Process and system for adsorptive desulfurization and upgrading of oil

Abstract

The present application relates to the technical field of petroleum chemical industry, and discloses an oil adsorption desulfurization and modification method and system.The method comprises the following steps: (1) reducing material containing fresh catalyst in the presence of hydrogen to obtain active catalyst; (2) contacting oil, reaction hydrogen and the active catalyst in a moving bed reactor to carry out adsorption desulfurization and modification reaction of the oil, and obtaining modified product and post-reaction catalyst; (3) separating the modified product to obtain modified product; (4) carrying out one-stage regeneration of the post-reaction catalyst to obtain one-stage regenerated catalyst; (5) carrying out two-stage regeneration of the one-stage regenerated catalyst to obtain regenerated catalyst; and (6) returning the regenerated catalyst to step (1) for recycling.The oil adsorption desulfurization and modification method provided by the present application has low reaction pressure, small hydrogen / oil ratio, low hydrogen consumption, small operation difficulty, high yield of modified product, small coke yield and suitability for industrial popularization.

Description

Technical Field

[0001] This invention relates to the field of petrochemical technology, specifically to a method and system for oil adsorption desulfurization and upgrading. Background Technology

[0002] With increasing environmental awareness, stricter sulfur limits for oil products are being introduced internationally. Reducing the sulfur content of oil products has become a necessary treatment method for efficient oil utilization. Taking marine fuel oil as an example, before the implementation of the IMO sulfur limit policy, the marine fuel oil market was mainly dominated by high-sulfur marine residue fuel oil and distillate marine fuel oil, with high-sulfur marine residue fuel oil accounting for about 85% and distillate marine fuel oil accounting for about 15%. After the sulfur limit policy was implemented on January 1, 2020, low-sulfur marine residue fuel oil accounted for about 45% of the total consumption.

[0003] CN105694949A discloses a method for gasoline adsorption desulfurization, which uses a gasoline adsorption desulfurizing agent to adsorb and desulfurize catalytically hydrogenated gasoline in a fixed-bed reactor. This method can reduce the octane number while increasing the desulfurization rate of gasoline adsorption desulfurization, achieving a desulfurization rate of over 80%. However, the desulfurization effect still needs further improvement.

[0004] CN104560138A discloses a fluidized bed heavy oil hydrotreating method, employing a mixed hydrotreating catalyst with two bimodal pore structures. Under heavy oil hydrotreating conditions, the heavy oil feedstock and hydrogen react in a fluidized bed reactor to achieve hydrodesulfurization, hydrodemetallization, residual carbon conversion, and asphaltene conversion. However, this method suffers from problems such as high reaction pressure, high hydrogen-to-oil ratio, high hydrogen consumption, and high operating costs.

[0005] CN1473910A discloses a diesel adsorption desulfurization method, in which diesel is contacted with an amorphous alloy adsorbent with nickel as the main active component in a fluidized bed or slurry bed reactor under conditions of room temperature -150℃ and atmospheric pressure -2.0MPa. This method achieves a desulfurization rate of over 60% by weight. While this method is effective for desulfurizing diesel, it is not suitable for treating heavy oils with high viscosity. Summary of the Invention

[0006] The purpose of this invention is to overcome the problems of high reaction pressure, high hydrogen-to-oil ratio, and high energy and hydrogen consumption in existing hydrodesulfurization processes, and to provide an oil adsorption desulfurization and upgrading method and system.

[0007] To achieve the above objectives, a first aspect of the present invention provides a method for oil adsorption desulfurization and upgrading, wherein the method includes the following steps:

[0008] (1) In the presence of hydrogen, the material containing fresh catalyst is reduced to obtain an active catalyst;

[0009] (2) The oil, reaction hydrogen and active catalyst are contacted in a moving bed reactor to carry out the adsorption desulfurization and upgrading reaction of the oil to obtain the upgrading product and the catalyst after the reaction.

[0010] (3) Separate the modified product to obtain the modified product;

[0011] (4) The catalyst after the reaction is regenerated to obtain a regenerator;

[0012] (5) The first-stage regenerator is subjected to a second-stage regeneration to obtain a regenerated catalyst;

[0013] (6) The regenerated catalyst is returned to the material from step (1) for recycling.

[0014] A second aspect of the present invention provides an oil adsorption desulfurization and upgrading system, the system comprising a reducer 1, a moving bed reactor 2, a separation tower 3, a primary regenerator 4, and a secondary regenerator 5;

[0015] The reducer 1 is connected end-to-end to the moving bed reactor 2, the first stage regenerator 4, and the second stage regenerator 5, and the separation tower 3 is connected to the moving bed reactor 2.

[0016] The beneficial technical effects achieved by the present invention through the above technical solution are as follows:

[0017] 1) The oil adsorption desulfurization and upgrading method provided by the present invention can perform adsorption desulfurization and upgrading treatment on oil in a moving bed reactor, and can realize continuous production.

[0018] 2) The oil adsorption desulfurization and upgrading method provided by the present invention performs two-stage regeneration treatment on the deactivated catalyst, which can improve the activity of the regenerated catalyst and greatly reduce the investment cost and production and operation cost of the equipment.

[0019] 3) The oil adsorption desulfurization and upgrading method provided by the present invention preferably uses oil washing liquid to clean the catalyst after the reaction. On the one hand, it can remove the oil adhering to the surface of the deactivated catalyst and the oil adsorbed in the pores, thereby reducing the emission of carbon dioxide during the regeneration process. On the other hand, it can maximize the recovery of the upgraded products.

[0020] 4) The oil adsorption desulfurization and upgrading method provided by the present invention preferably involves reactivating a portion of the first-stage regenerator, which can further remove metal ions attached to the catalyst surface and pores, thereby improving the activity of the regenerated catalyst.

[0021] 5) The oil adsorption desulfurization reforming method provided by the present invention separates the circulating hydrogen from the reforming products and recycles it back into the adsorption desulfurization reforming reaction, which can improve the utilization rate of hydrogen and reduce hydrogen consumption.

[0022] 6) The oil adsorption desulfurization and upgrading method provided by the present invention can significantly improve the viscosity reduction and desulfurization effect and reduce the metal content in the upgraded product by selecting a specific catalyst;

[0023] 7) The oil adsorption desulfurization and upgrading method and system provided by the present invention have low reaction pressure, small hydrogen-to-oil ratio, low hydrogen consumption, easy operation, high yield of upgraded products, and low coke yield, making them suitable for industrial promotion. Attached Figure Description

[0024] Figure 1 This is a schematic diagram of the structure of an oil adsorption desulfurization and upgrading system in a preferred embodiment of the present invention.

[0025] Explanation of reference numerals in the attached figures

[0026] 1. Reducer; 2. Moving bed reactor; 3. Separation tower

[0027] 4. First-stage regenerator; 5. Second-stage regenerator; 6. Oil washer.

[0028] 7. Fractionating tower; 8. Regenerator; 9. Separator

[0029] 10. Locking hopper; 11. Elevator; 12. Flue gas desulfurization system

[0030] 13. Hydrogen separation unit 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] A first aspect of the present invention provides a method for oil adsorption desulfurization and upgrading, wherein the method includes the following steps:

[0033] (1) In the presence of hydrogen, the material containing fresh catalyst is reduced to obtain an active catalyst;

[0034] (2) The oil, reaction hydrogen and the active catalyst are contacted in a moving bed reactor to carry out adsorption desulfurization and upgrading reaction to obtain the upgrading product and the catalyst after the reaction.

[0035] (3) Separate the modified product to obtain the modified product;

[0036] (4) The catalyst after the reaction is regenerated to obtain a regenerator;

[0037] (5) The first-stage regenerator is subjected to a second-stage regeneration to obtain a regenerated catalyst;

[0038] (6) The regenerated catalyst is returned to the material from step (1) for recycling.

[0039] In step (1):

[0040] In a preferred embodiment, based on the total amount of the fresh catalyst, the fresh catalyst comprises 65-85 wt%, preferably 70-80 wt%, of a support; 8-15 wt%, preferably 10-13 wt%, of the main active component; 2-5 wt%, preferably 3-4 wt%, of the co-active component; and 5-15 wt%, preferably 7-13 wt%, of a binder.

[0041] The carrier comprises, based on its total amount, 65-85 wt%, preferably 70-80 wt%, of zinc oxide; 8-22 wt%, preferably 12-18 wt%, of alumina; 0-5 wt%, preferably 3-5 wt%, of titanium oxide; 0-5 wt%, preferably 3-5 wt%, of zirconium oxide; and 4-10 wt%, preferably 5-7 wt%, of modified molecular sieve. The titanium oxide and zirconium oxide contents in the carrier are not simultaneously zero, and it is preferred to contain either titanium oxide or zirconium oxide. The modified molecular sieve comprises a molecular sieve and a modifying element. Based on the total amount of the modified molecular sieve, the modifying element, calculated as an oxide, has a content of 0.1-5 wt%, preferably 1-3 wt%. The molecular sieve is selected from USY molecular sieve and / or Beta molecular sieve, preferably USY molecular sieve or Beta molecular sieve. The modifying element is selected from lanthanum and / or palladium, preferably lanthanum and palladium. The mass ratio of lanthanum and palladium, calculated as oxides, is 1:0.05-0.5, preferably 1:0.1-0.3.

[0042] The main active component is selected from oxides of nickel and / or copper, and the auxiliary active component is selected from one or more oxides of potassium, cobalt, molybdenum, and iron; the binder is selected from alumina and / or silicon oxide.

[0043] In this invention, the content of each component in the fresh catalyst can be calculated based on the amount of raw materials fed during the preparation process of the fresh catalyst. It should be noted that the metal oxide corresponding to iron includes various types; in this invention, it is calculated as ferric oxide.

[0044] In a preferred embodiment, the crushing strength of the fresh catalyst is 40-50 N, preferably 43-48 N; the specific surface area of ​​the fresh catalyst is 100-160 m². 2 / g, preferably 115-145m 2 / g; the pore volume of the fresh catalyst is 0.3-0.4 cm³. 3 / g, preferably 0.33-0.38cm 3 / g; the average pore size of the fresh catalyst is 5-25nm, preferably 10-20nm.

[0045] In a preferred embodiment, the method for preparing the fresh catalyst includes the following steps:

[0046] S1, loading the modifying element onto the molecular sieve to obtain modified molecular sieve powder; wherein, the modifying element is selected from lanthanum and / or palladium, and the molecular sieve is selected from USY molecular sieve and / or Beta molecular sieve.

[0047] S2, the carrier precursor and the pore-forming agent are resonantly mixed to obtain mixture I; wherein, the carrier precursor includes zinc oxide powder, pseudoboehmite, optional titanium oxide powder, optional zirconium oxide powder, and the modified molecular sieve powder; wherein, at least one of the titanium oxide powder and zirconium oxide powder is present.

[0048] S3, the precursor of the main active component is added to the mixture I to obtain mixture II; wherein the precursor of the main active component is a nickel-containing and / or copper-containing soluble compound;

[0049] S4, after mixing the binder precursor with the mixture II, the mixture is granulated and calcined for the first time to obtain solid particles;

[0050] S5, the precursor of the co-active component is impregnated on the solid particles and then calcined a second time to obtain the fresh catalyst; wherein the precursor of the co-active component is selected from one or more soluble compounds of potassium, cobalt, molybdenum and iron.

