A method for producing naphtha from a diesel hydrotreating unit

Abstract

This invention discloses a method for producing naphtha from a diesel hydrotreating unit. The method includes the following steps: (I) Feedstock oil and hydrogen enter a diesel hydrotreating reactor, sequentially passing through diesel hydrorefining reaction zone I, diesel hydrocracking reaction zone I, diesel hydrocracking reaction zone II, and diesel hydrorefining reaction zone II within the reactor for hydrotreating; (II) The reaction stream obtained in step (I) enters a separator for gas-liquid separation; (III) The liquid obtained in step (II) enters a fractionation system to obtain naphtha; wherein the catalyst packed in diesel hydrocracking reaction zone II is a high-temperature hydrocracking catalyst. Using this method allows existing diesel hydrotreating units to flexibly adapt to changes in feedstock and increase naphtha production, achieving stable long-term naphtha production. This method also avoids equipment modification and reduces investment.

Description

Technical Field

[0001] This invention relates to the field of petroleum hydrotreating technology, specifically to a method for producing naphtha using a diesel hydrotreating unit. Background Technology

[0002] The existing main processing routes for diesel fractions are hydrorefining, hydrotreating, and hydrocracking. While existing diesel hydrorefining and hydrotreating processes can be slightly modified to increase naphtha production, diesel hydrotreating units are affected by process conditions and feedstock requirements. Initial cracking performance is good, but with extended operating cycles, hydrocracking catalyst deactivation and activity decline rapidly, impacting long-term unit operation.

[0003] The main diesel hydrorefining technology used in refining enterprises employs a trickle-bed gas-liquid-solid three-phase reaction process, where the gas is hydrogen, the liquid is diesel feedstock, and the solid is the hydrorefining catalyst. The products are primarily refined diesel fractions and wide-range naphtha. Currently, naphtha production in hydrorefining units mainly relies on hydrocracking catalysts, which typically contain molecular sieves and require relatively low nitrogen content. However, existing diesel hydrorefining units are mostly medium- or low-pressure hydrorefining units, which somewhat affects denitrification capacity, resulting in a decline in the cracking performance of the hydrocracking catalyst after a period of operation.

[0004] US5114562A discloses a two-stage diesel hydrotreating process. The first stage uses a conventional hydrorefining catalyst to remove impurities such as sulfur and nitrogen from the feedstock diesel. The second stage uses a catalyst with high hydrorefining saturation activity for deep aromatic removal. This process has high hydrogen consumption, high investment, poor economic efficiency, and low naphtha fraction in the product.

[0005] CN1940030A discloses a hydrocracking method for increasing diesel production from high-nitrogen-content heavy feedstock. This method employs a two-stage hydrocracking process to maximize naphtha production. However, existing diesel hydrocracking units require modification, resulting in high investment costs and issues such as the impact of nitrogen on hydrocracking catalysis during long-term operation.

[0006] CN104611029A discloses a method for hydroconversion of catalytic cracking diesel. This method involves mixing catalytic diesel with hydrogen, first feeding it into a hydrorefining reactor for hydrorefining, and then into a hydrocracking reactor for hydrocracking. While this hydroconversion process produces high-octane gasoline, it has certain limitations, and the operating cycle cannot be guaranteed under high-nitrogen feedstock conditions.

[0007] CN111321005A discloses a hydrogenation process for producing ultra-low sulfur diesel fuel with low energy consumption and long cycle time. It separates the feedstock diesel fuel into light and heavy components, which then enter different reaction zones. The heavy components pass through the first reaction zone and then enter the second reaction zone with the light components, enabling long-cycle operation. However, this method cannot meet the requirement of producing large amounts of naphtha.

[0008] In summary, existing technologies mainly consider the need for impurity removal in diesel hydrorefining, which has certain limitations in addressing the issue of increased naphtha production. In particular, the decrease or disappearance of hydrocracking catalytic activity after long-term operation has resulted in insufficient consideration of the impact on increased naphtha production. Summary of the Invention

[0009] To address the shortcomings of existing technologies, the present invention aims to provide a method for producing naphtha from a diesel hydrotreating unit, particularly for existing diesel hydrotreating units in refining and chemical enterprises that employ trickle bed hydrotreating processes. This method enables existing diesel hydrotreating units to flexibly adapt to changes in feedstock and increase naphtha production, while maintaining stable naphtha production over long periods. Furthermore, this method avoids equipment modifications and reduces investment.

