A method for preparing jet fuel components by hydrogenating chlorine-containing oil and fat raw materials

Abstract

The present invention relates to a method for preparing jet fuel components by hydrogenating chlorine-containing oil and fat feedstocks. After the chlorine-containing oil and fat feedstock is hydrogenated, the reaction effluent is subjected to high-pressure gas stripping. The resulting gaseous material is subjected to adsorption dechlorination. The liquid material obtained at the bottom of the high-pressure gas stripping tower is sequentially subjected to hydrocracking and hydroisomerization reactions. The resulting reaction effluent is separated to obtain jet fuel components. This method can maximize the production of high-quality jet fuel components, while also having low energy consumption and high environmental performance.

Description

Technical Field

[0001] The present invention relates to the technical field of hydrocarbon raw material processing, and in particular to a method for preparing jet fuel components by hydrogenating chlorine-containing oil and fat raw materials. Background Art

[0002] As traditional fossil fuel supplies tighten and pressure to reduce CO2 emissions grows, the refining industry faces a critical challenge: how to effectively reduce CO2 emissions while increasing fuel supply. The development of biomass fuel is considered an effective means of addressing this issue. For the aviation industry in particular, the use of biomass jet fuel is considered the most important way to address carbon emissions.

[0003] The production of jet fuel using oil as raw material usually adopts a two-step hydrogenation process. First, the oil raw material is hydrogenated to obtain hydrocarbon compounds; then the hydrocarbon compounds are hydrogenated and converted to prepare jet fuel components that meet jet fuel standards.

[0004] US20090158637A1 discloses a method for producing aviation fuel from renewable raw materials such as vegetable oils and animal fats and oils. The method comprises treating the renewable raw materials by hydrogenation and deoxygenation to provide normal paraffins having from about 8 to about 24 carbon atoms. At least some of the normal paraffins are isomerized to improve cold flow properties. At least a portion of the paraffins are selectively cracked to provide paraffins that meet the specifications of different aviation fuels, such as JP-8.

[0005] CN103059900A discloses a method for preparing jet fuel, comprising: (1) contacting vegetable oil and / or animal fat, hydrogen, and a hydrodeoxygenation catalyst under hydrodeoxygenation conditions to obtain C8-C24 normal alkanes; (2) contacting the C8-C24 normal alkanes and hydrogen with a hydroisomerization catalyst under hydroisomerization conditions; and (3) contacting the product of step (2) and hydrogen with a hydrorefining catalyst under hydrorefining conditions, followed by fractionation to obtain jet fuel. The jet fuel prepared using the method of the present invention is composed almost entirely of saturated alkanes and has stable properties, meeting the requirements for No. 3 jet fuel.

[0006] CN102124080A discloses a method for producing hydrocarbons using biomaterials, and hydrocarbons and fuels thereof. The method comprises subjecting the feed to a hydrogenation saturation reaction, a deoxygenation reaction, and an alkane isomerization reaction to generate reaction products, distilling the liquid hydrocarbon components, collecting the target hydrocarbon products, and recycling the liquid hydrocarbon components or the target hydrocarbon products to the reaction step; wherein the catalyst in the reaction step is a supported catalyst of a ten-membered ring one-dimensional pore molecular sieve supported by a Group VIII metal.

[0007] CN103059930A provides a method for preparing jet fuel. The method involves mixing a heavy fraction with a boiling point of 260°C or higher with C8-C24 normal alkanes to obtain a mixture, subjecting it to hydroisomerization and hydrofining. The reaction product is fractionated to obtain jet fuel and the heavy fraction with a boiling point of 260°C or higher. This method utilizes a cyclic conversion method for the heavy fraction to improve the yield of the jet fuel.

[0008] Waste oil and grease, especially waste cooking oil, contains a certain amount of sodium chloride during use. Oil and grease will generate water and hydrogen chloride during the hydrogenation process. Hydrogen chloride and water will generate highly concentrated hydrochloric acid under liquid phase conditions, especially under water dew point conditions, which will cause serious corrosion to the device, shorten the device operation cycle, and even bring safety hazards to the device operation. Summary of the Invention

[0009] The present invention aims to solve the problems in the prior art of high processing cost for dechlorination and easy corrosion of the device by hydrogen chloride during hydrogenation of chlorine-containing oil and fat raw materials.

[0010] The present invention provides a method for preparing jet fuel components by hydrogenating chlorine-containing oil and fat raw materials, comprising: (1) the chlorine-containing oil and fat raw materials enter a hydroprocessing reaction zone in the presence of hydrogen, contact with a hydroprocessing catalyst, and react under hydroprocessing reaction conditions, wherein the chlorine-containing oil and fat raw materials include one or more of various animal and plant oils and restaurant waste oils, and the chlorine content in the chlorine-containing oil and fat raw materials is greater than or equal to 10 mg / kg;

[0011] (2) The reaction effluent obtained in the hydrotreatment reaction zone of step (1) enters a high-pressure stripping tower, stripping gas is introduced into the bottom of the high-pressure stripping tower, and the gaseous material obtained at the top of the high-pressure stripping tower enters a dechlorination reactor and contacts with a dechlorinating agent for adsorption dechlorination. After cooling, the dechlorinated material obtained in the dechlorination reactor enters a dehydration separator, and is separated to obtain an aqueous phase material and a chlorine-free hydrogen-rich gas.

[0012] (3) The liquid phase material obtained at the bottom of the high-pressure gas stripping tower enters the hydrocracking reaction zone and the hydroisomerization reaction zone in the presence of hydrogen, and is respectively contacted with the hydrocracking catalyst and the hydroisomerization catalyst to react. The resulting reaction effluent is separated to obtain at least the jet fuel component;

[0013] The hydrocracking catalyst comprises a support and an active metal component. The support comprises alumina and silica-alumina. The content of alumina is 20 to 80% by weight, and the content of silica-alumina is 80 to 20% by weight, based on the support. The active metal component is one or more metals selected from Group VIII. The content of the active metal component is 0.1 to 5% by weight, calculated as oxide, based on the total weight of the hydrocracking catalyst. The conversion rate of the fraction above 270° C. in the hydrocracking reaction zone is controlled to be 55% to 80%.

