with a catalyst for two-stage hydrogenation refining. The catalyst comprises a first active component, an active promoter, and a first support. The first active component comprises Group VIB and Group VIII metals, preferably at least one of Ni, Co, Fe, Pt, and Pd, and Mo and / or W. The active promoter comprises lanthanide metals and / or Group VA elements, preferably La, P, and Ce. The catalyst for two-stage hydrogenation refining is pre-sulfurized. The catalyst has 3-4 layers of sulfide active phase stacked in a tower-like structure, preferably with a stack length of 3-6 nm, and the active sites at the edges, corners, and sides are fully exposed. The three-stage hydrocracking unit is filled with a three-stage hydrocracking catalyst, which includes a second active component and a second support. The second active component contains a Group VIII metal, preferably Ni. The second support contains layered HZSM-5 / MCM-41 molecular sieves and optional Ce oxide and Zr oxide.

[0088] like Figure 4 As shown: This invention provides an exemplary device for utilizing inferior aromatic distillate oil, comprising: ethylene tar undergoing deweighting, first-stage hydrogenation to remove dienes, and reduction of bromine value, the product is heated by a second-stage heater and then enters a second-stage hydrorefining reactor; the product passes through a second-stage high-precision separator, then hydrogen enters a second-stage compressor, and the liquid phase enters a second-stage oil-water separator, then passes through a second-stage low-precision separator and enters a stripping tower to remove hydrogen sulfide gas; the purified product from the second stage is pumped by a second-stage circulation pump into the second-stage hydrorefining reactor, and by a third-stage heater and then into a third-stage hydrocracking reactor; the three-stage products are sequentially processed through three high-precision and low-precision sections and then enter a distillation column to obtain the final products BTX and LPG.

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

[0090] In the following examples, the evaluation method for hydrogenation of aromatic cracked distillate oil includes:

[0091] Feedstock for hydrorefining of aromatic distillate oil: distillation range 165-300℃, sulfur content = 500ppm, nitrogen content = 40ppm; feedstock composition: naphthalene series ≥ 40%.

[0092] Evaluation criteria:

[0093] Hydrorefining: Liquid hourly space velocity V = 0.8; Pressure P = 4.0 MPa; H2 / Oil (v / v) = 1000; and / or the mass ratio of the second-stage refining product recycle to the second-stage liquid feed is 0–2;

[0094] Hydrocracking: Temperature T = 400℃; Liquid hourly space velocity V = 0.8; H2 / Oil(v / v) = 600; Pressure P = 4.0 MPa.

[0095] In this invention, the dispersion of the active group Ni is tested by hydrogen-oxygen titration.

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

[0097] In the formula:

[0098] R-----Dispersion degree of Ni;

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

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

[0101] V0-----Titration volume of hydrogen, mL;

[0102] N A -----Avogadro's 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 this invention, the test method for the reduction of TPR hydrocracking catalyst in a hydrogen atmosphere is the hydrogen-oxygen titration method.

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

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

[0109] NH3-TPD was determined using an Altamira AMI-3300 chemisorption analyzer manufactured by Micron Instruments, Inc., USA.

[0110] The pore structure parameters of the carrier were measured using a Micrometrics Tristar3000 surface area analyzer at a test temperature of -196℃. Before the test, the sample was vacuum activated at 300℃ for 6 hours.

[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 Corporation, USA.

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

[0113] Example 1

[0114] 1. Preparation of hydrorefining catalyst

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

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

[0117] 1 L of Al2O3-SiO2 support (pore size: 10 nm, Al2O3 95%, SiO2 5%) was mixed with 0.75 L of a mixed solution of ammonium molybdate, cobalt nitrate, cobalt acetate, 1,2-cyclohexanediaminetetraacetic acid, and phosphoric acid (containing 96 g MoO3, 25 g CoO, 5 g La2O3, 52 g of 1,2-cyclohexanediaminetetraacetic acid, urea, ethanol, and 5 g P2O5). 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] The A3 catalyst contains 96 g / L MoO3, 5 g / L La2O3, 25 g / L CoO, and 5 g / L P2O5.

[0119] 2. Preparation of hydrocracking catalyst

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

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

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

[0123] 3. Hydrorefining catalyst loading

[0124] Catalysts A1, A2, and A3 in a volume ratio of 3:5:2 were sequentially loaded into an adiabatic bed reactor (a two-stage hydrorefining reactor, hereinafter the same).

[0125] 4. Hydrorefining catalyst sulfidation

[0126] At a system pressure of 0.5 MPa, N2 was introduced at a rate of 30 °C / h to raise the temperature to 150 °C. The N2 flow rate was 850 ml / min. N2 was then stopped, and H2 and sulfurized oil (containing 5000 ppm CS2) were introduced. The system pressure was raised to 2.5 MPa, and the temperature was increased at 5 °C / h for 6 hours to 180 °C. After holding at this temperature for 5 hours, the temperature was increased at 5 °C / h for 10 hours to 230 °C. After holding at this temperature for 4 hours, the temperature was increased at 10 °C / h for 5 hours to 280 °C. After holding at this temperature for 4 hours, the temperature was increased at 10 °C / h for 4 hours to 320 °C. This temperature was held for 2 hours, completing the catalyst sulfidation. With the sulfurized oil introduced, the bed inlet temperature was lowered to the required reaction temperature.

[0127] TEM images of the A2 catalyst are shown in 1a, indicating that the active phase of the hydrogenation catalyst exhibits a tower-like structure after sulfidation, with 3-5 stacked layers and a length of 3-6 nm, fully exposing active sites such as edges, corners, and sides. Studies have shown that hydrogenation catalysts with this type of active phase structure have high hydrogenation activity and also possess stronger desulfurization and denitrification activity.

[0128] 5. Hydrocracking catalyst loading and reduction

[0129] Take B1 and load it into an adiabatic bed reactor (three-stage hydrocracking reactor, the same below). Heat it from room temperature to 500℃ at a rate of 20℃ / h and keep it at that temperature for 3 hours. Use a hydrogen atmosphere with a hydrogen gas flow rate of 500ml / min.

