Method for preparing saturated grease and byproduct alkane through two-stage low-pressure hydrogenation

Abstract

The invention discloses a method for preparing saturated grease and byproduct alkane through two-stage low-pressure hydrogenation, and belongs to the technical field of hydrogenation. According to the method, the problem that the poor-quality raw material is easy to coke in the heating process is solved by directly mixing the hot-state circulating returned material and the cold-state fed material for temperature equalization, and lower operation pressure and temperature are adopted, so that the equipment investment is saved, and the operation cost is reduced.

Description

Technical Field

[0001] The invention relates to the technical field of hydrogenation, and in particular to a method for preparing saturated oil and fat and by-product alkane by two-stage low-pressure hydrogenation and its application. Background Art

[0002] The world economy continues to develop. At present, the main source of chemical products worldwide is fossil energy, the most important of which are oil and coal. Both of these fossil energy sources are non-renewable energy sources. Not only are the resources increasingly depleted, but they are also becoming heavier and inferior, and the processing difficulty and cost are gradually increasing. In addition to improving the existing refining technology, new oil substitutes are being sought to produce products that meet the requirements at the lowest cost, especially the development and utilization of renewable resources, which are receiving more and more attention.

[0003] Unstable and inferior oils contain not only natural components such as chloride ions, heavy metal ions and unsaturated components, but also oxidative and deteriorative components such as alcohols, ketones, aldehydes, acids, and impurities such as dust, salt, and moisture. They are prone to oxidation during storage and transportation, producing harmful substances such as peroxides and aldehydes. These substances will not only reduce the sensory quality of oils, but also cause certain toxicological hazards to the human body; they will also affect the catalytic efficiency of liquid catalysts, making the catalytic hydrogenation treatment of unstable and inferior oils difficult, and the product yield and quality will decrease. Moreover, under high temperature conditions, unstable and inferior oils are very likely to produce polycyclic aromatic hydrocarbons such as shomene ketone, which have been proven to have potential carcinogenic and teratogenic effects. Some unstable and inferior oil production processes will cause the unsaturated fatty acids in the oils to isomerize and generate more trans fatty acids. In summary, unstable and inferior oils will not only bring many potential risks to human health, but also cause serious environmental pollution, and should be taken seriously.

[0004] Chinese patent CN103102906A discloses a two-stage hydrogenation method for producing high-quality solvent oil from bio-oil, and the specific method is: the bio-oil raw material and hydrogen pass through the first hydrogenation reaction zone containing at least two hydrogenation reactors, first through the reactor operated at low temperature, and then through the reactor operated at high temperature, the hydrogen-rich gas is recycled in the first stage, and the liquid enters the second stage hydrocracking reaction zone. The secondary hydrogen gas of hydrogenation reforming is recycled in the second stage, and the liquid product is fractionated to obtain high-quality solvent oil. Under the reaction state, the active component of the hydrogenation catalyst used in the first stage is Pt and / or Pd in ​​a reduced state, and the hydrogenation active component of the hydrogenation catalyst used in the second stage is one or more of W, Mo, Ni and Co in a sulfided state. However, the catalyst used in the first stage hydrogenation reaction zone of this technology is precious metal Pt and / or Pd, whether it is operated at low temperature or high temperature, and the reaction pressure is up to 20MPa; the reaction pressure of the second stage hydrogenation reaction is up to 18MPa, and the reaction operation process needs to continuously replenish sulfur-containing medium. The catalyst cost of this technology is relatively high, the operating conditions are harsh and complex, and the products are traditional naphtha and diesel.

[0005] Chinese patent CN 116254155 A provides a hydrogenation process for producing biomass wax from biomass oil, comprising the following steps: mixing a liquid catalyst containing molybdenum at 100-5000ppm with biomass oil, and then pressurizing and feeding it into a first hydrogenation reactor; the product enters a second hydrogenation reactor containing a solid catalyst; the product enters a high-pressure hot separator, the gas phase goes to a high-pressure cold separator to separate hydrogen and liquid phase, and the liquid phase is decompressed and exits the hydrogenation device for post-processing; the cold separator liquid phase and the hot separation liquid phase separated by post-processing are mixed and then enter a supplementary hydrogenation and upgrading reactor containing a solid catalyst; the product of step S4 is separated to obtain liquid biomass wax. However, the two-stage hydrogenation process of this technology has a maximum hydrogen pressure of 15MPa, and the reuse of the bottom oil of the separation tower is not considered. This technology has harsh operating conditions, low raw material utilization, and the product is an alkane with a carbon number of C14-C20.