[0014] The hydroisomerization catalyst contains a carrier and an active metal component, wherein the carrier contains a mesoporous molecular sieve and alumina, and the active metal component is selected from one or more of cobalt, nickel, palladium, platinum, molybdenum and tungsten. Based on the total weight of the hydroisomerization catalyst, the content of the mesoporous molecular sieve is 20 to 80% by weight, the content of the alumina is 15 to 75% by weight, and the content of the active metal component is 0.2 to 5% by weight, calculated as oxides. The sum of the contents of the above components is 100%.

[0015] In the present invention, the chlorine-containing oil and fat raw materials refer to animal and vegetable oils and / or catering waste oils containing chlorine, and the chlorine content in the chlorine-containing oil and fat raw materials is greater than or equal to 10 mg / kg. In the present invention, the animal and vegetable oils include vegetable oils and animal fats, as well as raw materials containing glycerides and free fatty acids, and fatty acid methyl esters or fatty acid ethyl esters prepared by ester exchange methods of vegetable oils and / or animal fats. The catering waste oil is oil waste that is no longer suitable for consumption and is generated during the processing and consumption of animal and vegetable oils. It includes fatty acids, acidified oils, etc. generated in the process of producing edible oil from oilseeds; various types of catering waste oils such as frying oil, kitchen waste oil, and sewage oil generated in the use of edible oil by households, hotels, catering industries, and food production enterprises; animal fats produced as by-products in meat production and processing, and edible oils that have exceeded their shelf life.

[0016] In one embodiment of the present invention, the chlorine-containing oil and fat raw material enters an optional water washing pretreatment unit for water washing treatment, the amount of water added is 2 to 20% by mass of the raw material, preferably 4 to 10% by mass, and oil-water separation is performed after standing at 50 to 100° C. for 30 to 120 minutes; the dehydrated chlorine-containing oil and fat raw material is sent to the hydroprocessing reaction zone.

[0017] In the present invention, "optional" means optional, not necessary. The chlorine-containing oil and fat raw materials can be pre-treated by water washing or directly enter the hydroprocessing reaction zone.

[0018] In the hydroprocessing reaction zone of step (1), the chlorine-containing oil and fat raw materials are contacted with a hydroprocessing catalyst in the presence of hydrogen, and mainly undergo reactions such as hydrodeoxygenation, olefin saturation, hydrodechlorination, hydrodesulfurization and hydrodenitrogenation to obtain a reaction effluent containing water, hydrogen chloride, hydrogen sulfide and ammonia.

[0019] In one embodiment of the present invention, the reaction conditions of the hydroprocessing reaction zone are: hydrogen partial pressure 1.0-15.0 MPa, reaction temperature 150-450°C, liquid hourly volume space velocity 0.1-10 h -1 , hydrogen-oil volume ratio 300-2000;

[0020] The preferred reaction conditions in the hydroprocessing reaction zone are: hydrogen partial pressure 3.0-10.0 MPa, reaction temperature 250-400°C, liquid hourly space velocity 0.5-5.0 h -1 , hydrogen-oil volume ratio 500~1500.

[0021] In one embodiment of the present invention, the hydroprocessing catalyst comprises a metal active component and a carrier, the carrier is selected from one or more of aluminum oxide, silicon oxide, zirconium oxide, and titanium oxide, the metal active component is at least one metal element selected from Group VIB and at least one metal element selected from Group VIII, the Group VIB metal element is molybdenum and / or tungsten, and the Group VIII metal element is cobalt and / or nickel; based on the total weight of the hydroprocessing catalyst, the content of the metal active component is 5 to 50 weight % in terms of oxide.

[0022] In a preferred embodiment, based on the total weight of the hydroprocessing catalyst, the content of the Group VIB metal element is 4 to 40% by weight, and the content of the Group VIII metal element is 1 to 10% by weight, calculated as oxides; and the metal active component contains cobalt and molybdenum.

[0023] More preferably, based on the total weight of the hydroprocessing catalyst, the content of the Group VIB metal element is 8 to 35% by weight, and the content of the Group VIII metal element is 2 to 5% by weight, calculated as oxide.

[0024] According to the method provided by the present invention, the hydroprocessing catalyst is preferably used as a hydroprocessing catalyst after being sulfurized. In one embodiment of the present invention, in order to maintain the sulfurized form of the catalyst, sulfur is supplemented in the chlorine-containing oil and fat raw material, or sulfur is added to the circulating hydrogen.

[0025] In one embodiment of the present invention, the chlorine-containing oil and fat feedstock contains a sulfiding agent, wherein the mass fraction of the sulfiding agent is 0.01% to 0.5% by weight. The sulfiding agent can be any sulfur-containing substance that can be vaporized under hydrogenation conditions and sulfidize the hydroprocessing catalyst. Preferably, the sulfiding agent is one or more of hydrogen sulfide, carbon disulfide, dimethyl disulfide, methyl sulfide, n-butyl sulfide, and thiophene.

[0026] In step (2), the reaction effluent obtained from the hydrotreatment reaction zone enters a high-pressure stripping tower, and stripping gas is introduced at the bottom of the high-pressure stripping tower to transfer gases such as water, hydrogen chloride, hydrogen sulfide and ammonia generated by the reaction into the gas phase.