[0130] 6. Catalyst Evaluation

[0131] (The feedstock of the low-quality aromatic fraction oil comes from the product stream of ethylene tar, with a final boiling point of 300℃.) The first-stage hydrogenation product (bromine value of 40mgBr / 100g oil and gum content of 48mg / 100g oil) is mixed with fresh hydrogen and second-stage circulating hydrogen, heated in the second-stage heater, and then fed into the second-stage hydrorefining reactor. The hydrorefined product is fed into the second-stage high-precision tank, the hydrogen is circulated in the second-stage compressor, the liquid phase is fed into the second-stage oil-water separator, and the oil phase is fed into the second-stage low-precision tank. After H2S is removed from the oil phase by the stripping tower, the circulation rate of the second-stage purified product is 0.5 to the mass ratio of the second-stage liquid phase feed. The second-stage purified product is heated in the third-stage heater and then fed into the third-stage hydrocracking reactor. The cracked product is fed into the third-stage high-precision tank, the hydrogen is circulated in the third-stage compressor, and the liquid phase is fractionated in the third-stage low-precision tank to obtain BTX and LPG.

[0132] The temperature of the hydrorefining feedstock stream is 225℃, and the temperature of the hydrorefining product stream is 275℃.

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

[0134] Example 2

[0135] Same as Example 1, except that A1 is replaced with A4, and the preparation of the hydrorefining catalyst A4 is as follows:

[0136] 1 L of Al₂O₃ support (pore size: 16 nm) was mixed with 0.75 L of a mixed solution of nickel acetate, nickel nitrate, ammonium molybdate, lanthanum nitrate, 1,2-cyclohexanediaminetetraacetic acid, urea, ethanol, and phosphoric acid (containing 40 g NiO, 108 g MoO₃, 5 g La₂O₃, 36 g 1,2-cyclohexanediaminetetraacetic acid, urea, ethanol, and 10 g P₂O₅). 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] The A4 catalyst contains 28 g / L NiO, 108 g / L MoO3, 5 g / L La2O3, and 10 g / L P2O5.

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

[0139] Example 3

[0140] 1. Hydrorefining catalyst loading

[0141] Catalysts A2 and A3, in a volume ratio of 8:2, were sequentially loaded into the adiabatic bed reactor.

[0142] 2. Hydrorefining catalyst sulfidation

[0143] At a system pressure of 0.5 MPa, N2 was introduced at a rate of 30 °C / h to raise the temperature to 150 °C. The N2 flow rate was 850 ml / min. N2 was then stopped, and H2 and sulfurized oil (containing 5000 ppm CS2) were introduced. The system pressure was raised to 2.5 MPa, and the temperature was increased at 5 °C / h for 6 hours to 180 °C. After holding at this temperature for 5 hours, the temperature was increased at 5 °C / h for 10 hours to 230 °C. After holding at this temperature for 4 hours, the temperature was increased at 10 °C / h for 5 hours to 280 °C. After holding at this temperature for 4 hours, the temperature was increased at 10 °C / h for 4 hours to 320 °C. This temperature was held for 2 hours, completing the catalyst sulfidation. With the sulfurized oil introduced, the bed inlet temperature was lowered to the required reaction temperature.

[0144] 3. Hydrocracking catalyst loading and reduction

[0145] B1 was loaded into an adiabatic bed reactor and heated from room temperature to 500℃ at a rate of 20℃ / h. The temperature was maintained for 3 hours under a hydrogen atmosphere at a hydrogen gas flow rate of 500 ml / min.

[0146] 4. Catalyst Evaluation

[0147] (The feedstock of the low-quality aromatic fraction oil comes from the product stream of ethylene tar, with a final boiling point of 300℃.) The first-stage hydrogenation product (bromine value of 40mgBr / 100g oil and gum content of 48mg / 100g oil) is mixed with fresh hydrogen and second-stage circulating hydrogen, heated in the second-stage heater, and then fed into the second-stage hydrorefining reactor. The hydrorefined product is fed into the second-stage high-precision tank, the hydrogen is circulated in the second-stage compressor, the liquid phase is fed into the second-stage oil-water separator, and the oil phase is fed into the second-stage low-precision tank. After H2S is removed from the oil phase by the stripping tower, the circulation rate of the second-stage purified product is 0.5 to the mass ratio of the second-stage liquid phase feed. The second-stage purified product is heated in the third-stage heater and then fed into the third-stage hydrocracking reactor. The cracked product is fed into the third-stage high-precision tank, the hydrogen is circulated in the third-stage compressor, and the liquid phase is fractionated in the third-stage low-precision tank to obtain BTX and LPG.

[0148] The temperature of the hydrorefining feedstock stream is 225℃, and the temperature of the hydrorefining product stream is 290℃.

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

[0150] Example 4

[0151] 1. Hydrorefining catalyst loading

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

[0153] 2. Hydrorefining catalyst sulfidation

[0154] At a system pressure of 0.5 MPa, N2 was introduced at a rate of 30 °C / h to raise the temperature to 150 °C. The N2 flow rate was 850 ml / min. N2 was then stopped, and H2 and sulfurized oil (containing 5000 ppm CS2) were introduced. The system pressure was raised to 2.5 MPa, and the temperature was increased at 5 °C / h for 6 hours to 180 °C. After holding at this temperature for 5 hours, the temperature was increased at 5 °C / h for 10 hours to 230 °C. After holding at this temperature for 4 hours, the temperature was increased at 10 °C / h for 5 hours to 280 °C. After holding at this temperature for 4 hours, the temperature was increased at 10 °C / h for 4 hours to 320 °C. This temperature was held for 2 hours, completing the catalyst sulfidation. With the sulfurized oil introduced, the bed inlet temperature was lowered to the required reaction temperature.

[0155] 3. Reduction by hydrocracking catalyst

[0156] B1 was loaded into an adiabatic bed reactor and heated from room temperature to 500℃ at a rate of 20℃ / h. The temperature was maintained for 3 hours under a hydrogen atmosphere at a hydrogen gas flow rate of 500 ml / min.