[0006] Chinese patent CN 116790288 A discloses a method for producing bio-jet fuel by hydrogenating waste oil, wherein the waste oil, hydrogen and hydrogenation pretreatment catalyst are mixed and then sent to a slurry bed hydrogenation reactor for hydrogenation, and the tailings obtained by the effluent through a hydrocyclone are partially circulated to the slurry bed hydrogenation reactor; the liquid phase product is subjected to solid-liquid separation, and the separated liquid phase product is sent to a series of hydrogenation refining reactors, hydrocracking reactors and post-refining reactors for hydrogenation, and finally the bio-jet fuel is obtained by fractionation. However, the hydrogen pressure of this technology can reach up to 15Mpa, and the catalyst is a conventional fixed catalyst. This technology has harsh operating conditions, low raw material utilization, and the product is saturated alkanes.

[0007] US Patent No. 17631803 discloses a highly dispersed hydrogenation catalyst and its preparation method and its application in preparing biofuel from palm oil or other oils. A highly dispersed hydrogenation catalyst is provided through a multi-stage impregnation and roasting process, which greatly reduces the particle size of the active component and improves the dispersibility of the active component and the number of active sites. However, the catalyst invented in this technology is a conventional fixed catalyst, and the hydrogenation process is not disclosed.

[0008] However, the above process still has some problems: (1) Low-quality oil has a high degree of deterioration and is difficult to process, and the industry is not willing to process it; (2) Since low-quality oil generally has poor stability and a high gum content, it has poor thermal stability and is prone to thermal cracking and thermal polycondensation during heating. It is very easy to form coking deposits in the heating equipment, causing blockage of the heating furnace tubes or related heat exchanger components, thereby causing problems such as deterioration of process operations and shortening of equipment service life; (3) The metal elements in the waste oil have a catalytic effect on coking deposits, making it difficult to heat using conventional indirect heat exchange methods. Summary of the invention

[0009] The purpose of the present invention is to provide a two-stage low-pressure hydrogenation method for preparing saturated oils and by-product alkanes. The method solves the problem of easy coking during the heating process of inferior raw materials by directly mixing hot circulating return materials with cold feed materials to equalize the temperature, and adopts lower operating pressure and temperature to save equipment investment and reduce operating costs. Specifically, the present invention mainly includes the following contents:

[0010] On the one hand, the present invention provides a method for preparing saturated oil and fat by-product alkanes by two-stage low-pressure hydrogenation, the method comprising the following steps:

[0011] (1) The unstable low-quality oil raw material is heated to 240-300° C. and enters a reactor A (6) without solid catalyst loading, and 200-2000 ppm of liquid catalyst a is mixed into the reactor A (6);

[0012] (2) The effluent from reactor A (6) is heated to 180-240°C and enters reactor B (9) without solid catalyst loading;

[0013] (3) The effluent from reactor B (9) enters reactor C (10) filled with solid catalyst b;

[0014] (4) The effluent from the reactor C (10) enters the hot high-pressure separator (11), and the gas phase flowing out of the hot high-pressure separator (11) is cooled and sequentially passes through the cold high-pressure separator (21) and the cold low-pressure separator (22) and enters the gas stabilization tower (23). Acid gas flows out from the top of the gas stabilization tower (23), and the effluent from the bottom enters the atmospheric pressure fractionation tower (15) to obtain saturated fatty acids, saturated fatty acid glycerides and by-product alkanes;

[0015] The liquid phase flowing out of the hot high-pressure separator (11) is divided into two parts, one part is used as circulating oil to go to the circulating oil heating furnace, mixed with the liquid catalyst a and then enters the reactor A (6) again, and the rest passes through the hot low-pressure separator (13) and then enters the vacuum distillation tower (14);