[0051] In step S1:

[0052] In a preferred embodiment, the loading process involves first immersing the molecular sieve in a solution containing the modifying element, and then drying and calcining it to obtain modified molecular sieve powder.

[0053] In a preferred embodiment, the solution containing the modified element is a solution containing a lanthanum-soluble compound and / or a palladium-soluble compound. The present invention does not specifically limit the lanthanum-soluble compound and the palladium-soluble compound; for example, it can be lanthanum nitrate and palladium nitrate.

[0054] In a preferred embodiment, the silica-to-alumina ratio of the molecular sieve is 10-100, preferably 30-60.

[0055] In a preferred embodiment, the molecular sieve is selected from USY molecular sieve or Beta molecular sieve.

[0056] In a preferred embodiment, the molecular sieve and the solution containing the modified element are fed such that, based on the total amount of the modified molecular sieve powder, the content of the modified element, calculated as oxides, is 0.1-5 wt%, preferably 1-3 wt%.

[0057] In a preferred embodiment, the modifying elements are lanthanum and palladium; wherein the feeding of the solution containing the modifying elements is such that the mass ratio of lanthanum and palladium, based on oxides, is 1:0.05-0.5, preferably 1:0.1-0.3.

[0058] In a preferred embodiment, the immersion conditions include treatment at 30-85°C for 2-10 hours, preferably at 50-70°C for 4-6 hours.

[0059] In step S2:

[0060] In a preferred embodiment, the zinc oxide powder is macroporous nano zinc oxide, and more preferably, the pore size of the zinc oxide powder is 10-50 nm, more preferably 20-40 nm.

[0061] In a preferred embodiment, the pseudoboehmite has a pore volume of 0.2-1 mL / g, preferably 0.4 mL / g-0.8 mL / g; and a specific surface area of ​​150-300 m² / g. 2 / g, preferably 180-260m 2 / g, with a colloidal index greater than or equal to 92%, preferably greater than or equal to 95%.

[0062] In a preferred embodiment, the pore-forming agent is selected from one or more of methylcellulose, carboxymethylcellulose, and guar gum.

[0063] In a preferred embodiment, the feeding of the support precursor results in the fresh catalyst support containing, by total mass, 65-85 wt% zinc oxide, 8-22 wt% alumina, 0-5 wt% titanium oxide or zirconium oxide, and 4-10 wt% modified molecular sieve; wherein the contents of titanium oxide and zirconium oxide are not simultaneously zero.

[0064] In a further preferred embodiment, the feeding of the carrier precursor results in the fresh catalyst carrier containing, by total carrier mass, 70-80 wt% zinc oxide, 12-18 wt% alumina, 3-5 wt% titanium oxide or zirconium oxide, and 5-7 wt% modified molecular sieve.

[0065] In a preferred embodiment, the mass ratio of the pore-forming agent to the zinc oxide powder in the carrier precursor is 1:10-45, preferably 1:20-35.

[0066] In step S3:

[0067] In a preferred embodiment, the nickel-containing soluble compound is selected from one or more of nickel nitrate, nickel acetate, and nickel chloride; the copper-containing soluble compound is selected from one or more of copper nitrate, copper sulfate, and copper chloride.

[0068] In a preferred embodiment, the main active component precursor is added to the mixture I in the form of an aqueous solution. The present invention does not impose a particular limitation on the concentration of the main active component precursor in the aqueous solution; the mass concentration can be 5-30%, for example 5%, 8%, 10%, 15%, 20%, 30%, or any value within any range between any two points.

[0069] In a preferred embodiment, the mass ratio of the main active metal precursor to the carrier precursor in the mixture I is 8-15:65-85, based on the mass of the oxide corresponding to the metal element in the main active metal precursor; preferably 10-13:70-80.

[0070] That is, in this invention, the feeding of the main active component precursor and mixture I results in the fresh catalyst containing, by total mass of catalyst, 8-15 wt%, preferably 10-13 wt%, of the main active component and 65-85 wt%, preferably 70-80 wt%, of the support.

[0071] In a preferred embodiment, to improve the uniformity of mixing, the precursor of the main active component is added to the mixture I and then resonant mixing is continued. The resonant mixing operating conditions in step (S3) are preferably the same as those in step (S2).

[0072] In step S4:

[0073] In a preferred embodiment, the binder precursor is selected from aluminum sol or silica sol, preferably aluminum sol.

[0074] In a preferred embodiment, the mass ratio of the binder precursor to the carrier precursor in the mixture II is 5-15:65-85, based on the mass of the oxide corresponding to the aluminum or silicon element in the binder precursor; preferably 7-13:70-80.

[0075] That is, in this invention, the feeding of the binder precursor results in the prepared fresh catalyst containing 5-15 wt%, preferably 7-13 wt%, of binder based on the total mass of the catalyst.

[0076] In a preferred embodiment, the present invention does not impose any special limitation on the mixing method of the binder precursor and mixture II, such as mechanical stirring.

[0077] In a preferred embodiment, the average particle size of the granules obtained after granulation is 0.5-10 mm, preferably 0.5-5 mm.

[0078] In a preferred embodiment, the operating conditions for the first roasting include: a first roasting temperature of 400-700℃ and a first roasting time of 1-10h; more preferably, the first roasting temperature is 400-550℃ and the first roasting time is 3-10h.

[0079] In a preferred embodiment, the granules obtained after granulation are dried at room temperature for 12-48 hours before the first calcination, and then dried at 60-120°C for 2-24 hours.

[0080] In step S5:

[0081] In a preferred embodiment, the present invention does not impose any particular limitation on the soluble compounds of potassium, cobalt, molybdenum, and iron, and commonly used soluble compounds of potassium, cobalt, molybdenum, and iron in the art can be used in the present invention.

[0082] In a preferred embodiment, the precursor of the co-active component is selected from soluble compounds of cobalt and / or molybdenum.

[0083] In a preferred embodiment, the mass ratio of the auxiliary active component precursor to the carrier precursor in the solid particles is 2-5:65-85, preferably 3-4:70-80, based on the mass of the metal oxide in the auxiliary active component precursor.

[0084] That is, in this invention, the feeding of the precursor of the co-active component results in the fresh catalyst containing 2-5 wt%, preferably 3-4 wt%, of the co-active component based on the total mass of the catalyst.

[0085] In a preferred embodiment, the present invention does not impose any special limitation on the impregnation method; it can be equal volume impregnation or excessive impregnation.

[0086] In a preferred embodiment, the impregnation conditions include: an impregnation temperature of 20-90°C and an impregnation time of 1-24 hours; more preferably, the impregnation temperature is 30-50°C and the impregnation time is 8-24 hours.

[0087] In a preferred embodiment, the impregnated solid particles are first dried at room temperature for 12-48 hours before being calcined a second time, and then dried at 60-120°C for 2-24 hours.

[0088] In a preferred embodiment, the operating conditions for the second roasting include: a second roasting temperature of 400-700℃ and a second roasting time of 1-10h; more preferably, the second roasting temperature is 400-550℃ and the second roasting time is 4-6h.

[0089] In a preferred embodiment, the material containing the fresh catalyst is a fresh catalyst or a mixture of a fresh catalyst and a regenerated catalyst.

[0090] In a preferred embodiment, the reduction operating conditions include: a reduction temperature of 250-550℃, preferably 300-400℃; a reduction pressure of atmospheric pressure -3.5MPa, preferably atmospheric pressure -1.25MPa; a reduction time of 2-8h, preferably 3-6h; and a hydrogen volume hourly space velocity of 500-4500h⁻¹. -1 Preferably 2000-4000h -1 In this invention, unless otherwise specified, all pressures are gauge pressures, and atmospheric pressure has a commonly known meaning in the art.

[0091] In a preferred embodiment, the reduced hydrogen can be high-purity hydrogen with a hydrogen gas fraction of 90-100%, or it can be reformed hydrogen from a catalytic cracking unit.

[0092] In step (2):

[0093] In a preferred embodiment, the oil product includes medium oil and heavy oil. The medium oil and heavy oil are products obtained after processing crude oil; the medium oil has a distillation range of 250-450°C, and the heavy oil has a distillation range >450°C. The oil product may also include directly extracted crude oil, especially heavy crude oil.

[0094] In a preferred embodiment, the sulfur content of the medium oil is 0.3-1 wt%, and the sulfur content of the heavy oil is 0.5-3.5 wt%.

[0095] In a preferred embodiment, the oil product is selected from straight-run diesel oil, catalytic diesel oil, catalytic slurry oil, coking feedstock, coking diesel oil, coking wax oil, atmospheric residue, vacuum residue, wax oil, heavy crude oil, and any mixture of the above oil products. The above-mentioned oil products all have conventional meanings in the art, and will not be described in detail here.

[0096] In a preferred embodiment, the moving bed reactor is preferably a medium-pressure gas-liquid-solid moving bed reactor. Here, "medium pressure" in a medium-pressure gas-liquid-solid moving bed reactor refers to 1.6 MPa-10.0 MPa.

[0097] This invention does not impose any special limitations on the contact method between the oil and the catalyst; they can be in co-current or counter-current contact. This invention also does not impose any special limitations on the feed direction of the oil in the moving bed reactor; it can be top-in, bottom-out or bottom-in, top-out.

[0098] In a preferred embodiment, the operating conditions of the adsorption desulfurization and upgrading reaction include: a reaction temperature of 300-450℃, preferably 350-450℃; a reaction pressure of 1-10 MPa, preferably 3-7 MPa; and a hydrogen volume hourly space velocity of 0.2-0.8 h⁻¹. -1 Preferably 0.4-0.6h -1 The hydrogen-to-oil ratio is 100-500, preferably 200-400. In this invention, the hydrogen-to-oil ratio is the volumetric flow rate ratio of reacting hydrogen to oil.

[0099] In step (3):

[0100] In a preferred embodiment, the separation is a flash evaporation, which is used to separate the modified product to obtain the modified product, C3 components, and hydrogen-rich tail gas. The C3 components are propane. This invention does not impose special limitations on the operating conditions of the flash evaporation; the modified product can be separated by flash evaporation according to conventional procedures in the art.

[0101] In a preferred embodiment, the hydrogen-rich tail gas is subjected to hydrogen separation to obtain dry gas and recycled hydrogen; wherein the recycled hydrogen is recycled and added to the reaction hydrogen in step (2). In this invention, separating the recycled hydrogen from the reforming product and recycling it back to step (2) as reaction hydrogen can improve the utilization rate of hydrogen and reduce hydrogen consumption.

[0102] In a preferred embodiment, step (3) further includes: dividing the modified product into part a and part b, wherein part a is cyclically added to the oil in step (2), and part b is output as the finished product.

[0103] In this invention, it is preferable to recycle a portion of the modified product back to step (2) as oil for re-adsorption desulfurization modification, which can further reduce the viscosity, API strength, sulfur content, and metal content of the product. In this invention, the modified product portion a is 0-80 wt% of the total mass of the modified product.