[0157] 4. Catalyst Evaluation

[0158] (The feedstock of the low-quality aromatic fraction oil comes from the product stream of ethylene tar, with a final boiling point of 300℃.) The first-stage hydrogenation product (bromine value of 40mgBr / 100g oil and gum content of 48mg / 100g oil) is mixed with fresh hydrogen and second-stage circulating hydrogen, heated in the second-stage heater, and then fed into the second-stage hydrorefining reactor. The hydrorefined product is fed into the second-stage high-precision tank, the hydrogen is circulated in the second-stage compressor, the liquid phase is fed into the second-stage oil-water separator, and the oil phase is fed into the second-stage low-precision tank. After H2S is removed from the oil phase by the stripping tower, the second-stage purified product is not circulated back to the inlet of the hydrorefining reactor. The second-stage purified product is heated in the third-stage heater and then fed into the third-stage hydrocracking reactor. The cracked product is fed into the third-stage high-precision tank, the hydrogen is circulated in the third-stage compressor, and the liquid phase is fed into the third-stage low-precision tank and then fractionated in the distillation column to obtain BTX and LPG.

[0159] The temperature of the hydrorefining feedstock stream is 225℃, and the temperature of the hydrorefining product stream is 300℃.

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

[0161] Example 5

[0162] 1. Hydrorefining catalyst loading

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

[0164] 2. Hydrorefining catalyst sulfidation

[0165] At a system pressure of 0.5 MPa, N2 was introduced at a rate of 30 °C / h to raise the temperature to 150 °C. The N2 flow rate was 850 ml / min. N2 was then stopped, and H2 and sulfurized oil (containing 5000 ppm CS2) were introduced. The system pressure was raised to 2.5 MPa, and the temperature was increased at 5 °C / h for 6 hours to 180 °C. After holding at this temperature for 5 hours, the temperature was increased at 5 °C / h for 10 hours to 230 °C. After holding at this temperature for 4 hours, the temperature was increased at 10 °C / h for 5 hours to 280 °C. After holding at this temperature for 4 hours, the temperature was increased at 10 °C / h for 4 hours to 320 °C. This temperature was held for 2 hours, completing the catalyst sulfidation. With the sulfurized oil introduced, the bed inlet temperature was lowered to the required reaction temperature.

[0166] 3. Hydrocracking catalyst loading and reduction

[0167] B1 was loaded into an adiabatic bed reactor and heated from room temperature to 500℃ at a rate of 20℃ / h. The temperature was maintained for 3 hours under a hydrogen atmosphere at a hydrogen gas flow rate of 500 ml / min.

[0168] 4. Catalyst Evaluation

[0169] (The feedstock of the low-quality aromatic fraction oil comes from the product stream of C9 cracking, with a final boiling point of 275℃.) The first-stage hydrogenation product (bromine value of 35mgBr / 100g oil and gum content of 40mg / 100g oil) is mixed with fresh hydrogen and second-stage circulating hydrogen, heated in the second-stage heater, and then fed into the second-stage hydrorefining reactor. The hydrorefined product is fed into the second-stage high-precision tank, the hydrogen is circulated in the second-stage compressor, the liquid phase is fed into the second-stage oil-water separator, and the oil phase is fed into the second-stage low-precision tank. After H2S is removed from the oil phase by the stripping tower, the circulation rate of the second-stage purified product is 0.5 to the mass ratio of the second-stage liquid phase feed. The second-stage purified product is heated in the third-stage heater and then fed into the third-stage hydrocracking reactor. The cracked product is fed into the third-stage high-precision tank, the hydrogen is circulated in the third-stage compressor, and the liquid phase is fed into the third-stage low-precision tank and then fractionated in the distillation column to obtain BTX and LPG.

[0170] The temperature of the hydrorefining feedstock stream is 225℃, and the temperature of the hydrorefining product stream is 270℃.

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

[0172] Example 6

[0173] 1. Hydrorefining catalyst loading

[0174] The A4 and A3 catalysts, in a volume ratio of 8:2, were sequentially loaded into the adiabatic bed reactor.

[0175] 2. Hydrorefining catalyst sulfidation

[0176] At a system pressure of 0.5 MPa, N2 was introduced at a rate of 30 °C / h to raise the temperature to 150 °C. The N2 flow rate was 850 ml / min. N2 was then stopped, and H2 and sulfurized oil (containing 5000 ppm CS2) were introduced. The system pressure was raised to 2.5 MPa, and the temperature was increased at 5 °C / h for 6 hours to 180 °C. After holding at this temperature for 5 hours, the temperature was increased at 5 °C / h for 10 hours to 230 °C. After holding at this temperature for 4 hours, the temperature was increased at 10 °C / h for 5 hours to 280 °C. After holding at this temperature for 4 hours, the temperature was increased at 10 °C / h for 4 hours to 320 °C. This temperature was held for 2 hours, completing the catalyst sulfidation. With the sulfurized oil introduced, the bed inlet temperature was lowered to the required reaction temperature.

[0177] 3. Hydrocracking catalyst loading and reduction

[0178] B1 was loaded into an adiabatic bed reactor and heated from room temperature to 500℃ at a rate of 20℃ / h. The temperature was maintained for 3 hours under a hydrogen atmosphere at a hydrogen gas flow rate of 500 ml / min.

[0179] 4. Catalyst Evaluation

[0180] (The feedstock of the low-quality aromatic fraction oil comes from the product stream of C9 cracking, with a final boiling point of 275℃.) The first-stage hydrogenation product (bromine value of 35mgBr / 100g oil and gum content of 40mg / 100g oil) is mixed with fresh hydrogen and second-stage circulating hydrogen, heated in the second-stage heater, and then fed into the second-stage hydrorefining reactor. The hydrorefined product is fed into the second-stage high-precision tank, the hydrogen is circulated in the second-stage compressor, the liquid phase is fed into the second-stage oil-water separator, and the oil phase is fed into the second-stage low-precision tank. After H2S is removed from the oil phase by the stripping tower, the circulation rate of the second-stage purified product is 0.5 to the mass ratio of the second-stage liquid phase feed. The second-stage purified product is heated in the third-stage heater and then fed into the third-stage hydrocracking reactor. The cracked product is fed into the third-stage high-precision tank, the hydrogen is circulated in the third-stage compressor, and the liquid phase is fed into the third-stage low-precision tank and then fractionated in the distillation column to obtain BTX and LPG.