[0016] (5) The bottom oil of the vacuum distillation tower (14) is centrifuged and filtered to obtain a light component which is recycled to the esterification oil intermediate tank (7), and then directly mixed with the effluent of the reactor A without heat exchange, heated to 180-240° C., and then enters the reactor B (9) without solid catalyst loading. The side stream of the vacuum distillation tower (14) is saturated fatty acids and saturated fatty acid glycerides;

[0017] (6) The bottom oil of the atmospheric distillation tower (15) is pressurized by a booster pump (17) and heated by an isomerization feed heating furnace (18), then mixed with fresh hydrogen and enters a reactor D (19) filled with a solid catalyst c at 160-240°C. The effluent of the reactor D (19) is mixed with the gas phase of the hot high-pressure separator (11), cooled, and then enters a cold high-pressure separator (21).

[0018] In the above method, the reaction in reactor A (6) and reactor B (9) is a one-stage reaction, the reaction in reactor C (10) is a two-stage reaction, and alkanes are produced in reactor D (19). The present application uses liquid catalyst a to perform catalytic hydrogenation pretreatment on the raw material. The liquid catalyst a can be evenly dispersed in the raw material and efficiently remove chlorine and metal elements in the raw material.

[0019] Optionally, the unstable low-quality oil is a biological oil containing unsaturated bonds, and the unstable low-quality oil includes any one or more of kitchen waste oil, sewage oil, acidified oil or palm oil.

[0020] Optionally, the hydrogen-to-oil volume ratio in the reactor A (6) is 400-1000, and the pressure is 1-6 MPa, preferably, the pressure is 2-5 MPa.

[0021] In order to improve the catalytic efficiency of the liquid catalyst and improve the yield and quality of the final unsaturated fatty acid glyceride, the chloride ions and heavy metal ions in the raw materials need to be treated. However, if the chlorine and metal elements are removed from the unstable and inferior oil raw materials in advance, it will increase the manpower and material costs. Therefore, a method is needed to remove the chlorine and metal elements in the raw materials without adding additional operations.

[0022] Optionally, the liquid catalyst a comprises a catalyst body and an additive, the catalyst body is a nickel-based room temperature fluidity catalyst, and the nickel-based room temperature fluidity catalyst comprises any one of a nickel-molybdenum fluidity catalyst, a nickel-cobalt fluidity catalyst, and a nickel-molybdenum-cobalt fluidity catalyst;

[0023] The weight of the auxiliary agent is 80-200% of the weight of the catalyst body, and the auxiliary agent is selected from a chelating agent, an organic acid, a carboxylate, and 1,2,4-triazole in a weight ratio of 1:(0.1-0.2):(0.1-0.2):(0.1-0.2).

[0024] The liquid catalyst a of the present application is composed of a catalyst body and an auxiliary agent, and the catalyst body can catalyze the raw material, improve the catalytic hydrogenation activity of the raw material and inhibit coke formation. However, due to the high metal content in the unstable and inferior oil raw material of the present application, the catalysis will cause the loss of active sites of the liquid catalyst and reduce the catalytic activity. Therefore, the present application adopts an auxiliary agent containing a chelating agent, which can chelate with the metal in the raw material to achieve the removal of the metal element, and ensure the normal functioning of the liquid catalyst a and its subsequent liquid catalyst b and solid catalyst c.

[0025] The auxiliary agent is formulated with organic acid, carboxylate, 1,2,4-triazole and chelating agent, which can improve the chelating effect of the chelating agent on the metal, and further improve the removal efficiency and final removal amount of the chelating agent on the metal.

[0026] Optionally, the organic acid is at least one of salicylic acid and glycolic acid;

[0027] The carboxylate is at least one of zinc acetate, sodium tartrate and sodium alginate.