[0010] This invention provides a method for producing naphtha using a diesel hydrotreating unit, comprising the following:

[0011] (I) The feedstock oil and hydrogen enter the diesel hydrotreating reactor and pass through the diesel hydrorefining reaction zone I, diesel hydrocracking reaction zone I, diesel hydrocracking reaction zone II and diesel hydrorefining reaction zone II in sequence to carry out the hydrogenation reaction;

[0012] (II) The reaction stream obtained in step (I) enters the separator for gas-liquid separation;

[0013] (III) The liquid obtained in step (II) enters the fractionation system to obtain naphtha;

[0014] The catalyst packed in diesel hydrocracking reaction zone II is a high-temperature hydrocracking catalyst. The high-temperature hydrocracking catalyst uses alumina as a support, on which molecular sieves and active components are loaded. The active components are Group VIB metal sulfides and Group VIII metal sulfides. The molar ratio of the Group VIII metal sulfide species on the Group VIB metal sulfide wafers to the total Group VIII metal sulfides is 60%-100%. The molecular sieves are loaded on the outer surface of the catalyst.

[0015] Further, the diesel hydrotreating reactor described in step (I) is a trickle bed reactor. Two or more diesel hydrotreating reactors can be used in series. The conditions for the hydrogenation reaction in the diesel hydrotreating reactor are as follows: hydrogen partial pressure of 2.0 MPa-18.0 MPa, preferably 4.0 MPa-12.0 MPa; average reaction temperature of 280℃-450℃, preferably 320℃-430℃; and volume hourly space velocity of 0.1 h⁻¹. -1 -4.0h -1 0.5h is preferred -1 -2h -1 The hydrogen-to-oil volume ratio is 100:1-2000:1, preferably 350:1-1000:1.

[0016] Further, the initial boiling point of the feedstock oil in step (I) is 130℃-300℃, preferably 150℃-230℃, and the final boiling point is 350℃-400℃. The total aromatic content in the feedstock oil is 10%-50wt%, and the content of aromatics with two or more rings is 10%-30wt%, preferably 15%-25wt%. The feedstock oil is mainly straight-run diesel oil, and may be mixed with one or more diesel fractions from secondary processed oils such as coking diesel oil, catalytic diesel oil, residue-added diesel oil, wax-added diesel oil, heavy aromatic oil, and residue oil from aromatic extraction units.

[0017] Furthermore, the hydrogen mentioned in step (I) can be recycled hydrogen and fresh hydrogen. The feedstock oil can first pass through a heat exchanger, and then be mixed with the recycled hydrogen and fresh hydrogen mixture heated by a heater before entering the diesel hydrotreating reactor.

[0018] Furthermore, both diesel hydrorefining reaction zone I and diesel hydrorefining reaction zone II are filled with conventional diesel hydrorefining catalysts (the diesel hydrorefining catalysts filled in the two can be the same or different). The diesel hydrorefining catalyst generally uses refractory porous oxides as supports, such as alumina, silica, titanium dioxide, and composite oxides or mixed oxide supports of several elements. Generally, non-acidic or weakly acidic materials are used as supports, and the hydrogenation activity combination of the catalyst is W-Ni, Mo-Ni, or W-Mo-Ni. The content of the hydrogenation active component, calculated as oxide, is generally 10wt%-50wt%, preferably 15wt%-40wt%. Among them, the nickel oxide content is 1.5wt%-10wt%, preferably 3.5wt%-6wt%. The diesel hydrorefining catalyst can be a suitable commercial catalyst selected according to the needs of the process flow, or it can be prepared by existing methods, or it can be a regenerated catalyst after the deactivated catalyst has been regenerated. For example, it can be FHUDS-8, FHUDS-7, etc.

[0019] Furthermore, the diesel hydrocracking reaction zone I is filled with a conventional hydrocracking catalyst.

[0020] Furthermore, the hydrocracking catalyst generally uses alumina as a support and is mixed with one or more of β-zeolites, Y-type zeolites, ZSM series zeolites, USY zeolites, and SAPO series zeolites. Based on the weight of the catalyst, the mass ratio of alumina to zeolite is 15:1-2:1, preferably 10:1-5:1; simultaneously, a hydrogenation active component is supported. The hydrogenation active component of the diesel hydrocracking catalyst is one or more of W, Mo, Ni, and Co, and its content as oxides is generally 3wt%-50wt%, preferably 15wt%-45wt%. The diesel hydrocracking catalyst can be a suitable commercial catalyst selected according to the needs of the process flow, or it can be prepared by existing methods, or it can be a regenerated catalyst after deactivated catalyst regeneration. For example, FC-50 produced by the Fushun Branch of Sinopec Catalyst Company.

[0021] Furthermore, based on the total weight of the high-temperature hydrocracking catalyst, the molecular sieve accounts for 1 wt%-20 wt%, preferably 1.5 wt%-12 wt%, more preferably 4 wt%-10 wt%; the Group VIB metal sulfides, calculated as sulfides, account for 10 wt%-30 wt%, preferably 15 wt%-28 wt%, and the Group VIII metal sulfides, calculated as sulfides, account for 2 wt%-10 wt%, preferably 4 wt%-8 wt%.