[0104] In step (4):

[0105] In a preferred embodiment, the method further includes: before regenerating the post-reaction catalyst, preferably washing the post-reaction catalyst with an oil wash solution; wherein the oil wash solution is selected from one or more of gasoline fractions, single aromatic fractions (including but not limited to toluene, xylene, trimethylbenzene and ethylbenzene), and mixed aromatic fractions (including but not limited to two or more combinations of toluene, xylene, trimethylbenzene and ethylbenzene), preferably gasoline fractions.

[0106] In this invention, before performing a first-stage regeneration of the catalyst after the reaction, the catalyst can be cleaned with an oil washing solution, or a first-stage regeneration treatment can be performed directly on the catalyst. Compared with directly performing a first-stage regeneration treatment on the catalyst after the reaction, the method of cleaning the catalyst with an oil washing solution before performing a first-stage regeneration can further remove the oil adhering to the surface of the deactivated catalyst and adsorbed in the pores, reduce carbon dioxide emissions during the regeneration process, and improve the degree of activity recovery of the regenerated catalyst.

[0107] In a preferred embodiment, the cleaning operating conditions include: the mass ratio of the oil washing solution to the catalyst after reaction is 1-2:1; the cleaning pressure is 1-10 MPa, preferably the same as the reaction pressure of the adsorption desulfurization and upgrading reaction; and the cleaning temperature is 350-430℃.

[0108] In a preferred embodiment, to improve the utilization rate of the oil washing solution and maximize the recovery of the modified product, the waste oil washing solution obtained after washing can be distilled. The waste oil washing solution includes the oil washing solution and components washed off from the catalyst after the reaction. Distillation can recover the oil washing solution and allow for its reuse. The remaining portion after distillation can be mixed with the modified product (or part b of the modified product) and output as the finished product.

[0109] In a preferred embodiment, the operating conditions for the first-stage regeneration include: a first-stage regeneration pressure of 0.01-1 MPa, preferably 0.1-0.5 MPa; a first-stage regeneration temperature of 320-480°C, preferably 360-440°C; and a first-stage regeneration gas space velocity of 500-4000 h⁻¹. -1 Preferably 1000-3000h -1 The oxygen content in the first-stage regeneration gas is 0.8-5.5 v%, preferably 1.2-4.2 v%. In this invention, the first-stage regeneration removes hydrocarbons and carbon deposits from the catalyst surface. The first-stage regeneration gas is a mixture of air and nitrogen; the oxygen content is adjusted by mixing nitrogen into the air.

[0110] In step (5):

[0111] In a preferred embodiment, the method further includes: dividing the regenerator into part A and part B; contacting part A with a regenerating liquid for regeneration treatment to obtain a regenerated product; and performing a two-stage regeneration on the regenerated product and part B.

[0112] In this invention, by contacting a portion of the regenerator with the reactivation liquid for regeneration, metal ions attached to the catalyst surface and pores can be removed, the catalyst structure can be restored, the deactivation problem caused by the reduction of catalyst specific surface area and pore blockage can be solved, and the catalyst can be restored to its reaction activity.

[0113] In a preferred embodiment, portion A accounts for 0-50% of the total mass of the first-stage regenerator. In this invention, when the metal content in the oil is <10 μg / g, the first-stage regenerator does not need reactivation treatment and can be directly regenerated in a second stage. When the metal content in the oil is ≥10 μg / g, it is preferable to reactivate 10%-50% of the first-stage regenerator.

[0114] In a preferred embodiment, the revitalizing solution is selected from one or more of acetic acid, formic acid, citric acid, oxalic acid, and nitric acid. Specifically, in this invention, the nitric acid is preferably dilute nitric acid with a mass concentration of 1-5%, and the revitalizing solution is preferably acetic acid, more preferably pure acetic acid with a mass concentration ≥98%.

[0115] In a preferred embodiment, the regeneration operating conditions include: the mass ratio of the regenerator in section A to the regeneration liquid is 1-2:1, the regeneration pressure is 0.1-1 MPa, and the regeneration temperature is 60-160℃.

[0116] In a preferred embodiment, the operating conditions for the two-stage regeneration include: a two-stage regeneration reaction pressure of 0.01-1 MPa, preferably 0.1-0.5 MPa; a two-stage regeneration temperature of 420-570°C, preferably 440-550°C; and a two-stage regeneration gas space velocity of 500-4500 h⁻¹. -1 Preferably 1000-3000h -1 The oxygen content in the second-stage regeneration gas is 3-15 vol%, preferably 4.2-12.6 vol%.

[0117] In this invention, the two-stage regeneration can convert zinc sulfide in the first-stage regenerator and / or regeneration product back into zinc oxide, thereby restoring the catalyst's activity. The second-stage regeneration gas is a mixture of air and nitrogen, with nitrogen mixed into the air to adjust the oxygen content.

[0118] In a preferred embodiment, the second-stage regeneration temperature is higher than the first-stage regeneration temperature; more preferably, the second-stage regeneration temperature is 60-110°C higher than the first-stage regeneration temperature.

[0119] In a preferred embodiment, the oxygen content in the second-stage regeneration gas is higher than that in the first-stage regeneration gas. More preferably, the oxygen content in the second-stage regeneration gas is 2.9-11% higher than that in the first-stage regeneration gas.

[0120] In a preferred embodiment, the method further includes: desulfurizing the regenerated flue gas generated in the first and second stages of regeneration before discharging it. Preferably, a flue gas desulfurization system is used to desulfurize the regenerated flue gas in this invention.

[0121] In step (6):

[0122] In a preferred embodiment, the method further includes: in step (6), the regenerated catalyst is divided into a recycled catalyst and a waste catalyst, the recycled catalyst is returned to the material in step (1) for recycling, and while returning the recycled catalyst to step (1), an equal amount of fresh catalyst is added to the waste catalyst.

[0123] In a preferred embodiment, the waste biocatalyst accounts for 0.01-5 wt% of the total mass of the regenerated catalyst.

[0124] The oil adsorption desulfurization and upgrading method provided in this invention is carried out in a moving bed reactor. Through the coupling of two-stage regeneration and preferred reactivation treatment, the activity of the regenerated catalyst can be significantly improved, the service life of the catalyst can be extended, and continuous production of oil adsorption desulfurization and upgrading can be realized.

[0125] A second aspect of the present invention provides an oil adsorption desulfurization and upgrading system, the system comprising a reducer 1, a moving bed reactor 2, a separation tower 3, a primary regenerator 4, and a secondary regenerator 5;

[0126] The reducer 1 is connected end-to-end to the moving bed reactor 2, the first-stage regenerator 4, and the second-stage regenerator 5, and the separation tower 3 is connected to the moving bed reactor 2, as shown below. Figure 1 As shown.

[0127] In a preferred embodiment, the moving bed reactor 2 is a medium-pressure gas-liquid-solid moving bed reactor.

[0128] In a preferred embodiment, the system further includes an oil washing device, which includes an oil washer 6 and a fractionation tower 7; wherein the oil washer 6 is disposed between the moving bed reactor 2 and the first stage regenerator 4, and the fractionation tower 7 is connected to the oil washer 6 for separating the waste oil washing liquid from the oil washer 6 and returning the separated oil washing liquid to the oil washer 6.

[0129] In a preferred embodiment, the system further includes a regenerator 8, which is connected in parallel to the pipeline of the first-stage regenerator 4 and the second-stage regenerator 5.

[0130] In a preferred embodiment, the system further includes a separator 9 disposed between the two-stage regenerator 5 and the reducer 1, for removing broken regenerated catalyst, i.e., waste catalyst.

[0131] In a preferred embodiment, a locking hopper 10 is provided in front of the feed inlets of the reducer 1, the moving bed reactor 2, the first stage regenerator 4, the second stage regenerator 5, the oil washer 6, and the regenerator 8, respectively.

[0132] In a preferred embodiment, the outlets of the two-stage regenerator 5 and the oil washing machine 6 are respectively provided with elevators 11, which are used to transport solid materials.

[0133] In a preferred embodiment, a heat exchanger is provided between the moving bed reactor 2 and the separation tower 3, and a flow meter is provided between the separator 9 and the reducer 1.

[0134] In a preferred embodiment, the system further includes a flue gas desulfurization system 12, wherein the flue gas desulfurization system 12 is connected to the first-stage regenerator 4 and the second-stage regenerator 5.

[0135] In a preferred embodiment, the system further includes a hydrogen separation device 13 disposed between the moving bed reactor 2 and the separation tower 3. In this invention, the hydrogen in the moving bed reactor 2 includes hydrogen from the reducer 1, externally supplied hydrogen, and recycled hydrogen separated from the hydrogen separation device 13.

[0136] In this invention, the specific type of device in the oil adsorption desulfurization and upgrading system is not specifically limited, and conventional selection can be made according to actual needs during actual use.

[0137] The present invention will be described in detail below through embodiments. The embodiments and comparative examples are as follows: Figure 1 The process is carried out in the oil adsorption desulfurization and upgrading system shown.

[0138] USY molecular sieves were purchased from Auscatalytic Materials (Dalian) Co., Ltd., with a silica-to-alumina ratio of 50 and a specific surface area of ​​>650 m². 2 g, pore size 0.74nm. Beta molecular sieve purchased from Auscatalytic Materials (Dalian) Co., Ltd., with a silica-to-alumina ratio of 50 and a specific surface area >500m². 2The zinc oxide powder, with a purity of 99.5% and a pore size of 30 nm, was purchased from Shandong Jinyuan New Material Technology Co., Ltd. The pore volume of the boehmite was 0.4 mL / g-0.8 mL / g, and its specific surface area was 180 m². 2 / g-260m 2 / g, colloidal index greater than 95%. Aluminum sol solid content is 20wt%, silica sol solid content is 20wt%.

[0139] The crushing strength of fresh catalyst was determined using a universal automatic strength tester. The tester had a range of 0–500 N and an accuracy of 0.1 N. During testing, a single sample was placed between two rigid platforms, one of which remained stationary while the other moved axially downwards at a very low speed. The crushing strength was defined as the maximum load measured before particle crushing. Fifty regular particles from each sample were selected for the strength test, and the average value was taken as the result.

[0140] The specific surface area, pore size, and pore volume of the fresh catalysts obtained in Examples 1-6 were characterized using BET. During testing: 10g of the sample was weighed and evacuated at 300℃ for 4.0h. Then, the N2 adsorption-desorption isotherm and pore size distribution of the catalyst were measured using a Tri Star II 3020 (Mike Instrument Company, USA) automated analyzer.