[0027] In one embodiment of the present invention, the operating conditions of the high-pressure stripping tower are: pressure 1.0-15.0 MPa, tower top temperature 200-300° C., tower bottom temperature 250-380° C., and the mass fraction of the stripping gas feed to the total amount of the high-pressure stripping tower feed is 0.1%-15%;

[0028] The stripping gas is selected from one or more of hydrogen, nitrogen and water vapor.

[0029] In one embodiment of the present invention, the chlorine content of the liquid phase material obtained at the bottom of the high-pressure gas stripping tower is less than 0.5 mg / kg, the water content is less than 100 mg / kg, and the hydrogen sulfide content is less than 5 mg / kg.

[0030] In step (2), the gaseous material obtained at the top of the high-pressure gas stripping tower enters the dechlorination reactor and contacts the dechlorinating agent for adsorption dechlorination. The inventors of the present invention have found through in-depth research that in the dechlorinating agent adsorption dechlorination process, when the hydrogen chloride gas is in the gas phase, the adsorption efficiency is high, while when the material is in the gas and liquid phases, the liquid covers the adsorbent, resulting in a decrease in the efficiency of the dechlorinating agent. In addition, if the mixed material is cooled and separated before dechlorination, the hydrogen chloride and condensed water vapor will cause severe dew point corrosion during the cooling process of the material. In order to solve the above problems, the present invention adopts a technical solution combining high-pressure gas stripping and high-temperature fractionation to separate the gas and liquid phases of the material at a higher temperature. Due to the gas stripping process, the hydrogen chloride contained in the liquid phase is fully separated and transferred to the gas phase. The gaseous phase material is then introduced into the dechlorination reactor and contacts the dechlorinating agent, and adsorption dechlorination is carried out under gas phase conditions, so that hydrogen chloride is removed in the gas phase with high efficiency and greatly improves the chlorine capacity of the dechlorinating agent. The present invention can effectively reduce the amount of dechlorinating agent used, reduce processing costs, and also avoid the risk of dew point corrosion.

[0031] In one embodiment of the present invention, the dechlorination agent comprises at least one Group IA metal compound and / or at least one Group IIA metal compound, optionally one or more metal oxides selected from Cu, Fe, Zn, and a support and / or a binder;

[0032] The carrier and / or binder is selected from one or more of silica, alumina, silica-alumina, zirconia, and clay. The clay is selected from one or more of kaolin, illite, montmorillonite, and bentonite; and the kaolin includes halloysite.

[0033] In the present invention, "optional" means selectable, and the optional one or several metal oxides selected from Cu, Fe, and Zn means that the one or several metal oxides selected from Cu, Fe, and Zn are optional components of the dechlorination agent.

[0034] In the present invention, the dechlorinating agent is preferably a high-temperature dechlorinating agent and / or a medium-temperature dechlorinating agent. The present invention does not limit the high-temperature dechlorinating agent and the medium-temperature dechlorinating agent in any way, and conventional high-temperature dechlorinating agents and medium-temperature dechlorinating agents can achieve the present invention. High-temperature dechlorinating agents and / or medium-temperature dechlorinating agents with a large chlorine capacity are further preferred.

[0035] In one embodiment of the present invention, the dechlorination agent comprises 50-80 mass % of alkaline earth metal oxide, 10-40 mass % of aluminum oxide, and 2-80 mass % of bentonite, based on the weight of the dechlorination agent.

[0036] In one embodiment of the present invention, the operating conditions of the dechlorination reactor are: pressure 1.0-15.0 MPa, reaction temperature 150-450°C, gas hourly volume space velocity 100-5000 h -1 ;

[0037] Preferably, the operating conditions of the dechlorination reactor are: reaction temperature 250-450°C, pressure 3.0-10.0 MPa, gas hourly volume space velocity 500-2000 h -1 .

[0038] In one embodiment of the present invention, the dechlorinated material obtained in the dechlorination reactor is cooled to a temperature of 30-80° C. and then enters a dehydration separator to obtain an aqueous phase material and a chlorine-free hydrogen-rich gas through separation. The chlorine-free hydrogen-rich gas obtained in the dehydration separator is divided into two paths, one of which is circulated to the inlet of the hydrogenation reactor as circulating hydrogen, and the other is entered into the bottom of the high-pressure stripping tower as stripping gas.

[0039] In step (3) of the present invention, the liquid phase material obtained at the bottom of the high-pressure gas stripping tower enters the hydrocracking reaction zone and the hydroisomerization reaction zone in the presence of hydrogen, and is contacted with the hydrocracking catalyst and the hydroisomerization catalyst respectively to react. After separation of the obtained reaction effluent, at least a jet fuel component is obtained. The mass yield of the jet fuel component obtained in the present invention can reach more than 55%.

[0040] In one embodiment of the present invention, the hydrocracking catalyst carrier comprises alumina and silica-alumina, with the alumina content being 25-75% by weight and the silica-alumina content being 75-25% by weight, based on the carrier. The alumina is an alumina having pores with a diameter of less than 60 angstroms accounting for at least 30% of the total pore volume. The active metal component of the hydrocracking catalyst comprises platinum. Preferably, the active metal components of the hydrocracking catalyst are platinum and palladium.

[0041] In one embodiment of the present invention, the BET specific surface area of silicon oxide-alumina is 150 to 320 m 2 / g, and a pore volume of 0.2 to 1.3 mL / g; the silica-alumina is a silica-alumina having a characteristic γ-alumina X-ray diffraction pattern, and based on the silica-alumina, the silica-alumina contains 5 to 60% by weight of silica and 40 to 95% by weight of alumina. Preferably, based on the silica-alumina, the silica-alumina contains 10 to 45% by weight of silica and 55 to 90% by weight of alumina.

[0042] The silicon oxide-aluminum oxide can be a commercial product or can be prepared using any existing technology.