[0022] Furthermore, when analyzed by CO-FTIR, the high-temperature hydrocracking catalyst has a molar ratio of Group VIII metal-Group VIB metal-S phase to Group VIII metal (i.e., the molar ratio of the Group VIII metal sulfide species on the Group VIB metal sulfide wafer to the total amount of Group VIII metal sulfides) of 60%-100%, preferably 65%-90%, more preferably 70%-90%, and most preferably 80%-90%.

[0023] Furthermore, the molecular sieve is selected from at least one of Y-type molecular sieve, ZSM-5 molecular sieve, β-type molecular sieve and MCM-41 molecular sieve, preferably Y-type molecular sieve.

[0024] Furthermore, the molecular sieve supported on the outer surface of the catalyst is a molecular sieve supported on a Group VIB metal sulfide, and / or a molecular sieve supported on a Group VIII metal sulfide, and / or a molecular sieve supported on alumina.

[0025] Furthermore, the Group VIB metal sulfide is molybdenum sulfide and / or tungsten sulfide, and the Group VIII metal sulfide is nickel sulfide and / or cobalt sulfide.

[0026] Furthermore, the preparation method of the high-temperature hydrocracking catalyst includes the following steps:

[0027] (1) The alumina support was impregnated with an impregnation solution containing a Group VIB metal salt, and the catalyst precursor A was obtained by drying and sulfidation.

[0028] (2) The catalyst precursor of step (1) is impregnated with an impregnation solution containing group VIII metal salts and organic additives, then dried in an inert atmosphere, and then sulfided to obtain catalyst precursor B.

[0029] (3) The catalyst precursor B from step (2) is subjected to hydrothermal treatment with the molecular sieve precursor, and then filtered, washed and dried and calcined in an inert atmosphere to obtain a high-temperature hydrocracking catalyst.

[0030] Further, the impregnation solution containing Group VIB metals in step (1) is a phosphate or ammonium salt solution of a Group VIB metal, the preparation method of which is well known to those skilled in the art, and employs equal-volume impregnation or supersaturated impregnation methods. The Group VIB metal is preferably Mo and / or W.

[0031] Furthermore, the drying conditions for step (1) are: drying temperature 90-200℃, drying time 3-6 hours.

[0032] Further, the vulcanization treatment in step (1) is either dry vulcanization or wet vulcanization. The dry vulcanizing agent is hydrogen sulfide, and the wet vulcanizing agent is one or two of carbon disulfide, dimethyl disulfide, methyl sulfide, and n-butyl sulfide; the vulcanization pressure is 3.2-6.4 MPa, the vulcanization temperature is 250-400℃, and the vulcanization time is 4-12 h.

[0033] Furthermore, the impregnation solution containing Group VIII metal salt in step (2) is at least one of the group VIII metal nitrate, acetate or sulfate solution, etc., and can be impregnated by equal volume. The group VIII metal is Ni and / or Co.

[0034] Further, the organic auxiliary agent in step (2) is at least one of alcohols or organic acids containing hydroxyl and / or carboxyl groups, wherein the number of carbon atoms is 3-10. Specifically, it is selected from at least one of ethylene glycol, glycerol, butanediol, pentanediol, acetic acid, citric acid, malonic acid, succinic acid, and glutaric acid.

[0035] Further, the inert atmosphere in step (2) is one or more of N2 and inert gases; the drying temperature in step (2) is 20-90℃, and the drying time is 4-16 hours.

[0036] Further, the vulcanization treatment in step (2) is either dry vulcanization or wet vulcanization. The dry vulcanizing agent is hydrogen sulfide, and the wet vulcanizing agent is one or two of carbon disulfide, dimethyl disulfide, methyl sulfide, and n-butyl sulfide; the vulcanization pressure is 3.2-6.4 MPa, the vulcanization temperature is 250-400℃, and the vulcanization time is 4-12 h.

[0037] Further, the molecular sieve precursor in step (3) is a gel formed by mixing a silicon source and / or an aluminum source, a precipitant, a template agent, and water. Its preparation method is well known to those skilled in the art, and the molecular sieve is formed using a precipitation method or a sol-gel method. The silicon source is selected from one or more of sodium silicate, tetraethyl orthosilicate, silica sol, and chromatographic silica gel; the aluminum source is selected from one or more of sodium aluminate, aluminum hydroxide, and boehmite; the precipitant is selected from at least one of sodium hydroxide, ammonia, and potassium hydroxide; and the template agent is selected from one or more of hexadecyltrimethylammonium bromide, ethylenediamine, n-butylamine, tetrapropylammonium bromide, ethanol, tetraethylammonium hydroxide, tetraethylammonium bromide, triethylamine, di-n-propylamine, diisopropylamine, and methylcellulose.

[0038] Further, the hydrothermal treatment conditions in step (3) are: temperature 90-200℃, preferably 130-200℃, pressure 0.1-2.0MPa, pH 7.5-9.0, and time 5-48 hours. After hydrothermal treatment, the molecular sieve precursor crystallizes into a molecular sieve and is loaded onto the outer surface of the catalyst.