[0141] In the examples and comparative examples, raw material AD was heavy oil, and raw materials E and F were medium oil. The properties of each raw material are as follows:

[0142] Raw material A: Desalted and dehydrated crude oil from the North China Oilfield, with a sulfur content of 3.1%, API gravity of 24, and viscosity of 56 mm. 2 / s (50℃), metal content 110μg / g;

[0143] Raw material B: Lanzhou Petrochemical catalytic feedstock (a mixture of vacuum residue and wax oil in a volume ratio of 3:7), sulfur content 0.6%, API gravity 18, viscosity 121 mm. 2 / s (100℃), metal content 56μg / g;

[0144] Raw material C: Sichuan Petrochemical vacuum residue, sulfur content 1.7%, API gravity 16, viscosity 614 mm. 2 / s (100℃), metal content 172μg / g;

[0145] Raw material D: Lanzhou Petrochemical catalytic oil slurry, sulfur content 1.1%, API gravity 20, viscosity 42 mm. 2 / s (100℃), metal content 154μg / g;

[0146] Raw material E: Sichuan Petrochemical wax oil, sulfur content 0.56%, API gravity 28, viscosity 8.5 mm. 2 / s (100℃);

[0147] Raw material F: Yumen Petrochemical catalytic diesel oil, sulfur content 0.86%, API gravity 26, viscosity 4mm. 2 / s (40℃).

[0148] Sulfur content determination: The total sulfur content in liquid samples was determined by ultraviolet fluorescence / chemiluminescence method under the conditions of 1000℃, argon gas at 30psi and oxygen at 10psi using a photonLAB total sulfur and total nitrogen analyzer, referring to the standard SH / T 0689-2000.

[0149] API density is a measure of the density of petroleum and petroleum products. It is calculated by measuring the relative density of the petroleum product at 15.6℃ (60°F) and then applying the formula: API = (141.5 / relative density) - 131.5.

[0150] Determination of kinematic viscosity: Refer to the standard GB / T 265-1988 Petroleum Products - Determination of kinematic viscosity and calculation of dynamic viscosity. Each sample is measured twice and the average value is taken.

[0151] Determination of metal content: Metal content refers to the total mass content of seven metals: sodium, calcium, iron, nickel, vanadium, copper, and lead.

[0152] The sodium content in oil samples was determined using a ContrAA 700 atomic absorption spectrometer from Jena Instruments, Germany. The working principle is that light radiating from a light source containing the characteristic spectral lines of the analyte element is absorbed by the ground-state atoms in the vapor of the sample. The degree of attenuation of the analyte's radiation indicates the sodium content in the oil sample. The oil sample was carbonized, ashed, and decomposed with hydrochloric acid. The solution was then poured into a volumetric flask and diluted to volume. The absorbance of sodium was measured using the atomic absorption spectrometer. The mass of sodium contained was then calculated using a formula and the weighed crude oil sample, expressed in μg / g.

[0153] Inductively coupled plasma atomic emission spectrometry (ICP-AES) was used to measure the content of trace metal elements (calcium, iron, nickel, vanadium, copper, and lead) in petroleum products. After complete combustion of crude oil, organic matter was removed in the blast furnace, and inorganic ash was dissolved in an acid solution. In a plasma environment, the molecules (atoms) in the analytes reached high-energy excited states. The content of various metal elements was qualitatively and quantitatively analyzed by observing the characteristic atomic spectra.

[0154] In the examples and comparative examples, the first and second stage regeneration gases were mixtures of air and nitrogen, respectively. The mass fraction of acetic acid was ≥98%.

[0155] Example 1

[0156] 1. Preparation of fresh catalyst:

[0157] S1, 0.18g of lanthanum nitrate hexahydrate and 0.07g of palladium nitrate dihydrate were added to 110g of water to prepare a mixed solution. 100g of USY molecular sieve and the mixed solution were placed together in a stirred reactor and immersed at 30℃ under normal pressure for 10h. After immersion, the mixture was filtered and the filter cake was dried at 60℃ for 24h. The dried filter cake was calcined at 450℃ for 8h to obtain modified USY molecular sieve powder with a lanthanum oxide (La2O3) content of 0.067wt% and a palladium oxide (PdO) content of 0.033wt%.

[0158] S2, 423g of zinc oxide powder, 140g of boehmite (calculated as alumina), 32g of titanium oxide powder, 55g of the above-mentioned lanthanum palladium modified USY molecular sieve powder and 12.7g of methylcellulose are mixed evenly in a resonant mixer to obtain mixture I;

[0159] S3, add 525.4g of nickel nitrate hexahydrate and 45.5g of copper nitrate trihydrate to 2800g of water to prepare a mixture, then add it to a resonant mixer and mix it with the above mixture I to obtain mixture II;

[0160] S4. Mix 750g of aluminum sol with the above mixture II in a kneader until uniform, then put it into a ball extruder to extrude into small balls with a particle diameter of 0.6mm. Dry the small balls naturally at room temperature for 12h, then dry them at 60℃ for 24h, and then calcine them at 400℃ for 10h to obtain solid particles.

[0161] S5, 107.4g of potassium nitrate was added to 2000g of water to prepare an impregnation solution. All the solid particles were placed in the impregnation solution and impregnated at room temperature for 24h. Then, they were first dried naturally at room temperature for 30h, then dried at 120℃ for 4h, and then calcined at 400℃ for 6h to obtain the catalyst.

[0162] The prepared catalyst contains 42.3 wt% zinc oxide, 14 wt% alumina, 3.2 wt% titanium oxide, 5.5 wt% modified molecular sieve, 13.5 wt% nickel oxide, 1.5 wt% copper oxide, 5 wt% potassium oxide, and 15 wt% binder (alumina). Of the modified molecular sieve, the sieve is USY molecular sieve, with 0.067 wt% lanthanum oxide (La₂O₃) and 0.033 wt% palladium oxide (PdO). The catalyst has a crushing strength of 46.3 N and a specific surface area of ​​132.4 μm. 2 / g, pore volume is 0.34cm³ 3 / g, with an average pore size of 14.3nm.

[0163] 2. Oil adsorption desulfurization and upgrading:

[0164] (1) The above-mentioned fresh catalyst is reduced in a reducer in contact with hydrogen to obtain an active catalyst; wherein the reduction pressure is atmospheric pressure, the reduction temperature is 300℃, the reduction time is 6h, and the hydrogen volume hourly space velocity is 4000h. -1 ;

[0165] (2) The above-mentioned active catalyst, oil product A, and reaction hydrogen enter the moving bed reactor from the top of the reactor in a co-current manner to carry out the adsorption desulfurization reforming reaction, and obtain the reforming product and the catalyst after the reaction; wherein, the reaction pressure of the adsorption desulfurization reforming reaction is 4 MPa, the reaction temperature is 410℃, and the mass hourly space velocity of oil product A is 0.25 h⁻¹. -1 The hydrogen-to-oil ratio is 400:1;

[0166] (3) The above-mentioned modified products are separated in a flash tank to obtain modified products, carbon three components and hydrogen-rich tail gas; among them, the extracted hydrogen-rich tail gas is introduced into a hydrogen separation device for hydrogen separation to obtain dry gas and circulating hydrogen. The obtained circulating hydrogen is pressurized and then added to the reaction hydrogen in step (2). The extracted modified products are output as finished products.

[0167] (4) After the above reaction, the catalyst is isolated by a closed hopper and then enters an oil washing machine to clean the catalyst using gasoline fractions; wherein the cleaning pressure is 4 MPa, the cleaning temperature is 400℃, and the mass ratio of oil washing liquid to catalyst after reaction is 1.5:1.

[0168] The waste oil washing liquid after cleaning enters the fractionation tower. The oil washing liquid collected from the top of the fractionation tower is recycled back to the oil washing unit. The components collected from the bottom of the fractionation tower are mixed with the modified product in step (2) and output as the finished product.

[0169] After cleaning, the reaction catalyst is fed into a locked hopper via an elevator. After being isolated by the locked hopper, it enters a first-stage regenerator for regeneration, yielding a first-stage regenerator. The first-stage regeneration pressure is 0.1 MPa, the first-stage regeneration temperature is 400℃, and the first-stage regeneration gas space velocity is 2000 h⁻¹. -1 The oxygen content in the first stage of regenerated gas is 2.1%;

[0170] (5) The above-mentioned regenerator is divided into part A and part B. Part A of the regenerator accounts for 20% of the total mass of the above-mentioned regenerator. Part A of the regenerator enters the regenerator and is contacted with acetic acid for regeneration treatment to obtain the regeneration product. The regeneration reaction pressure is 0.1 MPa, the regeneration reaction temperature is 100℃, and the mass ratio of part A of the regenerator to the regeneration liquid is 1.2:1.

[0171] The aforementioned regeneration product and the first-stage regenerator (part B) are isolated by a closed hopper and then enter a second-stage regenerator for second-stage regeneration to obtain a regenerated catalyst. The second-stage regeneration pressure is 0.1 MPa, the second-stage regeneration temperature is 480℃, and the second-stage regeneration gas space velocity is 2000 h⁻¹. -1 The oxygen content in the second-stage regenerated gas is 12.6% v.

[0172] Among them, the regenerated flue gas emitted from the first stage regenerator and the second stage regenerator is desulfurized by the flue gas desulfurization system and then released to the outside;

[0173] (6) The above-mentioned regenerated catalyst enters the separator to separate the recycled catalyst and the waste catalyst; wherein the waste catalyst accounts for 3% of the total mass of the above-mentioned regenerated catalyst; the separated recycled catalyst enters the lock hopper through the elevator, and after being isolated by the lock hopper, it enters the reducer. At the same time, fresh catalyst of the same mass as the separated waste catalyst is added to the reducer. The recycled catalyst and the added fresh catalyst are reduced by contacting hydrogen in the reducer. The reduced active catalyst is isolated by the lock hopper and returned to the moving bed reactor for recycling.

[0174] Analysis of the reaction reveals that:

[0175] The dry gas yield was 0.73 wt%, the C3 component yield was 0.26 wt%, the modified product yield was 94.81 wt%, and the coke yield was 1.32 wt%; the sulfur recovery rate in the flue gas desulfurization system was 2.88%.

[0176] Properties of the modified product: sulfur content 0.22%, API gravity 34, viscosity 11 mm. 2 / s (50℃), metal content 16μg / g.

[0177] Example 2

[0178] 1. Preparation of fresh catalyst:

[0179] S1, 2.42g of lanthanum nitrate hexahydrate and 0.2g of palladium nitrate dihydrate were added to 162g of water to prepare a mixed solution. 100g of Beta molecular sieve and the mixed solution were placed together in a stirred reactor and immersed at 50℃ under normal pressure for 6h. After immersion, the mixture was filtered and the filter cake was dried at 60℃ for 24h. The dried filter cake was calcined at 550℃ for 3h to obtain modified Beta molecular sieve powder with a lanthanum oxide (La2O3) content of 0.91wt% and a palladium oxide (PdO) content of 0.09wt%.

[0180] S2, 496g of zinc oxide powder, 131g of boehmite (calculated as alumina), 24g of zirconium oxide powder, 54g of the above-mentioned lanthanum palladium modified Beta molecular sieve and 19.9g of carboxymethyl cellulose are mixed evenly in a resonant mixer to obtain mixture I;

[0181] S3, add 333.1g of nickel acetate tetrahydrate and 31.4g of copper sulfate pentahydrate to 2600g of water to prepare a mixture, then add it to a resonant mixer and mix it with the above mixture I to obtain mixture II;

[0182] S4. Mix 700g of silica sol with the above mixture II in a kneader until homogeneous, then put it into a pelletizer and extrude it into small balls with a particle diameter of 2mm. First, dry the small balls naturally at room temperature for 12h, then dry them at 60℃ for 24h; then calcine them at 500℃ for 3h to obtain solid particles.