[0043] In one embodiment of the present invention, the preparation of the carrier includes mixing, shaping and calcining silica-alumina and / or its precursor with alumina and / or alumina precursor, wherein the amount of each component is such that the final carrier, the alumina precursor is selected from hydrated alumina, preferably pseudo-boehmite therein; the silica-alumina precursor is preferably silica-alumina with a pseudo-boehmite structure. They can be commercially available products or prepared using any existing technology. After calcination, the alumina precursor is an alumina with a pore volume of pores with a pore diameter of less than 60 angstroms accounting for more than 25% of the total pore volume, preferably more than 30%, and more preferably more than 35%. The calcination adopts the methods and conditions commonly used in the art, such as the calcination temperature can be 350-650°C; preferably 400-600°C, and the calcination time is 2-6 hours, preferably 3-5 hours.

[0044] In one embodiment of the present invention, the reaction conditions of the hydrocracking reaction zone are as follows: reaction temperature 250-500°C, hydrogen partial pressure 1.0-10.0 MPa, liquid hourly space velocity 0.1-5.0 h -1 , hydrogen-oil volume ratio 300~1500; preferred reaction conditions of the hydrocracking reaction zone: reaction temperature 300~400 ℃, hydrogen partial pressure 3.0~8.0MPa, liquid hourly volume space velocity 0.5~3.0h -1 , hydrogen-to-oil volume ratio is 500-1200.

[0045] In order to obtain the maximum amount of high-quality jet fuel components, the conversion rate of the fraction above 270°C in the hydrocracking reaction zone is controlled to be 55% to 80%. The conversion rate of the fraction above 270°C refers to:

[0046] Conversion rate of fraction >270°C = 100% * (mass fraction of fraction >270°C in the feed to the hydrocracking reaction zone - mass fraction of fraction >270°C in the reaction product of the hydrocracking reaction zone) / mass fraction of fraction >270°C in the feed to the hydrocracking reaction zone.

[0047] If the hydrocracking reaction unit controls the conversion rate of the fraction above 270°C too low, the high content of heavy fractions in the product will lead to a high resin content in the jet fuel components, necessitating the removal of the heavy fractions. This will prevent the maximum amount of jet fuel components from being obtained, and the product fraction quality indicators will not meet jet fuel quality requirements. If the hydrocracking reaction unit controls the conversion rate of the fraction above 270°C too high, not only will the jet fuel component yield be reduced, but the unit's hydrogen consumption will also be significantly increased, making the unit's product plan uneconomical and unreasonable.

[0048] In step (4) of the present invention, the reaction effluent from the hydrocracking unit enters the hydroisomerization unit and contacts the hydroisomerization catalyst to undergo an isomerization reaction. The reaction effluent is subjected to gas-liquid separation and fractionation to obtain a naphtha fraction and a jet fuel component. The jet fuel component obtained meets the quality requirements of Appendix C "Kerosene component produced by hydro-reforming of esters and fatty acids (HEFA-SPK)" of No. 3 jet fuel (GB 6537-2018). In addition, the reaction product of the hydroisomerization unit of the present invention does not contain a diesel component, and the oil feedstock is converted into the jet fuel component in the largest amount.

[0049] In a preferred embodiment, the active metal component of the hydroisomerization catalyst contains palladium and platinum, and the mesoporous molecular sieve is selected from one or more of ZSM-22, Nu-10, Thete-1, ISI-1, ZSM-23, SAPO-11, SAPO-31, and SAPO-41.

[0050] In one embodiment of the present invention, the reaction conditions of the hydroisomerization unit are: reaction temperature 150-500°C, hydrogen partial pressure 1.0-10.0 MPa, liquid hourly volume space velocity 0.2-5.0 h -1 , the hydrogen-oil volume ratio is 300-1000; the preferred reaction conditions of the hydroisomerization unit are: reaction temperature 200-300 ° C, hydrogen partial pressure 0.3-8.0 MPa, liquid hourly volume space velocity 0.5-4.0 h -1 , the hydrogen-to-oil volume ratio is 400-800.

[0051] Features of the present invention:

[0052] 1. The method provided by the present invention involves hydrotreating chlorine-containing oil and fat feedstock to produce normal alkanes. The normal alkanes are then hydrocracking and then hydroisomerized. The hydroisomerization products undergo gas-liquid separation and fractionation to produce jet fuel components. By controlling the conversion rate of the fraction above 270°C in the hydrocracking unit, the present invention maximizes the production of high-quality jet fuel components while minimizing energy consumption and effectively improving the economic benefits of the device.

[0053] 2. To address the problem of significantly reduced chlorine penetration capacity of dechlorinating agents in mixed gas and liquid phases, the present invention utilizes a high-pressure gas stripping tower to concentrate the hydrogen chloride generated in the gas phase of the high-pressure gas stripping tower through a combination of gas stripping and fractionation. The gas phase material from the high-pressure gas stripping tower is then introduced into a dechlorination reactor to contact the dechlorinating agent. In the dechlorination reactor of the present invention, adsorption dechlorination is performed under gas-phase conditions, significantly increasing the chlorine penetration capacity of the dechlorinating agent, saving a large amount of dechlorinating agent, thereby reducing the processing cost of high-chlorine raw materials and effectively solving the problem of hydrogen chloride corrosion on the equipment. BRIEF DESCRIPTION OF THE DRAWINGS

[0054] Figure 1 The figure is a schematic diagram of one embodiment of the method for preparing jet fuel components by hydrogenating chlorine-containing oil and fat raw materials provided by the present invention. DETAILED DESCRIPTION

[0055] The present invention will be further described below with reference to the accompanying drawings, but the present invention is not limited thereto.