[0028] At present, researchers have also developed many demetallizing chelating agents that can remove crude oil, vegetable oil or recycled oil. For example, patent CN112322340A has developed a deironifying agent, which is compounded with organic acid, carboxylate, Schiff base and chelating agent, and can remove iron from crude oil. However, when it is used in the present application, it is found that the chelating agent in the deironifying agent is not effective under high temperature and high salt conditions. Therefore, it can only be used in crude oil. The removal rate of iron metal in the catalytic hydrogenation of steps (1) and (2) of the present application is very low, and the removal rate of other metals is even lower. Therefore, the present application prepares a new type of chelating agent, which can be combined with 1,2,4-triazole to specifically remove iron, chromium, copper, lead and vanadium metals from the raw materials of the present application, thereby ensuring the effective catalysis of the liquid catalyst a.

[0029] Optionally, the preparation method of the chelating agent is:

[0030] S10: dissolving 2,4,6-trivinyl-2,4,6-trimethylcyclotrisiloxane and methyl allyl alcohol polyoxyethylene ether in a solvent, adding an initiator and reacting at 70-80° C. for 4-6 hours to obtain a solution containing intermediate A;

[0031] S20: adding sulfamic acid and Lewis acid to the solution containing intermediate A, reacting at 120-125° C. for 4-5 hours, then cooling to 100° C., adding ethanolamine for neutralization to obtain a solution containing intermediate B;

[0032] S30: adding a polystyrene solution and a Lewis acid to the solution containing the intermediate B, and reacting at 80-100° C. to obtain a crude product;

[0033] S40: washing, purifying and drying the crude product to obtain the chelating agent.

[0034] In the preparation step S10 of the chelating agent, 2,4,6-trivinyl-2,4,6-trimethylcyclotrisiloxane and methyl allyl alcohol polyoxyethylene ether are polymerized as the first main body of the chelating agent, and in the step S20, the aminosulfonic acid is used to sulfonate the hydroxyl group of the methyl allyl alcohol polyoxyethylene ether to introduce sulfonic acid and amino groups, which can improve the complexing effect on metals when combined with 1,2,4-triazole. In the step S30, polystyrene is used as the second main body of the chelating agent, and its benzene ring can undergo Friedel-Crafts alkylation reaction with the vinyl in the first main body to obtain a chelating agent. The introduction of the polystyrene can improve the temperature and salt resistance of the chelating agent, so that the metal can be removed under high temperature and high salt conditions. At the same time, the cyclotrisiloxane structure contained in the chelating agent structure can improve the dispersibility of the chelating agent in the raw material. The structure of the first body and the second body can make micropores exist in the chelating agent, which are used to store the metal adsorbed by the chelating agent itself and the metal adsorbed by the carboxylate and 1,2,4-triazole coordination, so as to prevent the metal from escaping into the system again at high temperature. The structure of the above-mentioned chelating agent can also achieve the removal of only metals such as iron, chromium, copper, lead and vanadium in the raw material, without removing nickel, molybdenum and cobalt in the liquid catalyst, so as to ensure continuous catalytic hydrogenation of the raw material.

[0035] Optionally, the solvent in step S10 is at least one of tetrahydrofuran, xylene or dichloromethane.

[0036] Optionally, in step S10, the molar ratio of 2,4,6-trivinyl-2,4,6-trimethylcyclotrisiloxane to methyl allyl alcohol polyoxyethylene ether is (1-2): 1;

[0037] In step S20, the molar ratio of methyl allyl alcohol polyoxyethylene ether to aminosulfonic acid is 1:(0.8-1.2);

[0038] In step S30, the molar ratio of polystyrene to 2,4,6-trivinyl-2,4,6-trimethylcyclotrisiloxane in the polystyrene solution is 1:(1500-1800).

[0039] Optionally, the molecular weight of the polystyrene is 150,000-200,000.

[0040] Optionally, the pH after neutralization with ethanolamine is 6-8.

[0041] Optionally, the amount of the initiator used accounts for 0.5-2wt% of the amount of methyl allyl alcohol polyoxyethylene ether used;

[0042] The amount of Lewis acid used in step S20 accounts for 1-2 wt % of the amount of aminosulfonic acid used;

[0043] The amount of Lewis acid used in step S30 is 2-5 wt % of the polystyrene.