[0039] Further, the inert atmosphere in step (3) is one or more of N2 and inert gases; the drying temperature in step (3) is 20-90℃ and the drying time is 4-16 hours; the calcination temperature is 300-500℃ and the calcination time is 2-5 hours.

[0040] Furthermore, in step (3), after hydrothermal treatment, an ammonium ion exchange process is required, which can be carried out using conventional methods in the art. For example, an ammonium salt aqueous solution can be used for ammonium ion exchange. The ammonium salt can be one or more of ammonium chloride, ammonium nitrate, ammonium sulfate, ammonium acetate, etc. The concentration of the ammonium salt aqueous solution is 0.05-3.0 mol / L, the exchange temperature is 55℃-90℃, the time for a single exchange is 1-15 h, and the number of exchanges is 1-6 times.

[0041] The volume ratio of catalysts packed in the diesel hydrorefining reaction zone I, diesel hydrocracking reaction zone I, diesel hydrocracking reaction zone II and diesel hydrorefining reaction zone II in the diesel hydrorefining reactor is 3-39:2-8:2-3:1, preferably 5-14:2-3:2-3:1.

[0042] Furthermore, in step (II), the separator is a conventional gas-liquid separator used in the diesel hydrotreating field. The gas separated by the separator undergoes gas desulfurization before being mixed with fresh hydrogen and feedstock oil. The gas desulfurization can be performed using conventional desulfurization methods used in the hydrotreating field.

[0043] Furthermore, in step (III), the fractionation system is a conventional fractionation system in the field of diesel hydrotreating.

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

[0045] 1. This invention fully utilizes existing diesel hydrotreating units without requiring process modifications. It boasts strong feedstock adaptability and can process low-quality secondary-processed diesel fractions. Particularly important is the initial stage of operation, where it is desirable to avoid excessively high hydrocracking activity, which can lead to a loss of hydrocracking performance of some cracking agents and a decrease in overall hydrocracking activity after prolonged operation. The hydrotreating reaction process of this invention effectively solves the problem of stable naphtha production during long-term operation of diesel hydrotreating units, ensuring the long-term operation of the unit.

[0046] 2. The hydrotreating reaction of diesel fractions is a strongly exothermic reaction. The reaction medium, in a hydrogen-rich state, contacts metal sulfides to remove impurities and undergo hydrogenation. Carbon deposits are generated during the hydrotreating process, primarily due to the high reaction temperature, but also accompanied by certain polymerization and condensation side reactions. With prolonged operation, these carbon deposits gradually accumulate on the hydrotreating catalyst, covering the active sites and causing a continuous decline in catalyst activity. To maintain catalyst activity, the reaction temperature increases with the duration of the reaction. Therefore, optimization of the configuration is crucial to consider the reaction characteristics of different types of catalysts at different reaction temperatures. This optimization aims to achieve a matching of catalyst activities across different types of catalysts as the reaction temperature increases, thereby extending the operating cycle and stabilizing naphtha production.

[0047] 3. This invention sets up different hydrocracking reaction zones and matches them with different types of hydrocracking catalysts. As the temperature increases, a stepped effect is formed, progressively increasing the cracking performance and achieving the goal of long-term stable production of naphtha. In the high-temperature hydrocracking catalyst of this invention, molecular sieves are added through a special method, and the molecular sieves are loaded on the outer surface of the catalyst rather than mixed with the support. This is beneficial in three ways: firstly, it increases the contact area between the molecular sieves and the active metal, and secondly, it facilitates the loading of the molecular sieves onto the active sites of the active metal, fully utilizing the ability of the molecular sieves to provide H protons and improving the hydrogenation activity of the catalyst; thirdly, it allows for precise control of the hydrogenation saturation and ring-opening / chain-severing activity of polycyclic aromatic hydrocarbons by controlling the type and content of the molecular sieves, resulting in high catalyst flexibility; and fourthly, it increases the utilization rate of the molecular sieves, thus reducing the amount of molecular sieves used and lowering the cost of the catalyst.

[0048] By selecting and loading conventional hydrocracking catalysts and high-temperature hydrocracking catalysts in a graded manner in the hydrocracking reaction zone, it is possible to achieve the following: during the initial low-temperature reaction at the start-up stage, the molecular sieve catalyst mainly exerts its cracking activity, while the high-temperature hydrocracking catalyst exerts little or no cracking activity; towards the end of operation, after the molecular sieve catalyst has partially lost its activity, the acidic centers of the high-temperature hydrocracking catalyst are activated at higher reaction temperatures to replenish the hydrocracking activity, thereby achieving the goal of stabilizing naphtha yield and effectively changing the problem of decreased hydrocracking activity of the catalyst system at the end of operation.

[0049] 4. Using the method provided by this invention, while producing naphtha stably, the diesel product is also stable and of high quality.