[0183] S5, 157.9g of cobalt nitrate hexahydrate was added to 2200g of water to prepare an impregnation solution. All the solid particles were placed in the impregnation solution and impregnated at 45°C for 10h. Then, the particles were first dried naturally at room temperature for 30h, then dried at 120°C for 4h, and then calcined at 450°C for 5h to obtain the catalyst.

[0184] The prepared catalyst contains 49.6 wt% zinc oxide, 13.1 wt% alumina, 2.4 wt% zirconium oxide, 5.4 wt% modified molecular sieve, 10 wt% nickel oxide, 1 wt% copper oxide, 4.5 wt% cobalt trioxide, and 14 wt% binder (silicon oxide). Of the modified molecular sieve, the sieve is Beta sieve, with 0.91 wt% lanthanum oxide (La₂O₃) and 0.09 wt% palladium oxide (PdO). The catalyst has a crushing strength of 47.1 N and a specific surface area of ​​126.8 μm. 2 / g, pore volume 0.35cm³ 3 / g, with an average pore size of 13.5nm.

[0185] 2. Oil adsorption desulfurization and upgrading:

[0186] (1) The above-mentioned fresh catalyst was reduced in a reducer by contacting hydrogen to obtain an active catalyst; wherein the reduction pressure was 0.75 MPa, the reduction temperature was 350 °C, the reduction time was 4 h, and the hydrogen volume hourly space velocity was 2000 h⁻¹. -1 ;

[0187] (2) The above-mentioned active catalyst, oil product B, and reaction hydrogen enter the moving bed reactor from the top of the reactor in a co-current manner to carry out the adsorption desulfurization and upgrading reaction, obtaining the upgraded product and the catalyst after the reaction; wherein, the reaction pressure of the adsorption desulfurization and upgrading reaction is 6 MPa, the reaction temperature is 430℃, and the mass hourly space velocity of oil product B is 0.5 h⁻¹. -1 The hydrogen-to-oil ratio is 300:1;

[0188] (3) The above-mentioned modified products are separated in a flash tank to obtain modified products, carbon three components and hydrogen-rich tail gas; among them, the extracted hydrogen-rich tail gas is introduced into a hydrogen separation device for hydrogen separation to obtain dry gas and circulating hydrogen. The obtained circulating hydrogen is pressurized and then added to the reaction hydrogen in step (2); the extracted modified products are divided into part a and part b. Part a is recycled back to the moving bed reactor, and part b is output as finished product; among them, part a accounts for 20% of the total mass of modified products, and part b accounts for 80% of the total mass of modified products.

[0189] (4) After the above reaction, the catalyst is isolated by a closed hopper and then enters an oil washing machine to clean the catalyst using gasoline fractions; wherein the cleaning pressure is 6 MPa, the cleaning temperature is 410℃, and the mass ratio of oil washing liquid to catalyst after reaction is 1.5:1.

[0190] The waste oil washing liquid after cleaning enters the fractionation tower. The oil washing liquid collected from the top of the fractionation tower is recycled back to the oil washing unit. The components collected from the bottom of the fractionation tower are mixed with the modified product in step (2) and output as the finished product.

[0191] After cleaning, the reaction catalyst is fed into a locked hopper via an elevator. After being isolated by the locked hopper, it enters a first-stage regenerator for regeneration, yielding a first-stage regenerator. The first-stage regeneration pressure is 0.5 MPa, the first-stage regeneration temperature is 420℃, and the first-stage regeneration gas space velocity is 1000 h⁻¹. -1 The oxygen content in the first stage of regenerated gas is 1.26v%.

[0192] (5) The above-mentioned regenerator is divided into part A and part B. Part A of the regenerator accounts for 10% of the total mass of the above-mentioned regenerator. Part A of the regenerator enters the regenerator and is contacted with acetic acid for regeneration treatment to obtain the regeneration product. The regeneration reaction pressure is 0.5 MPa, the regeneration reaction temperature is 130℃, and the mass ratio of part A of the regenerator to the regeneration liquid is 1.2:1.

[0193] The aforementioned regeneration product and the first-stage regenerator (part B) are isolated by a closed hopper and then enter a second-stage regenerator for second-stage regeneration to obtain a regenerated catalyst. The second-stage regeneration pressure is 0.5 MPa, the second-stage regeneration temperature is 510℃, and the second-stage regeneration gas space velocity is 1000 h⁻¹. -1 The oxygen content in the second-stage regenerated gas is 4.2% v.

[0194] Among them, the regenerated flue gas emitted from the first stage regenerator and the second stage regenerator is desulfurized by the flue gas desulfurization system and then released to the outside;

[0195] (6) The above-mentioned regenerated catalyst enters the separator to separate the recycled catalyst and the waste catalyst; wherein, the waste catalyst accounts for 4% of the total mass of the above-mentioned regenerated catalyst; the separated recycled catalyst enters the lock hopper through the elevator, and after being isolated by the lock hopper, it enters the reducer. At the same time, fresh catalyst of the same mass as the separated waste catalyst is added to the reducer. The recycled catalyst and the added fresh catalyst are reduced by contacting hydrogen in the reducer. The reduced active catalyst is isolated by the lock hopper and returned to the moving bed reactor for recycling.

[0196] Analysis of the reaction reveals that:

[0197] The dry gas yield was 0.49 wt%, the C3 component yield was 0.17 wt%, the modified product yield was 97.97 wt%, and the coke yield was 0.82 wt%; the sulfur recovery rate in the flue gas desulfurization system was 0.55%.

[0198] Properties of the modified product: sulfur content 0.054%, API gravity 30, viscosity 22 mm. 2 / s (100℃), metal content 8μg / g.

[0199] Example 3

[0200] 1. Preparation of fresh catalyst:

[0201] S1, 4.44g of lanthanum nitrate hexahydrate and 0.72g of palladium nitrate dihydrate were added to 180g of water to prepare a mixed solution. 100g of USY molecular sieve and the mixed solution were placed together in a stirred reactor and immersed at 60℃ under normal pressure for 5h. After immersion, the mixture was filtered and the filter cake was dried at 60℃ for 24h. The dried filter cake was calcined at 525℃ for 4.5h to obtain modified USY molecular sieve powder with a lanthanum oxide (La2O3) content of 1.67wt% and a palladium oxide (PdO) content of 0.33wt%.

[0202] S2, 568.5g of zinc oxide powder, 105g of boehmite (calculated as alumina), 31.5g of titanium oxide powder, 45g of the above-mentioned lanthanum-palladium modified USY molecular sieve and 28.4g of guar gum powder are mixed evenly in a resonant mixer to obtain mixture I;

[0203] S3, add 432g of nickel nitrate hexahydrate and 28.3g of copper sulfate pentahydrate to 3000g of water to prepare a mixture, then add it to a resonant mixer and mix it with the above mixture I to obtain mixture II;

[0204] S4. Mix 500g of aluminum sol with the above mixture II in a kneader until uniform, then put it into a ball extruder to extrude into small balls with a particle diameter of 2mm. Dry the small balls naturally at room temperature for 12h, then dry them at 60℃ for 24h, and then calcine them at 550℃ for 4h to obtain solid particles.

[0205] S5, 40.8g of ammonium molybdate was added to 2500g of water to prepare an impregnation solution. All the solid particles were placed in the impregnation solution and impregnated at 50°C for 9 hours. Then, the particles were first dried naturally at room temperature for 30 hours, then dried at 120°C for 4 hours, and then calcined at 550°C for 4 hours to obtain the catalyst.

[0206] The prepared catalyst contains 56.8 wt% zinc oxide, 10.5 wt% alumina, 3.2 wt% titanium oxide, 4.5 wt% modified molecular sieve, 11.1 wt% nickel oxide, 0.9 wt% copper oxide, 3 wt% molybdenum oxide, and 10 wt% binder (alumina). Of the modified molecular sieve, USY molecular sieve is the most abundant, with 1.67 wt% lanthanum oxide (La₂O₃) and 0.33 wt% palladium oxide (PdO). The catalyst has a crushing strength of 46.0 N and a specific surface area of ​​119.5 μm. 2 / g, pore volume is 0.33cm³ 3 / g, with an average pore size of 12.1nm.

[0207] 2. Oil adsorption desulfurization and upgrading:

[0208] (1) The above-mentioned fresh catalyst is reduced in a reducer by contacting hydrogen to obtain an active catalyst; wherein the reduction pressure is 1.0 MPa, the reduction temperature is 400℃, the reduction time is 3 h, and the hydrogen volume hourly space velocity is 2500 h⁻¹. -1 ;

[0209] (2) The above-mentioned active catalyst, oil C, and reducing hydrogen enter the moving bed reactor from the top of the reactor in a co-current manner to carry out the adsorption desulfurization reforming reaction, obtaining the reforming product and the catalyst after the reaction; wherein, the reaction pressure of the adsorption desulfurization reforming reaction is 8 MPa, the reaction temperature is 450℃, and the mass hourly space velocity of oil C is 0.5 h⁻¹. -1 The hydrogen-to-oil ratio is 500:1;

[0210] (3) The above-mentioned modified products are separated in a flash tank to obtain modified products, carbon three components and hydrogen-rich tail gas; among them, the extracted hydrogen-rich tail gas is introduced into a hydrogen separation device for hydrogen separation to obtain dry gas and circulating hydrogen. The obtained circulating hydrogen is pressurized and then added to the reaction hydrogen in step (2); the extracted modified products are divided into part a and part b. Part a is recycled back to the moving bed reactor and part b is output as finished product; among them, part a accounts for 50% of the total mass of modified products and part b accounts for 50% of the total mass of modified products.

[0211] (4) After the above reaction, the catalyst is isolated by a closed hopper and then enters an oil washing machine to clean the catalyst using gasoline fractions; wherein the cleaning pressure is 8 MPa, the cleaning temperature is 430℃, and the mass ratio of oil washing liquid to catalyst after reaction is 1.2:1.

[0212] The waste oil washing liquid after cleaning enters the fractionation tower. The oil washing liquid collected from the top of the fractionation tower is recycled back to the oil washing unit. The components collected from the bottom of the fractionation tower are mixed with the modified product in step (2) and output as the finished product.

[0213] After cleaning, the reaction catalyst is fed into a locked hopper via an elevator. After being isolated by the locked hopper, it enters a primary regenerator for primary regeneration, yielding a primary regenerator. The primary regeneration pressure is 1.0 MPa, the primary regeneration temperature is 440℃, and the primary regeneration gas space velocity is 3000 h⁻¹. -1 The oxygen content in the first stage of regenerated gas is 4.2%;

[0214] (5) The above-mentioned regenerator is divided into part A and part B. Part A of the regenerator accounts for 30% of the total mass of the above-mentioned regenerator. Part A of the regenerator enters the regenerator and is contacted with acetic acid for regeneration treatment to obtain the regeneration product. The regeneration reaction pressure is 1.0 MPa, the regeneration reaction temperature is 160℃, and the mass ratio of part A of the regenerator to the regeneration liquid is 1.2:1.