[0056] Figure 1 FIG. 1 is a schematic diagram of one embodiment of the method for preparing jet fuel components by hydrogenating chlorine-containing oil and fat raw materials provided by the present invention, as shown in FIG. Figure 1As shown, a chlorine-containing oil and fat feedstock 1 and recycled hydrogen 3a are mixed and then enter the hydroprocessing reaction zone. The hydroprocessing reaction zone is equipped with a fixed-bed hydroprocessing reactor 4, which is loaded with a hydroprocessing catalyst and has multiple catalyst beds. Recycled hydrogen 3b is added between the catalyst beds as cooling hydrogen to control the temperature rise of the catalyst beds. The reaction effluent 5 obtained from the hydroprocessing reaction zone enters a high-pressure stripping tower 6. Stripping gas 3c is introduced to the bottom of the high-pressure stripping tower. The gaseous material 7 obtained at the top of the high-pressure stripping tower 6 enters a dechlorination reactor 8 and contacts a dechlorinating agent for adsorption dechlorination. The dechlorinated material 9 obtained in the dechlorination reactor is cooled and then enters a dehydration separator 10. After separation, an aqueous material 11 and a chlorine-free hydrogen-rich gas 12 are obtained. After the chlorine-free hydrogen-rich gas 12 passes through a hydrogen purification unit 13, the resulting hydrogen-rich gas is mixed with fresh hydrogen 2 and pressurized by a recycle compressor 14 for use as recycle hydrogen. The liquid material 15 obtained at the bottom of the high-pressure stripping tower 6 is mixed with recycled hydrogen 25 and then enters the hydrocracking reaction zone. This hydrocracking reaction zone is equipped with a fixed-bed hydrocracking reactor 16, which is loaded with hydrocracking catalyst and has multiple catalyst beds. Recycled hydrogen 25 is added between the catalyst beds as cooling hydrogen to control the temperature rise of the catalyst beds. The reaction effluent 17 from the hydrocracking reactor enters the hydroisomerization reaction zone, which is equipped with a fixed-bed hydroisomerization reactor 18, which is loaded with hydroisomerization catalyst. The reaction effluent 19 from the hydroisomerization reactor enters a high-pressure separator 20, where gas-liquid separation is performed to produce a hydrogen-rich gaseous material 21 and liquid hydrocarbons 22. The gaseous material 21 is mixed with fresh hydrogen 23 and then enters a compressor 24 to produce recycled hydrogen 25. The liquid hydrocarbons 22 enter a fractionation tower 26, where fractionation produces liquefied gas 27, naphtha 28, and jet fuel components 29.

[0057] The present invention will be further described below with reference to the examples, but the present invention is not limited thereto.

[0058] The hydrogenation active components were determined by X-ray fluorescence spectrometry.

[0059] The hydrotreating catalyst used was RS-2000, produced by the Changling Catalyst Branch of Sinopec Catalyst Company. The dechlorination agent used was WGL-A, produced by Wenzhou Catalyst Factory.

[0060] The preparation method of the hydrocracking catalyst is as follows:

[0061] 68.5 g of CL-3 powder (a product of Changling Catalyst Plant, calcined at 550°C for 4 hours to obtain alumina D, whose specific surface area and pore distribution are listed in Table 1), CS powder (a product of Changling Refinery Catalyst Plant, containing 39.7% silicon oxide, 59.7% aluminum oxide, and a specific surface area of 332 m 2 / g, pore volume 1.11 mL / g, XRD characterization of a pseudo-boehmite phase) and 24.5 g of ammonium fluoride were mixed and extruded into trilobal bars with a circumscribed diameter of 1.5 mm. The wet bars were dried at 120°C for 4 hours and calcined at 600°C for 3 hours to obtain Support S. The composition of Support S is listed in Table 1. 100 g of Support S was impregnated with 96 mL of ammonium chloroplatinate solution with a Pt content of 5.2 g / L for 5 hours, then dried at 105°C for 4 hours and calcined at 350°C for 2 hours to obtain Catalyst C. The platinum content of Catalyst C is listed in Table 1. The catalyst was reduced before use at 450°C for 4 hours under a hydrogen pressure of 0.1 MPa.

[0062] The preparation method of the hydroisomerization catalyst is as follows:

[0063] A ZSM-22 molecular sieve (supplied by Changling Catalyst Factory, with a silicon-aluminum ratio of 56) was uniformly mixed with pseudo-boehmite P1-1 and sesbania powder. A nitric acid solution was added and thoroughly kneaded. The mixture was then extruded into clover-shaped strips with a diameter of 1.3 mm on an extruder. The strips were dried at 120°C for 4 hours and then calcined at 600°C in air for 2 hours to obtain a support. The support was then saturated with a solution containing Pt(NH3)4Cl2, dried at 110°C for 4 hours, and calcined at 400°C in air for 3 hours to obtain a hydroisomerization catalyst. The catalyst contained 50.2% ZSM-22 by weight, 49.0% alumina by weight, and 0.8% Pt by weight. The catalyst was reduced before use at a temperature of 350°C for 4 hours under a hydrogen pressure of 0.1 MPa.

[0064] The chlorine-containing oil and fat raw material used in the examples and comparative examples is waste cooking oil, and its main properties are shown in Table 2.

[0065] In the Examples and Comparative Examples, the hydrotreating reaction zone was equipped with a fixed-bed hydrotreating reactor, which was loaded with 100 mL of RS-2000 hydrotreating catalyst and sulfurized before use. The hydrocracking reaction zone was equipped with a fixed-bed hydrocracking reactor, which was loaded with 100 mL of hydrocracking catalyst C. The hydroisomerization reaction zone was equipped with a fixed-bed hydroisomerization reactor, which was loaded with 100 mL of hydroisomerization catalyst.

[0066] (1) Hydrogenation treatment: 100 mL of RS-2000 catalyst was placed in the hydrotreatment reactor. After the catalyst was sulfided, the reaction pressure was 6.0 MPa, the reaction temperature was 330 °C, and the volume space velocity was 2.0 h -1 Under the condition of hydrogen-to-oil volume ratio of 1200, chlorine-containing catering waste oil (properties shown in Table 1) and 0.3% dimethyl disulfide were hydrogenated.