[0181] The temperature of the hydrorefining feedstock stream is 225℃, and the temperature of the hydrorefining product stream is 258℃.

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

[0183] Example 7

[0184] 1. Hydrorefining catalyst loading

[0185] Commercially available NiMo catalyst C1, NiMoW catalyst C2, and CoMo catalyst C3 were loaded into an adiabatic bed reactor in a volume ratio of 3:5:2.

[0186] 2. Hydrorefining catalyst sulfidation

[0187] At a system pressure of 0.5 MPa, N2 was introduced at a rate of 30 °C / h to raise the temperature to 150 °C. The N2 flow rate was 850 ml / min. N2 was then stopped, and H2 and sulfurized oil (containing 5000 ppm CS2) were introduced. The system pressure was raised to 2.5 MPa, and the temperature was increased at 5 °C / h for 6 hours to 180 °C. After holding at this temperature for 5 hours, the temperature was increased at 5 °C / h for 10 hours to 230 °C. After holding at this temperature for 4 hours, the temperature was increased at 10 °C / h for 5 hours to 280 °C. After holding at this temperature for 4 hours, the temperature was increased at 10 °C / h for 4 hours to 320 °C. This temperature was held for 2 hours, completing the catalyst sulfidation. With the sulfurized oil introduced, the bed inlet temperature was lowered to the required reaction temperature.

[0188] 3. Hydrocracking catalyst loading and reduction

[0189] B1 was loaded into an adiabatic bed reactor and heated from room temperature to 500℃ at a rate of 20℃ / h. The temperature was maintained for 3 hours under a hydrogen atmosphere at a hydrogen gas flow rate of 500 ml / min.

[0190] 4. Catalyst Evaluation

[0191] (The feedstock of the low-quality aromatic fraction oil comes from the product stream of ethylene tar, with a final boiling point of 300℃.) The first-stage hydrogenation product (bromine value of 40mgBr / 100g oil and gum content of 48mg / 100g oil) is mixed with fresh hydrogen and second-stage circulating hydrogen, heated in the second-stage heater, and then fed into the second-stage hydrorefining reactor. The hydrorefined product is fed into the second-stage high-precision tank, the hydrogen is circulated in the second-stage compressor, the liquid phase is fed into the second-stage oil-water separator, and the oil phase is fed into the second-stage low-precision tank. After H2S is removed from the oil phase by the stripping tower, the circulation rate of the second-stage purified product is 0.5 to the mass ratio of the second-stage liquid phase feed. The second-stage purified product is heated in the third-stage heater and then fed into the third-stage hydrocracking reactor. The cracked product is fed into the third-stage high-precision tank, the hydrogen is circulated in the third-stage compressor, and the liquid phase is fractionated in the third-stage low-precision tank to obtain BTX and LPG.

[0192] The temperature of the hydrorefining feedstock stream is 225℃, and the temperature of the hydrorefining product stream is 278℃.

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

[0194] Example 8

[0195] 1. Hydrorefining catalyst loading

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

[0197] 2. Hydrorefining catalyst sulfidation

[0198] At a system pressure of 0.5 MPa, N2 was introduced at a rate of 30 °C / h to raise the temperature to 150 °C. The N2 flow rate was 850 ml / min. N2 was then stopped, and H2 and sulfurized oil (containing 5000 ppm CS2) were introduced. The system pressure was raised to 2.5 MPa, and the temperature was increased at 5 °C / h for 6 hours to 180 °C. After holding at this temperature for 5 hours, the temperature was increased at 5 °C / h for 10 hours to 230 °C. After holding at this temperature for 4 hours, the temperature was increased at 10 °C / h for 5 hours to 280 °C. After holding at this temperature for 4 hours, the temperature was increased at 10 °C / h for 4 hours to 320 °C. This temperature was held for 2 hours, completing the catalyst sulfidation. With the sulfurized oil introduced, the bed inlet temperature was lowered to the required reaction temperature.

[0199] 3. Hydrocracking catalyst loading and reduction

[0200] B1 was loaded into an adiabatic bed reactor and heated from room temperature to 500℃ at a rate of 20℃ / h. The temperature was maintained for 3 hours under a hydrogen atmosphere at a hydrogen gas flow rate of 500 ml / min.

[0201] 4. Catalyst Evaluation

[0202] (The feedstock of the low-quality aromatic fraction oil comes from the product stream of ethylene tar, with a final boiling point of 300℃.) The first-stage hydrogenation product (bromine value of 40mgBr / 100g oil and gum content of 48mg / 100g oil) is mixed with fresh hydrogen and second-stage circulating hydrogen, heated in the second-stage heater, and then fed into the second-stage hydrorefining reactor. The hydrorefined product is fed into the second-stage high-precision tank, the hydrogen is circulated in the second-stage compressor, the liquid phase is fed into the second-stage oil-water separator, and the oil phase is fed into the second-stage low-precision tank. After H2S is removed from the oil phase by the stripping tower, the circulation rate of the second-stage purified product to the second-stage liquid phase feed mass ratio is 2.1. The second-stage purified product is heated in the third-stage heater and then fed into the third-stage hydrocracking reactor. The cracked product is fed into the third-stage high-precision tank, the hydrogen is circulated in the third-stage compressor, and the liquid phase is fractionated in the third-stage low-precision tank to obtain BTX and LPG.

[0203] The temperature of the hydrorefining feedstock stream is 225℃, and the temperature of the hydrorefining product stream is 247℃.