[0050] 5. The method provided by this invention adjusts the catalyst gradation system without requiring any major adjustments to the overall equipment, thus reducing the cost of equipment modification. At the same time, it does not change much from the original operating steps, making it more suitable for operation. Attached Figure Description

[0051] Figure 1 This is a schematic diagram of the process used in embodiments 4-6 of the present invention;

[0052] The main reference numerals in the attached drawings are as follows: 1-Federal feedstock, 2-New hydrogen, 3-Circulating hydrogen, 4-Diesel hydrotreating reactor, 6-Cold high-pressure separator, 9-Circulating hydrogen compressor, 11-Cold low-pressure separator, 13-Fraction system. Detailed Implementation

[0053] The method disclosed in this invention will now be described in more detail with reference to the accompanying drawings and specific embodiments. Figure 1 Many essential equipment items, such as oil pumps, valves, and heating furnaces, have been omitted.

[0054] like Figure 1 As shown, the feedstock diesel 1, fresh hydrogen 2, and recycled hydrogen 3 are mixed and enter the reactor from the top of the diesel hydrotreating reactor 4. The material sequentially passes through the hydrorefining reaction zone I, the hydrocracking reaction zone I, the hydrocracking reaction zone II, and the hydrorefining reaction zone II, reacting with the hydrorefining catalyst, the hydrocracking catalyst, the high-temperature hydrocracking catalyst, and the hydrorefining catalyst. It then flows out from the bottom of the reactor and enters the cold high-pressure separator 6 through pipeline 5. The liquid phase passes through pipeline 8 and enters the cold low-pressure separator 11, and then enters the fractionation system 13 through pipeline 12. The recycled hydrogen separated by the cold high-pressure separator 6 enters the recycled hydrogen compressor 9 through pipeline 7 and mixes with the fresh hydrogen and feedstock oil through pipeline 10.

[0055] The present invention will be further described below with reference to preferred embodiments. It should be understood that the preferred embodiments described herein are for illustration and explanation only and are not intended to limit the present invention.

[0056] In the embodiments and comparative examples of this invention, the hydrorefining catalysts used in hydrorefining reaction zone I and hydrorefining reaction zone II were developed by the Fushun Petrochemical Research Institute, and both were FHUDS-8 produced by Sinopec Catalyst Company Fushun Branch. The hydrocracking catalyst loaded in hydrocracking reaction zone I was developed by the Fushun Petrochemical Research Institute, specifically FC-50 produced by Sinopec Catalyst Company Fushun Branch. The FHUDS-8 catalyst uses alumina as a support and Mo-Ni as the active metal component; the FC-50 catalyst uses alumina and molecular sieve as supports and undergoes acid modification, with W-Ni as the active metal component.

[0057] The diesel hydrogenation reactor used in the comparative examples of this invention is a trickle bed reactor.

[0058] Specific embodiments and comparative examples adopt Figure 1 The hydrogenation process flow.

[0059] The catalyst composition provided by this invention is characterized by a combination of inductively coupled plasma (ICP) and XPS spectroscopy. First, the total content of Group VIB metals and Group VIII metals in the catalyst is characterized by ICP. Then, the content of metal elements in different valence states in the catalyst is quantitatively characterized by XPS spectroscopy.

[0060] This invention uses CO-FTIR (carbon monoxide in situ infrared spectroscopy) to analyze the content of Group VIII metal oxides on Group VIB metal oxide wafers. The CO-FTIR measurement conditions include: the catalyst is ground, pressed into a Φ13mm self-supporting sheet, and placed on an in-situ sample holder. Sulfidation treatment with 3.0% H2S at 320℃ for 3 hours is performed, followed by cooling to room temperature in an H2S atmosphere, then purification at 300℃ under vacuum for 2 hours, and CO adsorption is carried out under liquid nitrogen cooling conditions. A small amount of CO gas is introduced into the in-situ cell, and after adsorption equilibrium for 30 minutes, desorption is performed to 10... -4 Pa. Infrared spectra were collected before and after CO adsorption; the difference spectrum between the two is the infrared spectrum of CO adsorption by the catalyst. The experiment used a Nicolet 6700 Fourier transform infrared spectrometer, with 32 scans and a resolution of 4 cm⁻¹. -1 4000-650cm -1 Measurements were taken using an MCT / A detector.

[0061] Example 1

[0062] (1) The ammonium heptamolybdate solution was impregnated into the alumina support, and then dried at 90°C for 4 hours. Then, it was sulfided with hydrogen containing 1.5% H2S at a sulfidation temperature of 290°C, a sulfidation pressure of 5.0 MPa, and a sulfidation time of 5 h. Then, it was cooled to room temperature in a N2 atmosphere to obtain catalyst precursor A.