[0215] The aforementioned regeneration product and the first-stage regenerator (part B) are fed into a second-stage regenerator for second-stage regeneration to obtain a regenerated catalyst; wherein the second-stage regeneration pressure is 1.0 MPa, the second-stage regeneration temperature is 550℃, and the second-stage regeneration gas space velocity is 3000 h⁻¹. -1 The oxygen content in the second-stage regenerated gas is 12.6% v.

[0216] Among them, the regenerated flue gas emitted from the first stage regenerator and the second stage regenerator is desulfurized by the flue gas desulfurization system and then released to the outside;

[0217] (6) The above-mentioned regenerated catalyst enters the separator to separate the recycled catalyst and the waste catalyst; wherein, the waste catalyst accounts for 5% of the total mass of the above-mentioned regenerated catalyst; the separated recycled catalyst enters the lock hopper through the elevator, and after being isolated by the lock hopper, it enters the reducer. At the same time, fresh catalyst of the same mass as the separated waste catalyst is added to the reducer. The recycled catalyst and the added fresh catalyst are reduced by contacting hydrogen in the reducer. The reduced active catalyst is isolated by the lock hopper and returned to the moving bed reactor for recycling.

[0218] Analysis of the reaction reveals that:

[0219] The dry gas yield was 0.84 wt%, the C3 component yield was 0.61 wt%, the modified product yield was 95.60 wt%, and the coke yield was 1.56 wt%; the sulfur recovery rate in the flue gas desulfurization system was 1.39%.

[0220] Properties of the modified product: sulfur content 0.31%, API gravity 26, viscosity 121 mm. 2 / s (100℃), metal content 34μg / g.

[0221] Example 4

[0222] 1. Preparation of fresh catalyst:

[0223] S1, 6.14g of lanthanum nitrate hexahydrate and 1.5g of palladium nitrate dihydrate were added to 200g of water to prepare a mixed solution. 100g of Beta molecular sieve and the mixed solution were placed together in a stirred reactor and immersed at 70℃ under normal pressure for 4h. After immersion, the mixture was filtered and the filter cake was dried at 60℃ for 24h. The dried filter cake was calcined at 500℃ for 6h to obtain modified Beta molecular sieve powder with a lanthanum oxide (La2O3) content of 2.31wt% and a palladium oxide (PdO) content of 0.69wt%.

[0224] S2, 640g of zinc oxide powder, 88g of pseudoboehmite (calculated as alumina), 32g of zirconium oxide powder, 40g of the above-mentioned lanthanum palladium modified Beta molecular sieve and 38.4g of methylcellulose are mixed evenly in a resonant mixer to obtain mixture I;

[0225] S3, add 328.1g of nickel acetate tetrahydrate and 19.7g of copper nitrate trihydrate to 2500g of water to prepare a mixture, then add it to a resonant mixer and mix it with the above mixture I to obtain mixture II;

[0226] S4. Mix 350g of silica sol with the above mixture II in a kneader until uniform, then put it into a ball extruder to extrude into small balls with a particle diameter of 2.5mm. Dry the small balls naturally at room temperature for 12h, then dry them at 60℃ for 24h, and then calcine them at 600℃ for 3h to obtain solid particles.

[0227] S5, 126.3g of ferric nitrate nonahydrate was added to 2700g of water to prepare an impregnation solution. All the solid particles were placed in the impregnation solution and impregnated at 55°C for 12h. Then, the particles were first dried naturally at room temperature for 30h, then dried at 120°C for 4h, and then calcined at 650°C for 3h to obtain the catalyst.

[0228] The prepared catalyst contains 64 wt% zinc oxide, 8.8 wt% alumina, 3.2 wt% zirconium oxide, 4 wt% modified molecular sieve, 9.8 wt% nickel oxide, 0.7 wt% copper oxide, 2.5 wt% ferric oxide, and 7 wt% binder (silicon oxide). Of the modified molecular sieve, the sieve is Beta sieve, with 2.31 wt% lanthanum oxide (La₂O₃) and 0.69 wt% palladium oxide (PdO). The catalyst has a crushing strength of 47.8 N and a specific surface area of ​​138.6 μm. 2 / g, pore volume is 0.38cm³ 3 / g, with an average pore size of 16.7nm.

[0229] 2. Oil adsorption desulfurization and upgrading:

[0230] (1) The above-mentioned fresh catalyst is reduced in a reducer by contacting hydrogen to obtain an active catalyst; wherein the reduction pressure is 1.25 MPa, the reduction temperature is 400℃, the reduction time is 4 h, and the hydrogen volume hourly space velocity is 3000 h⁻¹. -1 ;

[0231] (2) The above-mentioned active catalyst, oil product D, and reducing hydrogen enter the moving bed reactor from the top of the reactor in a co-current manner to carry out the adsorption desulfurization reforming reaction, obtaining the reforming product and the catalyst after the reaction; wherein, the reaction pressure of the adsorption desulfurization reforming reaction is 4 MPa, the reaction temperature is 420℃, and the mass hourly space velocity of oil product D is 1.0 h⁻¹. -1 The hydrogen-to-oil ratio is 400:1;

[0232] (3) The above-mentioned modified products are separated in a flash tank to obtain modified products, carbon three components and hydrogen-rich tail gas; among them, the extracted hydrogen-rich tail gas is introduced into a hydrogen separation device for hydrogen separation to obtain dry gas and circulating hydrogen. The obtained circulating hydrogen is pressurized and then added to the reaction hydrogen in step (2); the extracted modified products are divided into part a and part b. Part a is recycled back to the moving bed reactor and part b is output as finished product; among them, part a accounts for 30% of the total mass of modified products and part b accounts for 70% of the total mass of modified products.

[0233] (4) After the above reaction, the catalyst is isolated by a closed hopper and then enters an oil washing machine to clean the catalyst using gasoline fractions; wherein the cleaning pressure is 4 MPa, the cleaning temperature is 390℃, and the mass ratio of oil washing liquid to catalyst after reaction is 1.2:1.

[0234] The waste oil washing liquid after cleaning enters the fractionation tower. The oil washing liquid collected from the top of the fractionation tower is recycled back to the oil washing unit. The components collected from the bottom of the fractionation tower are mixed with the modified product in step (2) and output as the finished product.

[0235] After cleaning, the reaction catalyst is fed into a closed hopper via an elevator. After being isolated by the closed hopper, it enters a primary regenerator for primary regeneration, yielding a primary regenerator. The primary regeneration pressure is 1.0 MPa, the primary regeneration temperature is 410℃, and the primary regeneration gas space velocity is 1500 h⁻¹. -1 The oxygen content in the first stage of regenerated gas is 1.56v%.

[0236] (5) The above-mentioned regenerator is divided into part A and part B. Part A of the regenerator accounts for 10% of the total mass of the above-mentioned regenerator. Part A of the regenerator enters the regenerator and is contacted with acetic acid for regeneration treatment to obtain the regeneration product. The regeneration reaction pressure is 1.0 MPa, the regeneration reaction temperature is 100℃, and the mass ratio of part A of the regenerator to the regeneration liquid is 1.2:1.

[0237] The aforementioned regeneration product and the first-stage regenerator (part B) are fed into a second-stage regenerator for second-stage regeneration to obtain a regenerated catalyst; wherein the second-stage regeneration pressure is 1.0 MPa, the second-stage regeneration temperature is 530℃, and the second-stage regeneration gas space velocity is 1500 h⁻¹. -1 The oxygen content in the second-stage regeneration gas is 10.2% v.

[0238] Among them, the regenerated flue gas emitted from the first stage regenerator and the second stage regenerator is desulfurized by the flue gas desulfurization system and then released to the outside;

[0239] (6) The above-mentioned regenerated catalyst enters the separator to separate the recycled catalyst and the waste catalyst; wherein, the waste catalyst accounts for 2% of the total mass of the above-mentioned regenerated catalyst; the separated recycled catalyst enters the lock hopper through the elevator, and after being isolated by the lock hopper, it enters the reducer. At the same time, fresh catalyst of the same mass as the separated waste catalyst is added to the reducer. The recycled catalyst and the added fresh catalyst are reduced by contacting hydrogen in the reducer. The reduced active catalyst is isolated by the lock hopper and returned to the moving bed reactor for recycling.

[0240] Analysis of the reaction reveals that:

[0241] The dry gas yield was 0.69 wt%, the C3 component yield was 0.38 wt%, the modified product yield was 97.17 wt%, and the coke yield was 0.82 wt%; the sulfur recovery rate in the flue gas desulfurization system was 0.94%.

[0242] Properties of the modified product: sulfur content 0.16%, API gravity 28, viscosity 6mm. 2 / s (100℃), metal content 23μg / g.

[0243] Example 5

[0244] 1. Preparation of fresh catalyst

[0245] S1, 7.6g of lanthanum nitrate hexahydrate and 2.48g of palladium nitrate dihydrate were added to 194g of water to prepare a mixed solution. 100g of USY molecular sieve and the mixed solution were placed together in a stirred reactor and immersed at 40℃ under normal pressure for 8h. After immersion, the mixture was filtered and the filter cake was dried at 60℃ for 24h. The dried filter cake was calcined at 480℃ for 7h to obtain modified USY molecular sieve powder with a lanthanum oxide (La2O3) content of 2.86wt% and a palladium oxide (PdO) content of 1.14wt%.

[0246] S2, 707g of zinc oxide powder, 76g of pseudoboehmite (calculated as alumina), 31.8g of titanium oxide powder, 35.2g of the above-mentioned lanthanum palladium modified USY molecular sieve and 49.5g of carboxymethyl cellulose are mixed evenly in a resonant mixer to obtain mixture I;

[0247] S3, add 295g of nickel nitrate hexahydrate and 13.2g of copper sulfate pentahydrate to 2200g of water to prepare a mixture, then add it to a resonant mixer and mix it with the above mixture I to obtain mixture II;

[0248] S4. Mix 250g of aluminum sol with the above mixture II in a kneader until uniform, then put it into a ball extruder to extrude into small balls with a particle diameter of 5mm. Dry the small balls naturally at room temperature for 12h, then dry them at 60℃ for 24h, and then calcine them at 650℃ for 4h to obtain solid particles.

[0249] S5, add 43g of potassium nitrate to 2800g of water to prepare an impregnation solution, place all the above solid particles in the impregnation solution, and impregnate at 60℃ for 8h; then, first dry naturally at room temperature for 30h, then dry at 120℃ for 4h, and then calcine at 600℃ for 4h to obtain the catalyst.

[0250] The prepared catalyst contains 70.7 wt% zinc oxide, 7.6 wt% alumina, 3.2 wt% titanium oxide, 3.5 wt% modified molecular sieve, 7.6 wt% nickel oxide, 0.4 wt% copper oxide, 2 wt% potassium oxide, and 5 wt% binder (alumina). Of the modified molecular sieve, USY molecular sieve is the most abundant, with lanthanum oxide (La₂O₃) content of 2.86 wt% and palladium oxide (PdO) content of 1.14 wt%. The catalyst has a crushing strength of 45.8 N and a specific surface area of ​​122.9 μm. 2 / g, pore volume is 0.33cm³ 3 / g, with an average pore size of 18.8nm.