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

[0205] Example 9

[0206] 1. Hydrorefining catalyst loading

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

[0208] 2. Hydrorefining catalyst sulfidation

[0209] At a system pressure of 0.5 MPa, N2 was introduced at a rate of 30 °C / h to raise the temperature to 150 °C. The N2 flow rate was 850 ml / min. N2 was then stopped, and H2 and sulfurized oil (containing 5000 ppm CS2) were introduced. The system pressure was raised to 2.5 MPa, and the temperature was increased at 5 °C / h for 6 hours to 180 °C. After holding at this temperature for 5 hours, the temperature was increased at 5 °C / h for 10 hours to 230 °C. After holding at this temperature for 4 hours, the temperature was increased at 10 °C / h for 5 hours to 280 °C. After holding at this temperature for 4 hours, the temperature was increased at 10 °C / h for 4 hours to 320 °C. This temperature was held for 2 hours, completing the catalyst sulfidation. With the sulfurized oil introduced, the bed inlet temperature was lowered to the required reaction temperature.

[0210] 3. Hydrocracking catalyst loading and reduction

[0211] B1 was loaded into an adiabatic bed reactor and heated from room temperature to 500℃ at a rate of 20℃ / h. The temperature was maintained for 3 hours under a hydrogen atmosphere at a hydrogen gas flow rate of 500 ml / min.

[0212] 4. Catalyst Evaluation

[0213] (The feedstock of the low-quality aromatic fraction oil comes from the product stream of ethylene tar, with a final boiling point of 300℃.) The first-stage hydrogenation product (bromine value of 40mgBr / 100g oil and gum content of 48mg / 100g oil) is mixed with fresh hydrogen and second-stage circulating hydrogen, heated in the second-stage heater, and then fed into the second-stage hydrorefining reactor. The hydrorefined product is fed into the second-stage high-precision tank, the hydrogen is circulated in the second-stage compressor, the liquid phase is fed into the second-stage oil-water separator, and the oil phase is fed into the second-stage low-precision tank. After H2S is removed from the oil phase by the stripping tower, the circulation rate of the second-stage purified product is 0.1 to the mass ratio of the second-stage liquid phase feed. The second-stage purified product is heated in the third-stage heater and then fed into the third-stage hydrocracking reactor. The cracked product is fed into the third-stage high-precision tank, the hydrogen is circulated in the third-stage compressor, and the liquid phase is fractionated in the third-stage low-precision tank to obtain BTX and LPG.

[0214] The temperature of the hydrorefining feedstock stream is 225℃, and the temperature of the hydrorefining product stream is 305℃.

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

[0216] Example 10

[0217] 1. Preparation of hydrorefining catalyst

[0218] 1 L of Al2O3-TiO2-SiO2 (pore size: 11 nm, Al2O3 67%, TiO2 5%, SiO2 28%) was mixed with 0.75 L of a mixed solution of nickel acetate, ammonium metatungstate, lanthanum nitrate, ammonium molybdate, and phosphoric acid (containing 40 g NiO, 60 g MoO3, 5 g La2O3, 165 g WO3, and 10 g P2O5). The mixture was dried at 110 °C for 6 hours and calcined at 450 °C for 4 hours to prepare catalyst A5. The TEM image after sulfidation is shown below. Figure 1b As can be seen from the figure, the sulfide catalyst partially agglomerates, and the crystallite length is 5-15 nm.

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

[0220] 2. Hydrorefining catalyst loading

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

[0222] 3. Hydrorefining catalyst sulfidation

[0223] At a system pressure of 0.5 MPa, N2 was introduced at a rate of 30 °C / h to raise the temperature to 150 °C. The N2 flow rate was 850 ml / min. N2 was then stopped, and H2 and sulfurized oil (containing 5000 ppm CS2) were introduced. The system pressure was raised to 2.5 MPa, and the temperature was increased at 5 °C / h for 6 hours to 180 °C. After holding at this temperature for 5 hours, the temperature was increased at 5 °C / h for 10 hours to 230 °C. After holding at this temperature for 4 hours, the temperature was increased at 10 °C / h for 5 hours to 280 °C. After holding at this temperature for 4 hours, the temperature was increased at 10 °C / h for 4 hours to 320 °C. This temperature was held for 2 hours, completing the catalyst sulfidation. With the sulfurized oil introduced, the bed inlet temperature was lowered to the required reaction temperature.

[0224] 4. Reduction by hydrocracking catalyst

[0225] B1 was loaded into an adiabatic bed reactor and heated from room temperature to 500℃ at a rate of 20℃ / h. The temperature was maintained for 3 hours under a hydrogen atmosphere at a hydrogen gas flow rate of 500 ml / min.

[0226] 5. Catalyst Evaluation

[0227] (The feedstock of the low-quality aromatic fraction oil comes from the product stream of ethylene tar, with a final boiling point of 300℃.) The first-stage hydrogenation product (bromine value of 40mgBr / 100g oil and gum content of 48mg / 100g oil) is mixed with fresh hydrogen and second-stage circulating hydrogen, heated in the second-stage heater, and then fed into the second-stage hydrorefining reactor. The hydrorefined product is fed into the second-stage high-precision tank, the hydrogen is circulated in the second-stage compressor, the liquid phase is fed into the second-stage oil-water separator, and the oil phase is fed into the second-stage low-precision tank. After H2S is removed from the oil phase by the stripping tower, the circulation rate of the second-stage purified product is 0.5 to the mass ratio of the second-stage liquid phase feed. The second-stage purified product is heated in the third-stage heater and then fed into the third-stage hydrocracking reactor. The cracked product is fed into the third-stage high-precision tank, the hydrogen is circulated in the third-stage compressor, and the liquid phase is fractionated in the third-stage low-precision tank to obtain BTX and LPG.

[0228] The temperature of the hydrorefining feedstock stream is 225℃, and the temperature of the hydrorefining product stream is 275℃.