[0063] (2) The nickel nitrate and glycerol solution was impregnated into the catalyst precursor A prepared in step (1), and then dried at 80°C for 5 h in a nitrogen atmosphere. Then, it was sulfided with hydrogen containing 1.5% H2S at a sulfidation temperature of 310°C, a sulfidation pressure of 5.0 MPa, and a sulfidation time of 5 h. Finally, it was cooled to room temperature in a N2 atmosphere to obtain catalyst precursor B.

[0064] (3) Sodium hydroxide, silica sol, sodium aluminate and ethylenediamine were added to deionized water, with the molar ratio of each component being n(SiO2):n(Al2O3):n(Na2O):n(ethylenediamine):n(H2O) = 11:2:5:3:190. The mixture was stirred until a homogeneous sol was formed, which is the precursor of the molecular sieve. It was then mixed with the catalyst precursor B prepared in step (2), and then hydrothermally treated for 8 hours at 150℃, 2.0MPa and pH=8.1. The resulting material was placed in a 0.5mol / L NH4Cl aqueous solution and mixed at a liquid-to-solid volume ratio of 3:1. The exchange temperature was 80℃, the single exchange time was 3 hours, and the total number of exchanges was 3. The mixture was then filtered, washed three times with deionized water, and dried at 90℃ for 5 hours in a nitrogen atmosphere to obtain the high-temperature hydrocracking catalyst C-1.

[0065] The weight percentages of the components in the high-temperature hydrocracking catalyst C-1 are as follows: MoS2 21%, NiS 3.9%, Y molecular sieve 6.5%, and the remainder is alumina support. Analysis showed that the content of Group VIII metal oxides on the Group VIB metal oxide wafer (Ni-Mo-S content) was 85%.

[0066] Example 2

[0067] (1) The ammonium heptamolybdate solution was impregnated into the alumina support, and then dried at 90°C for 6 hours. Then, it was sulfided with hydrogen containing 1.5% H2S at a sulfidation temperature of 280°C, a sulfidation pressure of 5.0 MPa, and a sulfidation time of 5 hours. Then, it was cooled to room temperature in a N2 atmosphere to obtain catalyst precursor A.

[0068] (2) The nickel nitrate and malonic acid solution was impregnated into the catalyst precursor A prepared in step (1), and then dried at 80°C for 6 hours in a nitrogen atmosphere. Then, a sulfidation treatment was performed at a sulfidation temperature of 350°C, a sulfidation pressure of 5.0 MPa, and a sulfidation time of 5 hours. Finally, the mixture was cooled to room temperature in a N2 atmosphere to obtain catalyst precursor B.

[0069] (3) Sodium hydroxide, silica sol, sodium aluminate and ethylenediamine were added to deionized water, with the molar ratio of each component being n(SiO2):n(Al2O3):n(Na2O):n(n-butylamine):n(H2O) = 25:2:7:4:200. The mixture was stirred to form a homogeneous sol, which is the precursor of the molecular sieve. It was then mixed with the catalyst precursor B prepared in step (2) and hydrothermally treated for 10 h at 150 °C, 1.5 MPa and pH = 8.0. The resulting material was placed in a 0.5 mol / L NH4Cl aqueous solution and mixed at a liquid-to-solid volume ratio of 3:1. The exchange temperature was 80 °C and the single exchange time was 3 h. The total number of exchanges was 3. The mixture was then filtered, washed three times with deionized water, and dried at 90 °C for 4 h in a nitrogen atmosphere to obtain the high-temperature hydrocracking catalyst C-2.

[0070] The weight percentages of the components in the high-temperature hydrocracking catalyst C-2 are as follows: MoS2 25%, NiS 4.5%, Y molecular sieve 7.0%, and the remainder is alumina support. Analysis showed that the content of Group VIII metal oxides on the Group VIB metal oxide wafer (Ni-Mo-S content) was 83%.

[0071] Example 3

[0072] (1) The ammonium heptamolybdate solution was impregnated into the alumina support, and then dried at 150°C for 3 hours. Then, it was sulfided with hydrogen containing 1.5% H2S at a sulfidation temperature of 280°C, a sulfidation pressure of 5.0 MPa, and a sulfidation time of 5 hours. Then, it was cooled to room temperature in a N2 atmosphere to obtain catalyst precursor A.

[0073] (2) The catalyst precursor A prepared in step (1) is impregnated with nickel nitrate and acetic acid solution, then dried at 80°C for 5 hours in a nitrogen atmosphere, and then subjected to sulfidation treatment at a temperature of 330°C, a pressure of 5.0 MPa, and a time of 5 hours. Finally, it is cooled to room temperature in a N2 atmosphere to obtain catalyst precursor B.