[0251] 2. Oil adsorption desulfurization and upgrading:

[0252] (1) The above-mentioned fresh catalyst is reduced in a reducer by contacting hydrogen to obtain an active catalyst; wherein the reduction pressure is 1.5 MPa, the reduction temperature is 350 °C, the reduction time is 3 h, and the hydrogen volume hourly space velocity is 2500 h⁻¹. -1 ;

[0253] (2) The above-mentioned active catalyst, oil product E, and reactive hydrogen enter the moving bed reactor from the top of the reactor in a co-current manner to carry out the adsorption desulfurization reforming reaction, obtaining the reforming product and the catalyst after the reaction; wherein, the reaction pressure of the adsorption desulfurization reforming reaction is 2 MPa, the reaction temperature is 400℃, and the mass hourly space velocity of oil product E is 1.5 h⁻¹. -1 The hydrogen-to-oil ratio is 300:1;

[0254] (3) The above-mentioned modified products are separated in a flash tank to obtain modified products, carbon three components and hydrogen-rich tail gas; among them, the extracted hydrogen-rich tail gas is introduced into a hydrogen separation device for hydrogen separation to obtain dry gas and circulating hydrogen. The obtained circulating hydrogen is pressurized and then added to the reaction hydrogen in step (2); the extracted modified products are divided into part a and part b. Part a is recycled back to the moving bed reactor, and part b is output as finished product; among them, part a accounts for 10% of the total mass of modified products, and part b accounts for 90% of the total mass of modified products.

[0255] (4) After the above reaction, the catalyst is isolated by a closed hopper and then enters an oil washing machine to clean the catalyst using gasoline fractions; wherein the cleaning pressure is 2 MPa, the cleaning temperature is 370℃, and the mass ratio of oil washing liquid to catalyst after reaction is 1.2:1.

[0256] The waste oil washing liquid after cleaning enters the fractionation tower. The oil washing liquid collected from the top of the fractionation tower is recycled back to the oil washing unit. The components collected from the bottom of the fractionation tower are mixed with the modified product in step (2) and output as the finished product.

[0257] After cleaning, the reaction catalyst is fed into a locked hopper via an elevator. After being isolated by the locked hopper, it enters a primary regenerator for primary regeneration, yielding a primary regenerator. The primary regeneration pressure is 0.5 MPa, the primary regeneration temperature is 380℃, and the primary regeneration gas space velocity is 2500 h⁻¹. -1 The oxygen content in the first stage of regenerated gas is 2.38v%.

[0258] (5) The above-mentioned first-stage regenerator enters the second-stage regenerator for second-stage regeneration to obtain the regenerated catalyst; wherein, the second-stage regeneration pressure is 0.5 MPa, the second-stage regeneration temperature is 500℃, and the second-stage regeneration gas space velocity is 2500 h⁻¹. -1 The oxygen content in the second-stage regenerated gas is 6.8% v.

[0259] Among them, the regenerated flue gas emitted from the first stage regenerator and the second stage regenerator is desulfurized by the flue gas desulfurization system and then released to the outside;

[0260] (6) The above-mentioned regenerated catalyst enters the separator to separate the recycled catalyst and the waste catalyst; wherein, the waste catalyst accounts for 1% of the total mass of the above-mentioned regenerated catalyst; the separated recycled catalyst enters the lock hopper through the elevator, and after being isolated by the lock hopper, it enters the reducer. At the same time, fresh catalyst of the same mass as the separated waste catalyst is added to the reducer. The recycled catalyst and the added fresh catalyst are reduced by contacting hydrogen in the reducer. The reduced active catalyst is isolated by the lock hopper and returned to the moving bed reactor for recycling.

[0261] Analysis of the reaction reveals that:

[0262] The dry gas yield was 0.53 wt%, the C3 component yield was 0.45 wt%, the modified product yield was 97.81 wt%, and the coke yield was 0.69 wt%; the sulfur recovery rate in the flue gas desulfurization system was 0.52%.

[0263] Properties of the modified product: sulfur content 0.04%, API gravity 34, viscosity 1.3 mm. 2 / s (100℃).

[0264] Example 6

[0265] 1. Preparation of fresh catalyst:

[0266] S1, 12.65g of lanthanum nitrate hexahydrate and 0.52g of palladium nitrate dihydrate were added to 198g of water to prepare a mixed solution. 100g of Beta molecular sieve and the mixed solution were placed together in a stirred reactor for immersion at 85℃ under normal pressure for 2h. After immersion, the mixture was filtered and the filter cake was dried at 60℃ for 24h. The dried filter cake was calcined at 600℃ for 2h to obtain modified Beta molecular sieve powder with a lanthanum oxide (La2O3) content of 4.76wt% and a palladium oxide (PdO) content of 0.24wt%.

[0267] S2, 613.4g of zinc oxide powder, 97.6g of boehmite (calculated as alumina), 36.1g of zirconium oxide powder, 32.9g of the above-mentioned lanthanum-palladium modified Beta molecular sieve and 49.1g of guar gum powder are mixed evenly in a resonant mixer to obtain mixture I;

[0268] S3, add 266.5g of nickel acetate tetrahydrate and 30.4g of copper nitrate trihydrate to 2000g of water to prepare a mixture, then add it to a resonant mixer and mix it with the above mixture I to obtain mixture II;

[0269] S4. Mix 450g of silica sol with the above mixture II in a kneader until uniform, then put it into a ball extruder and extrude it into small balls with a particle diameter of 4mm. First, dry the small balls naturally at room temperature for 12h, then dry them at 60℃ for 24h, and then calcine them at 700℃ for 1h to obtain solid particles.

[0270] S5, 54.4g of ammonium molybdate was added to 3000g of water to prepare an impregnation solution. All the solid particles were placed in the impregnation solution and impregnated at 90°C for 1 hour. Then, the particles were first dried naturally at room temperature for 30 hours, then dried at 120°C for 4 hours, and then calcined at 700°C for 2 hours to obtain the catalyst.

[0271] The prepared catalyst contains 61.3 wt% zinc oxide, 9.8 wt% alumina, 3.6 wt% zirconium oxide, 3.3 wt% modified molecular sieve, 8 wt% nickel oxide, 1 wt% copper oxide, 4 wt% molybdenum oxide, and 9 wt% binder (silicon oxide). Of the modified molecular sieve, the sieve is Beta sieve, with 4.76 wt% lanthanum oxide (La₂O₃) and 0.24 wt% palladium oxide (PdO). The catalyst has a crushing strength of 46.6 N and a specific surface area of ​​130.2 μm. 2 / g, pore volume is 0.36cm³ 3 / g, with an average pore size of 14.6nm.

[0272] 2. Oil adsorption desulfurization and upgrading:

[0273] (1) The above-mentioned fresh catalyst is reduced in a reducer in contact with hydrogen to obtain an active catalyst; wherein the reduction pressure is 3 MPa, the reduction temperature is 500℃, the reduction time is 6 h, and the hydrogen volume hourly space velocity is 500 h⁻¹. -1 ;

[0274] (2) The above-mentioned active catalyst, oil product F, and reaction hydrogen enter the moving bed reactor from the top of the reactor in a co-current manner to carry out the adsorption desulfurization and upgrading reaction, obtaining the upgraded product and the catalyst after the reaction; wherein, the reaction pressure of the adsorption desulfurization and upgrading reaction is 1 MPa, the reaction temperature is 380℃, and the mass hourly space velocity of oil product F is 2.0 h⁻¹. -1 The hydrogen-to-oil ratio is 200:1;

[0275] (3) The above-mentioned modified products are separated in a flash tank to obtain modified products, carbon three components and hydrogen-rich tail gas; among them, the extracted hydrogen-rich tail gas is introduced into a hydrogen separation device for hydrogen separation to obtain dry gas and circulating hydrogen. The obtained circulating hydrogen is pressurized and then added to the reaction hydrogen in step (2); the extracted modified products are divided into part a and part b. Part a is recycled back to the moving bed reactor, and part b is output as finished product; among them, part a accounts for 10% of the total mass of modified products, and part b accounts for 90% of the total mass of modified products.

[0276] (4) After the above reaction, the catalyst is isolated by a closed hopper and then enters an oil washing machine to clean the catalyst using gasoline fractions; wherein the cleaning pressure is 1 MPa, the cleaning temperature is 350℃, and the mass ratio of oil washing liquid to catalyst after reaction is 1.2:1.

[0277] The waste oil washing liquid after cleaning enters the fractionation tower. The oil washing liquid collected from the top of the fractionation tower is recycled back to the oil washing unit. The components collected from the bottom of the fractionation tower are mixed with the modified product in step (2) and output as the finished product.

[0278] After cleaning, the reaction catalyst is fed into a closed hopper via an elevator. After being isolated by the closed hopper, it enters a primary regenerator for primary regeneration, yielding a primary regenerator. The primary regeneration pressure is 0.1 MPa, the primary regeneration temperature is 360℃, and the primary regeneration gas space velocity is 2000 h⁻¹. -1 The oxygen content in the first stage of regenerated gas is 3.64v%.

[0279] (5) The above-mentioned first-stage regenerator enters the second-stage regenerator for second-stage regeneration to obtain the regenerated catalyst; wherein, the second-stage regeneration pressure is 0.1 MPa, the second-stage regeneration temperature is 440℃, and the second-stage regeneration gas space velocity is 2000 h⁻¹. -1 The oxygen content in the second-stage regenerated gas is 8.4% v.

[0280] Among them, the regenerated flue gas emitted from the first stage regenerator and the second stage regenerator is desulfurized by the flue gas desulfurization system and then released to the outside;

[0281] (6) The above-mentioned regenerated catalyst enters the separator to separate the recycled catalyst and the waste catalyst; wherein, the waste catalyst accounts for 1% of the total mass of the above-mentioned regenerated catalyst; the separated recycled catalyst enters the lock hopper through the elevator, and after being isolated by the lock hopper, it enters the reducer. At the same time, fresh catalyst of the same mass as the separated waste catalyst is added to the reducer. The recycled catalyst and the added fresh catalyst are reduced by contacting hydrogen in the reducer. The reduced active catalyst is isolated by the lock hopper and returned to the moving bed reactor for recycling.

[0282] Analysis of the reaction reveals that:

[0283] The dry gas yield was 0.36 wt%, the C3 component yield was 0.18 wt%, the modified product yield was 98.25 wt%, and the coke yield was 0.37 wt%; the sulfur recovery rate in the flue gas desulfurization system was 0.84%.

[0284] Properties of the modified product: sulfur content 0.02%, API gravity 30, viscosity 3.32 mm. 2 / s (40℃).

[0285] Example 7

[0286] Similar to Example 1, the difference is that the reactivation treatment in step (5) of the oil adsorption desulfurization and upgrading is omitted, and the first-stage regenerator is directly regenerated in the second stage.