[0067] The dechlorination reaction temperature was 300°C and the pressure was 5.5 MPa. The properties of the hydrotreated refined oil are shown in Table 2. The chlorine content in the water was <0.5 mg / kg.

[0068] Example 1

[0069] (1) Waste cooking oil and dimethyl disulfide (accounting for 0.3% by weight of the mixed raw materials) are put into the hydrotreating reactor together with the hydrotreating catalyst. The reaction pressure is 6.0 MPa, the reaction temperature is 330 ° C, and the volume space velocity is 2.0 h -1 The reaction was carried out under the conditions of a hydrogen-to-oil volume ratio of 1200.

[0070] (2) The reaction effluent obtained in the hydrotreatment reaction zone enters a high-pressure gas stripping tower. The chlorine-free hydrogen-rich gas obtained by the dehydration separator at the bottom of the high-pressure gas stripping tower is used as the gas stripping gas. The gaseous material obtained at the top of the high-pressure gas stripping tower enters the dechlorination reactor and contacts with the dechlorinating agent for adsorption dechlorination. After the dechlorinated material obtained in the dechlorination reactor is cooled to 50°C, it enters the dehydration separator and is separated to obtain an aqueous phase material and chlorine-free hydrogen-rich gas. The main operating conditions of the high-pressure gas stripping tower and the dechlorination reactor are shown in Table 3. The main properties of the liquid phase material obtained at the bottom of the high-pressure gas stripping tower are shown in Table 4.

[0071] (3) The liquid phase material obtained at the bottom of the high-pressure gas stripping tower enters the hydrocracking reaction zone in the presence of hydrogen and contacts the hydrocracking catalyst at a reaction pressure of 5.0 MPa, a reaction temperature of 358°C, and a volume space velocity of 1.0 h -1 The hydrocracking reaction was carried out under the conditions of hydrogen-to-oil volume ratio of 800, and the conversion rate of the fraction above 270°C in the hydrocracking reaction zone was controlled to be 59%. The reaction effluent of the hydrocracking reactor was fed into the hydroisomerization reactor together, and contacted with the hydroisomerization catalyst. The reaction pressure was 5.0 MPa, the reaction temperature was 290°C, and the volume space velocity was 1.5 h -1 The hydroisomerization reaction was carried out under the following conditions: The resulting liquid reaction product was fractionated at a cut point of 140°C. The fraction above 140°C was the jet fuel component. The reaction results are shown in Table 5.

[0072] Example 2

[0073] Step (1) and step (2) are the same as in Example 1.

[0074] (3) The liquid phase material obtained at the bottom of the high-pressure gas stripping tower enters the hydrocracking reaction zone in the presence of hydrogen and contacts the hydrocracking catalyst at a reaction pressure of 5.0 MPa, a reaction temperature of 368°C, and a volume space velocity of 1.0 h -1The hydrocracking reaction was carried out under the conditions of hydrogen-to-oil volume ratio of 500, and the conversion rate of fractions above 270°C in the hydrocracking reaction zone was controlled to be 79%. The reaction effluent of the hydrocracking reactor was fed into the hydroisomerization reactor together, and contacted with the hydroisomerization catalyst. The reaction pressure was 5.0 MPa, the reaction temperature was 280°C, and the volume space velocity was 1.5 h -1 The hydroisomerization reaction was carried out under the following conditions. The resulting liquid reaction product was fractionated at a cut point of 140°C. The fraction above 140°C was the jet fuel component. The reaction results are shown in Table 5.

[0075] Comparative Example 1

[0076] Step (1) and step (2) are the same as in Example 1.

[0077] (3) The liquid phase obtained at the bottom of the high-pressure gas stripping column enters the hydrocracking reaction zone in the presence of hydrogen and contacts the hydrocracking catalyst. Hydrocracking is carried out under the conditions of a reaction pressure of 5.0 MPa, a reaction temperature of 358°C, a volume space velocity of 1.0 h⁻¹, and a hydrogen-to-oil volume ratio of 500. The conversion of the fraction above 270°C in the hydrocracking reaction zone is controlled to be 59%. The liquid phase reaction products obtained in the hydrocracking reaction zone are fractionated at a cut point of 140°C. The fraction above 140°C is the jet fuel component. The reaction results are shown in Table 5.

[0078] Comparative Example 2

[0079] Step (1) and step (2) are the same as in Example 1.

[0080] (3) The liquid phase obtained at the bottom of the high-pressure gas stripping tower enters the hydrocracking reaction zone in the presence of hydrogen and contacts the hydrocracking catalyst. The hydrocracking reaction is carried out under the conditions of a reaction pressure of 5.0 MPa, a reaction temperature of 348°C, a volume space velocity of 1.0 h⁻¹, and a hydrogen-to-oil volume ratio of 500. The conversion rate of the fraction above 270°C in the hydrocracking reaction zone is controlled to be 39%. The reaction effluent from the hydrocracking reactor enters the hydroisomerization reactor together, contacts the hydroisomerization catalyst, and undergoes a hydroisomerization reaction under the conditions of a reaction pressure of 5.0 MPa, a reaction temperature of 300°C, and a volume space velocity of 1.5 h⁻¹. The resulting liquid phase reaction products are fractionated according to a cut point of 140°C. The fraction above 140°C is the jet fuel component. The reaction results are shown in Table 5.

[0081] Comparative Example 3

[0082] Step (1) and step (2) are the same as in Example 1.