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

[0230] Example 11

[0231] 1. Hydrorefining catalyst loading

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

[0233] 2. Hydrorefining catalyst sulfidation

[0234] At a system pressure of 0.5 MPa, N2 was introduced at a rate of 30 °C / h to raise the temperature to 150 °C. The N2 flow rate was 850 ml / min. N2 was then stopped, and H2 and sulfurized oil (containing 5000 ppm CS2) were introduced. The system pressure was raised to 2.5 MPa, and the temperature was increased at 5 °C / h for 6 hours to 180 °C. After holding at this temperature for 5 hours, the temperature was increased at 5 °C / h for 10 hours to 230 °C. After holding at this temperature for 4 hours, the temperature was increased at 10 °C / h for 5 hours to 280 °C. After holding at this temperature for 4 hours, the temperature was increased at 10 °C / h for 4 hours to 320 °C. This temperature was held for 2 hours, completing the catalyst sulfidation. With the sulfurized oil introduced, the bed inlet temperature was lowered to the required reaction temperature.

[0235] 3. Reduction by hydrocracking catalyst

[0236] B1 was loaded into an adiabatic bed reactor and heated from room temperature to 500℃ at a rate of 20℃ / h. The temperature was maintained for 3 hours under a hydrogen atmosphere at a hydrogen gas flow rate of 500 ml / min.

[0237] 4. Catalyst Evaluation

[0238] (The feedstock of the low-quality aromatic fraction oil comes from the product stream of ethylene tar, with a final boiling point of 300℃.) The first-stage hydrogenation product (bromine value of 40mgBr / 100g oil and gum content of 48mg / 100g oil) is mixed with fresh hydrogen and second-stage circulating hydrogen, heated in the second-stage heater, and then fed into the second-stage hydrorefining reactor. The hydrorefined product is fed into the second-stage high-precision tank, the hydrogen is circulated in the second-stage compressor, the liquid phase is fed into the second-stage oil-water separator, and the oil phase is fed into the second-stage low-precision tank. After H2S is removed from the oil phase by the stripping tower, the circulation rate of the second-stage purified product is 0.5 to the mass ratio of the second-stage liquid phase feed. The second-stage purified product is heated in the third-stage heater and then fed into the third-stage hydrocracking reactor. The cracked product is fed into the third-stage high-precision tank, the hydrogen is circulated in the third-stage compressor, and the liquid phase is fractionated in the third-stage low-precision tank to obtain BTX and LPG.

[0239] The temperature of the hydrorefining feedstock stream is 225℃, and the temperature of the hydrorefining product stream is 260℃.

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

[0241] Example 12

[0242] 1. Hydrorefining catalyst loading

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

[0244] 2. Hydrorefining catalyst sulfidation

[0245] At a system pressure of 0.5 MPa, N2 was introduced at a rate of 30 °C / h to raise the temperature to 150 °C. The N2 flow rate was 850 ml / min. N2 was then stopped, and H2 and sulfurized oil (containing 5000 ppm CS2) were introduced. The system pressure was raised to 2.5 MPa, and the temperature was increased at 5 °C / h for 6 hours to 180 °C. After holding at this temperature for 5 hours, the temperature was increased at 5 °C / h for 10 hours to 230 °C. After holding at this temperature for 4 hours, the temperature was increased at 10 °C / h for 5 hours to 280 °C. After holding at this temperature for 4 hours, the temperature was increased at 10 °C / h for 4 hours to 320 °C. This temperature was held for 2 hours, completing the catalyst sulfidation. With the sulfurized oil introduced, the bed inlet temperature was lowered to the required reaction temperature.

[0246] 3. Hydrocracking catalyst loading and reduction

[0247] B1 was loaded into an adiabatic bed reactor and heated from room temperature to 500℃ at a rate of 20℃ / h. The temperature was maintained for 3 hours under a hydrogen atmosphere at a hydrogen gas flow rate of 500 ml / min.

[0248] 4. Catalyst Evaluation

[0249] (The feedstock of the low-quality aromatic fraction oil comes from the product stream of ethylene tar, with a final boiling point of 300℃.) After the feedstock is deweighted (bromine value of 150mgBr / 100g oil, gum content of 100mg / 100g oil), it is mixed with fresh hydrogen and second-stage circulating hydrogen, heated in the second-stage heater, and then fed into the second-stage hydrorefining reactor. The hydrorefining product is fed into the second-stage high-precision tank, the hydrogen is circulated in the second-stage compressor, the liquid phase is fed into the second-stage oil-water separator, and the oil phase is fed into the second-stage low-precision tank. After H2S is removed from the oil phase by the stripping tower, the circulation rate of the second-stage purified product is 0.5 to the mass ratio of the second-stage liquid phase feed. The second-stage purified product is heated in the third-stage heater and then fed into the third-stage hydrocracking reactor. The cracking product is fed into the third-stage high-precision tank, the hydrogen is circulated in the third-stage compressor, and the liquid phase is fractionated in the third-stage low-precision tank to obtain BTX and LPG.

[0250] The temperature of the hydrorefining feedstock stream is 225℃, and the temperature of the hydrorefining product stream is 317℃.

[0251] The evaluation results are listed in Table 1 (refining) 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 Example 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 have been described in detail above; however, the present invention is not limited thereto. Within the scope of the inventive concept, various simple modifications can be made to the technical solutions of the present invention, including combinations of various technical features in any other suitable manner. These simple modifications and combinations should also be considered as the content disclosed in the present invention and are all within the protection scope of the present invention.

Claims

1. A method for utilizing inferior aromatic-rich distillate oil, characterized in that, The method includes sequentially subjecting a low-quality aromatic-rich distillate feedstock to a first-stage hydrotreating process to convert dienes and reduce bromine value, a second-stage hydrorefining process, a first purification process, a third-stage hydrocracking process, and a second purification process to separate BTX and LPG. The final boiling point of the low-quality aromatic-rich distillate feedstock is ≤300℃. The low-quality aromatic-rich distillate feedstock is selected from at least one of the following: cracked C9 product stream, cracked ethylene tar product stream, and cracked coal tar product stream. The temperature of the product stream from the second-stage hydrorefining process is higher than the temperature of the feedstock stream before the second-stage hydrorefining process, and the temperature difference is not greater than 75℃. The catalyst for the second-stage hydrorefining process includes a first active component, an active additive, and a first support. The first active component contains Group VIB and Group VIII metals, and the active additive contains lanthanide metals and / or Group VA elements. After pre-sulfurization, the catalyst for the second-stage hydrorefining process has 3-4 layers of sulfide active phase stacked in a tower-like structure.