[0287] Analysis of the reaction reveals that:

[0288] The dry gas yield was 0.67 wt%, the C3 component yield was 0.19 wt%, the modified product yield was 95.76 wt%, and the coke yield was 1.36 wt%; the sulfur recovery rate in the flue gas desulfurization system was 2.02%.

[0289] Properties of the modified product: sulfur content 1.08%, API gravity 30, viscosity 20 mm. 2 / s (50℃), metal content 49μg / g.

[0290] Comparative Example 1

[0291] Similar to Example 1, the difference is that step (5) in the oil adsorption desulfurization and upgrading is omitted, that is, the regeneration and second-stage regeneration are omitted, and the first-stage regenerator in step (4) is used as the regeneration catalyst to enter the separator in step (6).

[0292] Analysis of the reaction reveals that:

[0293] The dry gas yield was 0.64 wt%, the C3 component yield was 0.16 wt%, the modified product yield was 97.34 wt%, and the coke yield was 1.44 wt%; the sulfur recovery rate in the flue gas desulfurization system was 0.42%.

[0294] Properties of the modified product: sulfur content 2.46%, API gravity 26, viscosity 49 mm. 2 / s (50℃), metal content 51μg / g.

[0295] As can be seen from the above embodiments, the oil adsorption desulfurization and upgrading method provided in this invention can achieve continuous production through the combination of catalyst and process, significantly improve the desulfurization and viscosity reduction effect, reduce the metal content in the upgraded product, increase the API degree of the upgraded product, improve the utilization rate of hydrogen, reduce hydrogen consumption, and is suitable for industrial promotion.

[0296] Comparing Example 1 and Comparative Example 1, it can be seen that the desulfurization effect was poor only during the first stage of regeneration. Elemental analysis of the regenerated catalyst revealed that the sulfur oxide content in the regenerated catalyst was as high as 16%, indicating that the catalyst was not effectively regenerated, thus leading to a decrease in the catalyst's desulfurization activity.

[0297] 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 oil adsorption desulfurization and upgrading, characterized in that, The method includes the following steps: (1) In the presence of hydrogen, the material containing fresh catalyst is reduced to obtain an active catalyst; (2) The oil, the reaction hydrogen and the active catalyst are contacted in a moving bed reactor to carry out an adsorption desulfurization and upgrading reaction to obtain the upgrading product and the catalyst after the reaction; (3) Separate the modified product to obtain the modified product; (4) The catalyst after the reaction is regenerated in one stage to obtain a regenerator; (5) The first-stage regenerator is subjected to a second-stage regeneration to obtain a regenerated catalyst; (6) The regenerated catalyst is returned to the material from step (1) for recycling; In step (1), based on the total amount of the fresh catalyst, the fresh catalyst includes 65-85 wt% of support, 8-15 wt% of main active component, 2-5 wt% of co-active component and 5-15 wt% of binder. In step (1), based on the total amount of the carrier, the carrier includes 65-85 wt% zinc oxide, 8-22 wt% alumina, 0-5 wt% titanium oxide, 0-5 wt% zirconium oxide, and 4-10 wt% modified molecular sieve; wherein the contents of titanium oxide and zirconium oxide are not simultaneously zero, the modified molecular sieve includes molecular sieve and modifying elements, the molecular sieve is selected from USY molecular sieve and / or Beta molecular sieve; the modifying elements are selected from lanthanum and / or palladium; In step (1), the main active component is selected from oxides of nickel and / or copper, the auxiliary active component is selected from one or more oxides of potassium, cobalt, molybdenum, and iron, and the binder is selected from alumina and / or silicon oxide. In step (4), the operating conditions for the first-stage regeneration include: a first-stage regeneration pressure of 0.01-1 MPa; a first-stage regeneration temperature of 320-480℃; and a first-stage regeneration gas space velocity of 500-4000 h⁻¹. -1 The oxygen content in the first stage of regenerated gas is 0.8-5.5% vol%. In step (5), the operating conditions for the second-stage regeneration include: a second-stage regeneration reaction pressure of 0.01-1 MPa; a second-stage regeneration temperature of 420-570℃; and a second-stage regeneration gas space velocity of 500-4500 h⁻¹. -1 The oxygen content in the second-stage regeneration gas is 3-15%.

2. The method according to claim 1, wherein, In step (1), the reduction operating conditions include: reduction temperature of 250-550℃; reduction pressure of atmospheric pressure to 3.5MPa; reduction time of 2-8h; and hydrogen volume hourly space velocity of 500-4500h⁻¹. -1 .

3. The method according to claim 2, wherein, In step (1), the reduction operating conditions include: a reduction temperature of 300-400℃; a reduction pressure of atmospheric pressure to 1.25 MPa; a reduction time of 3-6 h; and a hydrogen volume hourly space velocity of 2000-4000 h⁻¹. -1 .

4. The method according to claim 1, wherein, In step (2), the oil includes medium oil and heavy oil.

5. The method according to claim 4, wherein, In step (2), the distillation range of the medium oil is 250-450℃, and the distillation range of the heavy oil is >450℃; And / or, the sulfur content of the medium oil is 0.3-1 wt%, and the sulfur content of the heavy oil is 0.5-3.5 wt%. And / or, the oil product is selected from straight-run diesel, catalytic diesel, catalytic slurry, coking diesel, atmospheric residue, vacuum residue, wax oil, heavy crude oil, and any mixture of the above oil products; And / or, the operating conditions for the adsorption desulfurization and upgrading reaction include: a reaction temperature of 300-450℃; a reaction pressure of 1-10 MPa; and a hydrogen volume hourly space velocity of 0.2-0.8 h⁻¹. -1 The hydrogen-to-oil volume ratio is 100-500.

6. The method according to claim 5, wherein, In step (2), the operating conditions for the adsorption desulfurization and upgrading reaction include: a reaction temperature of 350-450℃; a reaction pressure of 3-7 MPa; and a hydrogen volume hourly space velocity of 0.4-0.6 h⁻¹. -1 The hydrogen-to-oil volume ratio is 200-400.

7. The method according to claim 1, wherein, In step (3), the separation is flash evaporation, which is used to separate the modified product to obtain the modified product, the three carbon components and the hydrogen-rich tail gas.

8. The method according to claim 7, wherein, In step (3), the hydrogen-rich tail gas is separated into dry gas and recycled hydrogen; wherein the recycled hydrogen is recycled into the reaction hydrogen in step (2); And / or, the method further includes: dividing the modified product into part a and part b, wherein part a is cyclically added to the oil in step (2), and part b is output as a finished product; part a is 0-80 wt% of the total mass of the modified product.

9. The method according to claim 1, wherein, The method further includes: in step (4), before the regeneration of the first section, cleaning the catalyst after the reaction.

10. The method according to claim 9, wherein, The method further includes: in step (4), washing the catalyst after the reaction with an oil washing solution; wherein the oil washing solution is selected from one or more of gasoline fractions, single aromatic fractions, and mixed aromatic fractions; And / or, the cleaning operating conditions include: the mass ratio of the oil washing solution to the catalyst after reaction is 1-2:1, the cleaning pressure is 1-10 MPa, and the cleaning temperature is 350-430℃.

11. The method according to claim 10, wherein, The oil wash fluid is a gasoline fraction.

12. The method according to claim 1, wherein, In step (4), the operating conditions for the first-stage regeneration include: a first-stage regeneration pressure of 0.1-0.5 MPa; a first-stage regeneration temperature of 360-440℃; and a first-stage regeneration gas space velocity of 1000-3000 h⁻¹. -1 The oxygen content in the first stage of regenerated gas is 1.2-4.2%.

13. The method according to claim 1, wherein, The method further includes: in step (5), the first stage of regenerator is divided into part A and part B, part A is contacted with regenerating liquid for regeneration treatment to obtain regenerating product; the regenerating product and part B are subjected to the second stage of regeneration.

14. The method according to claim 13, wherein, The A portion accounts for 0-50% of the total mass of the regenerant in the first section; And / or, the revitalizing solution is selected from one or more of acetic acid, formic acid, citric acid, oxalic acid, and nitric acid; And / or, the regeneration operation conditions include: the mass ratio of the regenerator in section A to the regeneration liquid is 1-2:1, the regeneration pressure is 0.1-1 MPa, and the regeneration temperature is 60-160℃; And / or, the second-stage regeneration temperature is higher than the first-stage regeneration temperature, and the oxygen content in the second-stage regeneration gas is higher than the oxygen content in the first-stage regeneration gas.

15. The method according to claim 14, wherein, In step (5), the revitalizing solution is acetic acid.

16. The method according to claim 1, wherein, In step (5), the operating conditions for the second-stage regeneration include: the second-stage regeneration reaction pressure is 0.1-0.5 MPa; the second-stage regeneration temperature is 440-550℃; and the second-stage regeneration gas space velocity is 1000-3000 h⁻¹. -1 The oxygen content in the second-stage regenerated gas is 4.2-12.6%.

17. The method according to claim 1, wherein, In step (6), the regenerated catalyst is divided into a recycled catalyst and a waste catalyst. The recycled catalyst is returned to the material in step (1) for recycling. At the same time as the recycled catalyst is returned to step (1), an equal amount of fresh catalyst is added to the waste catalyst.

18. The method according to claim 17, wherein, In step (6), the waste catalyst accounts for 0.01-5 wt% of the total mass of the regenerated catalyst.

19. A system for use in the oil adsorption desulfurization and upgrading method according to any one of claims 1-18, characterized in that, The system includes a reducer (1), a moving bed reactor (2), a separation tower (3), a first-stage regenerator (4), and a second-stage regenerator (5). The reducer (1) is connected end-to-end to the moving bed reactor (2), the first stage regenerator (4) and the second stage regenerator (5), and the separation tower (3) is connected to the moving bed reactor (2).

20. The system according to claim 19, wherein, The system also includes an oil washing device, which includes an oil washing unit (6) and a fractionation tower (7); wherein the oil washing unit (6) is located between the moving bed reactor (2) and a first-stage regenerator (4), and the fractionation tower (7) is connected to the oil washing unit (6) for separating the waste washing oil from the oil washing unit (6) and returning the separated washing oil to the oil washing unit (6).

21. The system according to claim 19, wherein, The system also includes a regenerator (8), which is connected in parallel to the pipelines of the first-stage regenerator (4) and the second-stage regenerator (5).

22. The system according to claim 21, wherein, The system also includes a separator (9) disposed between the two-stage regenerator (5) and the reducer (1); And / or, a locking hopper (10) is provided in front of the feed inlet of the reducer (1), the moving bed reactor (2), the first stage regenerator (4), the second stage regenerator (5), the oil washer (6) and the regenerator (8). And / or, the outlets of the two-stage regenerator (5) and the oil washing machine (6) are respectively provided with elevators (11), which are used to transport solid materials.

23. The system according to any one of claims 19-22, wherein, The system also includes a flue gas desulfurization system (12), wherein the flue gas desulfurization system (12) is connected to the first stage regenerator (4) and the second stage regenerator (5).

24. The system according to any one of claims 19-22, wherein, The system also includes a hydrogen separation device (13), which is disposed between the moving bed reactor (2) and the separation tower (3).

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