[0083] (3) The liquid phase obtained at the bottom of the high-pressure gas stripping tower enters the hydrocracking reaction zone in the presence of hydrogen and contacts the hydrocracking catalyst. The hydrocracking reaction is carried out under the conditions of a reaction pressure of 5.0 MPa, a reaction temperature of 378°C, a volume space velocity of 1.0 h⁻¹, and a hydrogen-to-oil volume ratio of 500. The conversion rate of the fraction above 270°C in the hydrocracking reaction zone is controlled to be 89%. The reaction effluent from the hydrocracking reactor enters the hydroisomerization reactor together, contacts the hydroisomerization catalyst, and undergoes a hydroisomerization reaction under the conditions of a reaction pressure of 5.0 MPa, a reaction temperature of 300°C, and a volume space velocity of 1.5 h⁻¹. The resulting liquid phase reaction product is fractionated at a cut point of 140°C, and the fraction above 140°C is the jet fuel component. The reaction results are shown in Table 5.

[0084] Comparative Example 4

[0085] (1) Waste cooking oil and dimethyl disulfide (accounting for 0.3% by weight of the mixed raw materials) are put into the hydrotreating reactor together with the hydrotreating catalyst. The reaction pressure is 6.0 MPa, the reaction temperature is 330 ° C, and the volume space velocity is 2.0 h -1 The reaction was carried out under the conditions of a hydrogen-to-oil volume ratio of 1200.

[0086] (2) The reaction effluent obtained in the hydrotreatment reaction zone enters a high-pressure gas stripping tower. Fresh hydrogen without chlorine is introduced into the bottom of the high-pressure gas stripping tower as the stripping gas. The gaseous material obtained at the top of the high-pressure gas stripping tower is hydrogen-rich gas. The mass fraction of chlorine in the hydrogen-rich gas is 328 mg / kg. The main operating conditions of the high-pressure gas stripping tower are shown in Table 3.

[0087] Table 1

[0088]

[0089]

[0090] Table 2

[0091] project Waste cooking oil <![CDATA[Density (20 °C), kg / m 3 > 910.6 Total acid value, mgKOH / g 34 Oxygen content, % 11.98 Chlorine content, mg / kg 59

[0092] Table 3

[0093]

[0094]

[0095] Table 4

[0096] project Liquid material obtained at the bottom of the high-pressure gas stripping tower <![CDATA[Density (20 °C), g / cm 3 > 0.7785 Freezing point, ℃ 16 Sulfur mass fraction, mg / kg <1.0 Chlorine mass fraction, mg / kg <0.5 C15-C18 alkane mass fraction, % 94.4 Distillation range (D2887), ℃ IBP 252 10％ 275 30％ 292 50％ 306 70％ 320 90％ 329 FBP 346

[0097] Table 5

[0098]

Claims

1. A method for preparing jet fuel components by hydrogenating chlorine-containing oil and fat raw materials, comprising: (1) Chlorine-containing oil and fat raw materials enter the hydrotreatment reaction zone in the presence of hydrogen and contact with the hydrotreatment catalyst, and react under the hydrotreatment reaction conditions. The chlorine-containing oil and fat raw materials include one or more of various animal and plant oils and catering waste oils, and the chlorine content in the chlorine-containing oil and fat raw materials is greater than or equal to 10 mg / kg. (2) The reaction effluent obtained in the hydrotreatment reaction zone of step (1) enters a high-pressure stripping tower, stripping gas is introduced into the bottom of the high-pressure stripping tower, and the gaseous material obtained at the top of the high-pressure stripping tower enters a dechlorination reactor and contacts with a dechlorinating agent for adsorption dechlorination. After cooling, the dechlorinated material obtained in the dechlorination reactor enters a dehydration separator, and is separated to obtain an aqueous phase material and a chlorine-free hydrogen-rich gas. (3) The liquid phase material obtained at the bottom of the high-pressure gas stripping tower enters the hydrocracking reaction zone and the hydroisomerization reaction zone in the presence of hydrogen, and is respectively contacted with the hydrocracking catalyst and the hydroisomerization catalyst to react. The resulting reaction effluent is separated to obtain at least the jet fuel component; The hydrocracking catalyst comprises a support and an active metal component, wherein the support comprises alumina and silica-alumina, wherein the content of the alumina is 20 to 80% by weight, and the content of the silica-alumina is 80 to 20% by weight, based on the support; the active metal component is one or more metals selected from Group VIII, wherein the content of the active metal component is 0.1 to 5% by weight, calculated as oxide, based on the total weight of the hydrocracking catalyst; and the conversion rate of the fraction having a temperature greater than 270° C. in the hydrocracking reaction zone is controlled to be 55% to 80%; The hydroisomerization catalyst contains a carrier and an active metal component, wherein the carrier contains a mesoporous molecular sieve and alumina, and the active metal component is selected from one or more of cobalt, nickel, palladium, platinum, molybdenum and tungsten. Based on the total weight of the hydroisomerization catalyst, the content of the mesoporous molecular sieve is 20 to 80% by weight, the content of the alumina is 15 to 75% by weight, and the content of the active metal component is 0.2 to 5% by weight, calculated as oxides. The sum of the contents of the above components is 100%.

2. The method according to claim 1, characterized in that The reaction conditions of the hydrotreating reaction zone are: hydrogen partial pressure 1.0-15.0 MPa, reaction temperature 150-450°C, liquid hourly volume space velocity 0.1-10 h -1 , hydrogen-oil volume ratio 300~2000.

3. The method according to claim 1, characterized in that The reaction conditions of the hydrotreating reaction zone are: hydrogen partial pressure 3.0-10.0 MPa, reaction temperature 250-400°C, liquid hourly space velocity 0.5-5.0 h -1 , hydrogen-oil volume ratio 500~1500.