[0074] (3) Dissolve sodium aluminate and sodium hydroxide in deionized water, then add tetraethylammonium bromide, stir vigorously, slowly add silica sol, and age for 3 hours. The molar ratio of each component is n(SiO2):n(Al2O3):n(Na2O):n(tetraethylammonium bromide):n(H2O) = 26:2:9:7:280 to form a molecular sieve precursor. Then mix it with the catalyst precursor B prepared in step (2), and then hydrothermally treat it for 10 hours at 150℃, 2.0MPa, and pH=8.5. The resulting material is placed in a 0.5mol / L NH4Cl aqueous solution and mixed according to a liquid-solid volume ratio of 3:1. The exchange temperature is 80℃, the single exchange time is 3 hours, and the total number of exchanges is 3. Then filter, wash with deionized water three times, and dry at 90℃ in a nitrogen atmosphere for 5 hours to obtain high-temperature hydrocracking catalyst C-3.

[0075] The weight percentages of the components in the high-temperature hydrocracking catalyst C-3 are as follows: MoS2 25%, NiS 4.2%, Y molecular sieve 7.5%, and the remainder is alumina support. Analysis showed that the content of Group VIII metal oxides on the Group VIB metal oxide wafers (Ni-Mo-S content) was 82%.

[0076] Example 4

[0077] use Figure 1 The hydrogenation process involves loading diesel hydrorefining catalyst FHUDS-8, hydrocracking catalyst FC-50, and high-temperature hydrocracking catalyst C-1 prepared in Example 1 into the reactor from top to bottom. The main operating conditions are: total space velocity (SHSV) of 1.1 h⁻¹. -1 The reaction pressure was 8.0 MPa, the hydrogen-to-oil volume ratio was 500:1, and the reaction results are shown in Table 2.

[0078] Example 5

[0079] use Figure 1 The hydrogenation process flow differs from that of Example 4 in that the high-temperature hydrocracking catalyst used is the catalyst prepared in Example 2. The reaction results are shown in Table 2.

[0080] Example 6

[0081] use Figure 1 The hydrogenation process flow differs from that of Example 4 in that the high-temperature hydrocracking catalyst used is the catalyst prepared in Example 3. The reaction results are shown in Table 2.

[0082] Comparative Example 1

[0083] use Figure 1The hydrogenation process flow differs from that in Example 4 in that both hydrocracking reaction zone I and hydrocracking reaction zone II are filled with hydrocracking catalyst FC-50. The reaction results are shown in Table 2.

[0084] Comparative Example 2

[0085] use Figure 1 The hydrogenation process flow differs from Example 4 in that the reactor is loaded from top to bottom with diesel hydrorefining catalyst FHUDS-8, high-temperature hydrocracking catalyst C-1 prepared in Example 1, hydrocracking catalyst FC-50, and diesel hydrorefining catalyst FHUDS-8. That is, the reactants first contact catalyst C-1 and then hydrocracking catalyst FC-50. The reaction results are shown in Table 2.

[0086] Meanwhile, the feedstock oils used in the embodiments and comparative examples of the present invention are shown in Table 1. The main operating conditions and gasoline product properties in the embodiments and comparative examples of the present invention are shown in Table 2.

[0087] Table 1 Properties of the Crude Oil

[0088] Blended diesel <![CDATA[Density (20 °C), g / cm 3 > 0.8756 Distillation range, °C (ASTM D86) Initial boiling point / 10% 170 / 195 50％ / 90％ 293 / 346 95% / Dry point (final boiling point) 363 / 375 Sulfur, μg / g 9655 Nitrogen, μg / g 356

[0089] Table 2. Main operating conditions and diesel product properties at different times

[0090]

[0091]

[0092] The data results from the examples and comparative examples show that the present invention uses a conventional diesel hydrorefining unit and a trickle bed reactor process for graded cracking. By grading the catalyst, it achieves matching of cracking activities at different temperatures, thereby ensuring that the naphtha yield does not decrease and the product quality remains stable during long-term operation, reducing unit energy consumption and achieving the goal of long-term operation. This process, by optimizing the use of catalysts, reduces the cost of unit modification and operational risks while keeping the operation steps unchanged, and has significant practical application advantages.