2. The method according to claim 1, wherein, The temperature difference between the product stream from the second-stage hydrorefining and the raw material stream before the second-stage hydrorefining is 30-75°C.

3. The method according to claim 2, wherein, The temperature difference between the product stream from the second-stage hydrorefining and the raw material stream before the second-stage hydrorefining is 35-60℃.

4. The method according to any one of claims 1-3, wherein, The temperature of the product stream from the two-stage hydrorefining process is 250-310°C; and / or The temperature of the feed stream before the second-stage hydrorefining is 200-260℃.

5. The method according to claim 4, wherein, The temperature of the product stream from the two-stage hydrorefining process is 260-300°C; and / or The temperature of the feedstock stream before the second-stage hydrorefining is 210-250℃.

6. The method according to claim 1, wherein, The feedstock stream before the two-stage hydrogenation refining includes the first purified feedstock stream.

7. The method according to claim 6, wherein, The mass content of the first purified stream in the feed stream before the two-stage hydrogenation refining is >0% and ≤90%.

8. The method according to claim 7, wherein, The mass content of the first purified stream in the feed stream before the two-stage hydrogenation refining is >0% and ≤50%.

9. The method according to claim 1, wherein, The bromine value of the product stream after hydrogenation is ≤40 mgBr / 100g oil, and the gum content is ≤50 mg / 100g oil; and / or The sulfur content in the first purified product stream is less than 1 ppm by mass, and / or the nitrogen content is less than 1 ppm by mass.

10. The method according to claim 9, wherein, The sulfur content in the first purified product stream is less than 0.85 ppm by mass, and / or the nitrogen content is less than 0.8 ppm by mass.

11. The method according to claim 1, wherein, The first active component comprises at least one of Ni, Co, Fe, Pt and Pd, as well as Mo and / or W; the active additive is La, P, or Ce; and the sulfide active phase has a stacking length of 3-6 nm.

12. The method according to claim 1, wherein, The first support in the catalyst for the two-stage hydrorefining is selected from at least one of Al2O3, Al2O3-TiO2, Al2O3-TiO2-SiO2, and Al2O3-SiO2; and / or The average pore diameter of the first carrier in the catalyst for the two-stage hydrofining is 7-20 nm; and / or the pore volume is 0.5-0.98 cm 3 / g; and / or the specific surface area is 190-380 m 2 / g.

13. The method according to claim 12, wherein, The first support in the catalyst for the two-stage hydrorefining is selected from one or two of Al2O3, Al2O3-TiO2, and Al2O3-TiO2-SiO2.

14. The method according to claim 1, wherein, In the catalyst for the two-stage hydrorefining, the density of the first active component, calculated as oxide, is 50-300 g / L, and the density of the active auxiliary agent, calculated as oxide, is 5-50 g / L.

15. The method according to claim 14, wherein, In the catalyst for two-stage hydrorefining, the density of the first active component, calculated as oxide, is 100-300 g / L, and the density of the active auxiliary agent, calculated as oxide, is 5-40 g / L.

16. The method according to claim 1, wherein, The preparation method of the catalyst for the two-stage hydrorefining includes: mixing a first active component source, an active auxiliary agent source, a first support source and an additive to obtain a mixed solution, drying and calcining; in, The concentration of Group VIB metals in the mixed solution, calculated as oxides, is 150-850 g / L; the concentration of Group VIII metals, calculated as oxides, is 50-300 g / L. The additive is a mixture of amino compounds, organic acid metal complexing agents, and alcohols.

17. The method according to claim 16, wherein, The amino compound is selected from at least one of urea, acetamide, ethylenediamine, ethanolamine, diethanolamine, and triethanolamine; The organic acid-metal complexing agent is selected from at least one of 1,2-cyclohexanediaminetetraacetic acid, citric acid, tartaric acid, and ethylenediaminetetraacetic acid; The alcohol is ethylene glycol and / or ethanol; and / or The mass fraction of the additive in the mixture is 0.01-10%.

18. The method according to claim 17, wherein, The mass fraction of the additive in the mixture is 1.5-7.5%.

19. The method of claim 17, wherein, The mass ratio of amino compound, organic acid metal complexing agent and alcohol in the additive is 1-2:1-2:1-2.

20. The method according to claim 1, wherein, The catalyst for three-stage hydrocracking includes a second active component and a second support. The second active component contains a Group VIII metal, and the second support contains a layered HZSM-5 / MCM-41 molecular sieve.

21. The method according to claim 20, wherein, The second carrier also contains Ce oxide and Zr oxide.

22. The method according to claim 20 or 21, wherein, The second active component is Ni.

23. The method according to claim 20 or 21, wherein, The interlayer thickness of the layered HZSM-5 / MCM-41 molecular sieve is 5-10 nm; and / or The pore size of the layered HZSM-5 / MCM-41 molecular sieve is 3-18 nm; and / or The medium-strong acid content / total acid content of the layered HZSM-5 / MCM-41 molecular sieve is 0.55-0.85; and / or The specific surface area of the layered HZSM-5 / MCM-41 molecular sieve is ≧400 m 2 ·g -1 ; and / or The SiO2 / Al2O3 molar ratio of the layered HZSM-5 / MCM-41 molecular sieve is 20-200.

24. The method according to claim 23, wherein, The interlayer thickness of the layered HZSM-5 / MCM-41 molecular sieve is 5-7 nm; and / or The pore size of the layered HZSM-5 / MCM-41 molecular sieve is 3-11 nm; and / or The medium-strong acid content / total acid content of the layered HZSM-5 / MCM-41 molecular sieve is 0.55-0.65; and / or The specific surface area of the layered HZSM-5 / MCM-41 molecular sieve is 400-550 m 2 ·g -1 ; and / or The SiO2 / Al2O3 molar ratio of the layered HZSM-5 / MCM-41 molecular sieve is 20-60.