4. The method according to claim 1, wherein The hydroprocessing catalyst comprises a metal active component and a carrier, wherein the carrier is selected from one or more of aluminum oxide, silicon oxide, zirconium oxide, and titanium oxide; the metal active component is at least one metal element selected from Group VIB and at least one metal element selected from Group VIII, wherein the Group VIB metal element is molybdenum and / or tungsten, and the Group VIII metal element is cobalt and / or nickel; and the content of the metal active component is 5 to 50% by weight, calculated as oxide, based on the total weight of the hydroprocessing catalyst.

5. The method according to claim 4, characterized in that Based on the total weight of the hydroprocessing catalyst, the content of the VIB group metal element is 4-40% by weight, the content of the VIII group metal element is 1-10% by weight, and the metal active component contains cobalt and molybdenum, calculated as oxide.

6. The method according to claim 1, wherein The operating conditions of the high-pressure stripping tower are: pressure 1.0-15.0 MPa, tower top temperature 200-300°C, tower bottom temperature 250-380°C, and the mass fraction of the stripping gas feed to the total feed of the high-pressure stripping tower is 0.1%-15%; The stripping gas is selected from one or more of hydrogen, nitrogen and water vapor.

7. The method according to claim 1, characterized in that The chlorine content of the liquid phase material obtained at the bottom of the high-pressure gas stripping tower is less than 0.5 mg / kg, the water content is less than 100 mg / kg, and the hydrogen sulfide content is less than 5 mg / kg.

8. The method according to claim 1, characterized in that The dechlorination agent contains at least one Group IA metal compound and / or at least one Group IIA metal compound, optionally one or more metal oxides selected from Cu, Fe, and Zn, and a carrier and / or a binder; The carrier and / or binder is selected from one or more of silica, alumina, silica-alumina, zirconia, and clay.

9. The method according to claim 7, characterized in that The dechlorination agent comprises 50-80 mass % of alkaline earth metal oxide, 10-40 mass % of aluminum oxide and 2-80 mass % of bentonite, based on the weight of the dechlorination agent.

10. The method according to claim 1, characterized in that The operating conditions of the dechlorination reactor are: pressure 1.0-15.0 MPa, reaction temperature 150-450°C, gas hourly volume space velocity 100-5000 h -1 ; The dechlorinated material obtained from the dechlorination reactor is cooled to a temperature of 30-80°C and then enters the dehydration separator.

11. The method according to claim 1, wherein The operating conditions of the dechlorination reactor are: reaction temperature 250-450°C, pressure 3.0-10.0 MPa, gas hourly volume space velocity 500-2000 h -1 12. The method according to claim 1, characterized in that The chlorine-free hydrogen-rich gas obtained from the dehydration separator is divided into two paths, one path is circulated to the inlet of the hydrogenation reactor as circulating hydrogen, and the other path is used as stripping gas to enter the bottom of the high-pressure stripping tower.

13. The method according to claim 1, wherein The carrier of the hydrocracking catalyst contains alumina and silica-alumina. Based on the carrier, the content of alumina is 25 to 75% by weight, and the content of silica-alumina is 75 to 25% by weight. The alumina is an alumina in which the pore volume of pores with a pore diameter of less than 60 angstroms accounts for more than 30% of the total pore volume. The active metal component of the hydrocracking catalyst includes platinum.

14. The method according to claim 1 or 13, characterized in that The BET specific surface area of silica-alumina is 150 to 320 m 2 / g, pore volume 0.2~1.3mL / g; the silica-alumina is a silica-alumina having a characteristic γ-alumina X-ray diffraction spectrum, and based on the silica-alumina, the silica-alumina contains 5~60% by weight of silica and 40~95% by weight of alumina.

15. The method according to claim 1, wherein Reaction conditions of the hydrocracking reaction zone: reaction temperature 250-500°C, hydrogen partial pressure 1.0-10.0 MPa, liquid hourly space velocity 0.1-5.0 h -1 , hydrogen-oil volume ratio 300~1500, The conversion rate of the fraction above 270°C in the hydrocracking reaction zone is: Conversion rate of fraction >270°C = 100% * (mass fraction of fraction >270°C in the feed to the hydrocracking reaction zone - mass fraction of fraction >270°C in the reaction product of the hydrocracking reaction zone) / mass fraction of fraction >270°C in the feed to the hydrocracking reaction zone.

16. The method according to claim 1, characterized in that Reaction conditions of the hydrocracking reaction zone: reaction temperature 300-400°C, hydrogen partial pressure 3.0-8.0 MPa, liquid hourly space velocity 0.5-3.0 h -1 , hydrogen-to-oil volume ratio is 500-1200.

17. The method according to claim 1, wherein The active metal components of the hydroisomerization catalyst contain palladium and platinum, and the mesoporous molecular sieve is selected from one or more of ZSM-22, Nu-10, Thete-1, ISI-1, ZSM-23, SAPO-11, SAPO-31, and SAPO-41.

18. The method according to claim 1, wherein The reaction conditions of the hydroisomerization reaction zone are: reaction temperature 150-500°C, hydrogen partial pressure 1.0-10.0 MPa, liquid hourly space velocity 0.2-5.0 h -1 , the hydrogen-to-oil volume ratio is 300-1000.

19. The method according to claim 1, wherein The reaction conditions of the hydroisomerization reaction zone are: reaction temperature 200-300°C, hydrogen partial pressure 0.3-8.0 MPa, liquid hourly volume space velocity 0.5-4.0 h -1 , the hydrogen-to-oil volume ratio is 400-800.

20. The method according to claim 1, wherein The chlorine-containing oil and fat raw materials enter the optional water washing pretreatment unit for water washing treatment. The amount of water added is 2 to 20% by weight of the raw material. After standing at 50 to 100°C for 30 to 120 minutes, oil-water separation is performed; the dehydrated chlorine-containing oil and fat raw materials are sent to the hydrogenation reaction zone.

21. The method according to claim 20, characterized in that The amount of water added is 4 to 10% by mass of the raw material.

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