Claims

1. A method for producing naphtha in a diesel hydroprocessing unit, comprising the following steps: (I) feeding a feedstock oil and hydrogen into a diesel hydroprocessing reactor, and sequentially passing through a diesel hydrofining reaction zone I, a diesel hydrocracking reaction zone I, a diesel hydrocracking reaction zone II and a diesel hydrofining reaction zone II in the diesel hydroprocessing reactor to perform a hydrogenation reaction; (II) passing the reaction stream obtained in step (I) into a separator to perform a gas-liquid separation; (III) passing the liquid obtained in step (II) into a fractionation system to obtain naphtha; wherein The diesel hydrocracking reaction zone II is packed with a high-temperature hydrocracking catalyst, the high-temperature hydrocracking catalyst has an alumina carrier, a molecular sieve and an active component supported on the carrier, the active component is a Group VIB metal sulfide and a Group VIII metal sulfide; when the high-temperature hydrocracking catalyst is analyzed by CO-FTIR, the molar ratio of the species of the Group VIII metal sulfide on the Group VIB metal sulfide wafer to the total amount of the Group VIII metal sulfide is 60%-100%; the molecular sieve is supported on the Group VIB metal sulfide, and / or the molecular sieve is supported on the Group VIII metal sulfide, and / or the molecular sieve is supported on the alumina; The molecular sieve accounts for 1wt%-20wt%, the Group VIB metal sulfide accounts for 10wt%-30wt% and the Group VIII metal sulfide accounts for 2wt%-10wt% based on the total weight of the high-temperature hydrocracking catalyst; The high-temperature hydrocracking catalyst is prepared by the following method, comprising: (1) impregnating an alumina carrier with an impregnation solution containing a Group VIB metal salt, and then performing drying and sulfidation treatment to obtain a catalyst precursor A; (2) impregnating the catalyst precursor of step (1) with an impregnation solution containing a Group VIII metal salt and an organic auxiliary agent, and then performing drying in an inert atmosphere, and then performing sulfidation to obtain a catalyst precursor B; (3) performing hydrothermal treatment on the catalyst precursor B of step (2) and a molecular sieve precursor, and then performing filtration, washing, drying in an inert atmosphere and calcination to obtain a high-temperature hydrocracking catalyst.

2. The method of claim 1, wherein, The conditions of the hydrogenation reaction in step (I) are as follows: hydrogen partial pressure is 2.0 MPa to 18.0 MPa, average reaction temperature is 280°C to 450°C, volume space velocity is 0.1 h -1 -4.0 h -1 , and hydrogen to oil volume ratio is 100:1 to 2000:

1.

3. The method of claim 2, wherein, The conditions for the hydrogenation reaction in step (I) are as follows: hydrogen partial pressure is 4.0 MPa to 12.0 MPa; average reaction temperature is 320°C to 430°C; volume space velocity is 0.5 h -1 -2 h -1 ; hydrogen to oil volume ratio is 350:1 to 1000:

1.

4. The method of claim 1, wherein, The initial boiling point of the feedstock oil in step (I) is 130°C-300°C, and the final boiling point is 350°C-400°C; the total content of aromatic hydrocarbons in the feedstock oil is 10wt%-50wt%, and the content of aromatic hydrocarbons with two or more rings is 10wt%-30wt%.

5. The method of claim 4, wherein, The initial boiling point of the feedstock oil in step (I) is 150°C-230°C, and the content of aromatic hydrocarbons with two or more rings in the feedstock oil is 15wt%-25wt%.

6. The method of claim 1, wherein, The diesel hydrofining reaction zone I and the diesel hydrofining reaction zone II are both packed with a diesel hydrofining catalyst, and the diesel hydrocracking reaction zone I is packed with a hydrocracking catalyst.

7. The method of claim 1, wherein, The molecular sieve accounts for 1.5wt%-12wt%, the Group VIB metal sulfide accounts for 15wt%-28wt% and the Group VIII metal sulfide accounts for 4wt%-8wt% based on the total weight of the high-temperature hydrocracking catalyst.

8. The method of claim 7, wherein, The molecular sieve accounts for 4wt%-10wt% based on the total weight of the high-temperature hydrocracking catalyst.

9. The method of claim 1, wherein, The Group VIB metal sulfide is molybdenum sulfide or / and tungsten sulfide, and the Group VIII metal sulfide is nickel sulfide or / and cobalt sulfide.

10. The method according to claim 1 or 7 or 8, characterized in that, The molecular sieve is selected from at least one of Y-type molecular sieve, ZSM-5 molecular sieve, beta-type molecular sieve and MCM-41 molecular sieve.

11. The method of claim 1, wherein, The molar ratio of the Group VIII metal sulfide species on the Group VIB metal sulfide crystalline sheet to the total amount of the Group VIII metal sulfide is 65%-90% when the high-temperature hydrocracking catalyst is analyzed by CO-FTIR.

12. The method of claim 11, wherein, The molar ratio of the Group VIII metal sulfide species on the Group VIB metal sulfide crystalline sheet to the total amount of the Group VIII metal sulfide is 70%-90%.

13. The method of claim 12, wherein, The molar ratio of the Group VIII metal sulfide species on the Group VIB metal sulfide crystalline sheet to the total amount of the Group VIII metal sulfide is 80%-90%.

14. The method of claim 1, wherein, The volume ratio of the catalysts loaded in the diesel hydrofining reaction zone I, the diesel hydrocracking reaction zone I, the diesel hydrocracking reaction zone II and the diesel hydrofining reaction zone II is 3-39: 2-8: 2-3:

1.

15. The method of claim 14, wherein, The volume ratio of the catalysts loaded in the diesel hydrofining reaction zone I, the diesel hydrocracking reaction zone I, the diesel hydrocracking reaction zone II and the diesel hydrofining reaction zone II is 5-14: 2-3: 2-3: 1.

Images