25. The method according to claim 23, wherein, In the catalyst for the three-stage hydrocracking, the content of the second active component, calculated as oxide, is 50-260 g / L, and the content of Ce oxide and Zr oxide, calculated as oxide, is 0.1-100 g / L respectively.

26. The method of claim 25, wherein, In the catalyst for the three-stage hydrocracking, the content of the second active component, calculated as oxide, is 50-150 g / L, and the content of Ce oxide and Zr oxide, calculated as oxide, is 0.1-50 g / L respectively.

27. The method according to claim 23, wherein, The TPR hydrogen atmosphere reduction temperature of the catalyst used in the three-stage hydrocracking process is below 400°C; and / or The dispersion of the second active component element in the catalyst for the three-stage hydrocracking is greater than 8%.

28. The method according to claim 27, wherein, The TPR hydrogen atmosphere reduction temperature of the catalyst used in the three-stage hydrocracking process is below 390°C; and / or The dispersion of the second active component element in the catalyst for the three-stage hydrocracking is 8-15%.

29. The method according to claim 20, wherein, The preparation method of the catalyst for three-stage hydrocracking includes: mixing, drying and calcining a second active component source, a second support source and a chelating surfactant, wherein the chelating surfactant is selected from at least one of 1,2-cyclohexanediaminetetraacetic acid, citric acid, tartaric acid and ethylenediaminetetraacetic acid.

30. The method according to claim 29, wherein, The chelating surfactant is one of 1,2-cyclohexanediaminetetraacetic acid, citric acid, and ethylenediaminetetraacetic acid.

31. The method according to claim 29, wherein, The amount of chelated surfactant in the mixed material is 0.01-10 wt%.

32. The method according to claim 31, wherein, The amount of chelated surfactant in the mixed material is 1.5-5 wt%.

33. The method according to claim 1, 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 / oil volume ratio of 400-3000; and / or liquid phase volume space velocity of 0.2-3 h -1 ; and / or The conditions for the three-stage hydrocracking include: a pressure of 2-8 MPa; and / or a liquid hourly space velocity of 0.2-3 h⁻¹. -1 ; and / or a temperature of 260-500℃; and / or a hydrogen-to-oil volume ratio of 400-3000.

34. The method according to claim 33, wherein, The conditions for the two-stage hydrorefining process include: a temperature of 210-250℃; and / or a pressure of 2.2-6.0 MPa; and / or a hydrogen-to-oil volume ratio of 600-2000; and / or a liquid hourly space velocity of 0.6-1.2 h⁻¹. -1 ; The conditions for the three-stage hydrocracking include: a pressure of 3.0-6.5 MPa; and / or a liquid hourly space velocity of 0.6-1.2 h⁻¹. -1 ; and / or a temperature of 260-480℃; and / or a hydrogen-to-oil volume ratio of 600-2000.

35. The method according to claim 1, wherein, The two-stage hydrorefining method involves the material passing through at least two layers of catalyst along the flow direction, with the pore size of the first support in the catalyst decreasing sequentially.

36. The method according to claim 1, wherein, The method is carried out in a low-quality aromatic distillate oil utilization unit, which includes a first-stage hydrogenation unit, a second-stage hydrorefining unit, a first purification unit, a third-stage hydrocracking unit, and a second purification unit arranged in series; wherein... The first-stage hydrogenation unit is used to convert dienes in inferior aromatic-rich distillate feedstock and reduce bromine value to obtain a first-stage hydrogenation product stream. The temperature of the product stream outlet of the two-stage hydrorefining unit is higher than that of the raw material inlet, and the temperature difference is not greater than 75°C. It is used to refine the product stream of the first-stage hydrorefining unit, so that the naphthalene series compounds are selectively generated into tetrahydronaphthalene series compounds and the sulfur and nitrogen of the refined product are less than 1 ppm by mass, thus obtaining the product stream of the two-stage hydrorefining unit. The first purification device is used to remove hydrogen sulfide from the product stream of the second-stage hydrorefining process to obtain a mixture of monocyclic aromatic hydrocarbons. The three-stage hydrocracking unit is used to crack a mixture of monocyclic aromatic hydrocarbons, so that tetrahydronaphthalene compounds are hydrocracking to generate BTX and LPG, and a three-stage hydrocracking product stream is obtained. The second purification unit is used to purify and separate the three-stage hydrocracking product streams to obtain BTX and LPG; The two-stage hydrorefining unit is filled with a catalyst for two-stage hydrorefining. The catalyst for two-stage hydrorefining includes a first active component, an active auxiliary agent, and a first support. The first active component contains Group VIB metals and Group VIII metals, and the active auxiliary agent contains lanthanide metals and / or Group VA elements. After pre-sulfurization, the catalyst for two-stage hydrorefining has 3-4 layers of sulfide active phase stacked in a tower-like structure.

37. The method of claim 36, wherein, The outlet of the monocyclic aromatic hydrocarbon mixture of the first purification unit is connected to the feed inlet of the second-stage hydrorefining unit via a circulation pipeline.

38. The method according to claim 36, wherein, The first purification device and the second purification device are each independently selected from at least one of the following: high-resolution tank, oil-water separator, stripping tower, low-resolution tank, and distillation tower; and / or The first active component contains at least one of Ni, Co, Fe, Pt and Pd, as well as Mo and / or W; the active additive is La, P, or Ce; the sulfide active phase has a stacking length of 3-6 nm, with edges, corners, and active sites fully exposed; The three-stage hydrocracking unit is loaded with a three-stage hydrocracking catalyst, which includes a second active component and a second support. The second active component contains a Group VIII metal, and the second support contains a layered HZSM-5 / MCM-41 molecular sieve.

39. The method according to claim 38, wherein, The second carrier also contains Ce oxide and Zr oxide.

40. The method according to claim 38 or 39, wherein, The second active component is Ni.

Images