[0044] Optionally, the initiator is selected from azobisisobutyronitrile.

[0045] Optionally, the Lewis acid is AlCl 3 , BF 3 、SbCl 5 , FeBr 3 、FeCl 3 SnCl 4 、TiCl 4 、ZnCl 2 Any one or more of .

[0046] Optionally, the hydrogen-to-oil volume ratio in the reactor B (9) is 200-800, and the pressure is 1-6 MPa, preferably, the pressure is 2-5 MPa.

[0047] Optionally, the hydrogen-oil volume ratio in the reactor C (10) is 300-1000, and the volume space velocity is 0.6-6h -1 , the pressure is 1-6MPa, preferably, the pressure is 2-5MPa.

[0048] Optionally, the iodine value of the saturated fatty acid glyceride is not greater than 5g / 100g or the bromine value is not greater than 5g / 100g, and the sulfur content is lower than 6ppm.

[0049] Optionally, the solid catalyst b is a molybdenum-nickel-based alumina catalyst, and the molybdenum-nickel-based alumina catalyst is any one of a molybdenum-nickel alumina catalyst and a cobalt-molybdenum-nickel alumina catalyst.

[0050] Optionally, the hydrogen-oil volume ratio in the reactor D (19) is 300-700, and the volume space velocity is 0.5-5h -1, the pressure is 1-5MPa, preferably, the pressure is 2-4MPa.

[0051] Optionally, the solid catalyst c is a reduced nickel-alumina catalyst.

[0052] On the other hand, the present application provides an application of saturated fatty acid glyceride prepared by the method for preparing saturated oils and by-product alkanes by two-stage low-pressure hydrogenation as described in any of the above in hydrocarbon-based biodiesel, high-end lubricating oils, emulsifiers, etc.

[0053] Compared with the prior art, the present invention achieves at least one of the following beneficial effects:

[0054] 1. The two-stage low-pressure hydrogenation method of the present application adopts a hot oil circulation direct heating method. Most traditional processes directly heat the bio-oil to a high temperature, which is easy to cause overheating and thermal cracking side reactions. This method adopts a hot oil circulation and bio-oil mixed heating method, which can better control the heating temperature and avoid excessive thermal cracking reactions. At the same time, the method of directly mixing the hot circulating return material with the cold feed to even the temperature solves the problem of easy coking during the heating process of inferior raw materials.

[0055] 2. The two-stage low-pressure hydrogenation method of the present application has an innovative reactor combination method. The traditional process generally uses only a single reactor. This method uses a preparation system composed of multiple reactors and hot and cold separators, which can precisely control the reaction conditions and paths at different stages, improve the reaction selectivity, and prepare better quality products.

[0056] 3. The two-stage low-pressure hydrogenation method of the present application is carried out under ultra-low pressure of 1-6 MPa and low temperature of 160-300°C. Compared with the existing oil hydrogenation process, the lower operating pressure and temperature are conducive to saving equipment investment and reducing operating costs, and can also avoid the deterioration of the raw oil during the reaction process, thereby improving the quality stability and production safety of mass-produced finished products.

[0057] 4. The two-stage low-pressure hydrogenation method of the present application has a liquid catalyst with strong activity, small dosage and low cost. Due to its unique shape, it will not have the essential problems of solid catalysts such as active centers being occupied or covered, and pores being blocked by dust and carbon deposits, and the use environment and conditions are very relaxed.

[0058] 5. The two-stage low-pressure hydrogenation method of the present application utilizes unstable and low-quality oils such as waste cooking oil, sewage oil, and animal fats, expands the scope of bio-oil-based raw materials, and provides a new engineering technology method for the resource recycling of unstable and low-quality oils, which is not only beneficial to environmental governance, but also contributes to the green carbon cycle.

[0059] 6. The two-stage low-pressure hydrogenation method of the present application has a high product yield and quality, thereby improving the process economy. BRIEF DESCRIPTION OF THE DRAWINGS

[0060] The drawings described herein are used to provide a further understanding of the present invention and constitute a part of the present invention. The exemplary embodiments of the present invention and their descriptions are used to explain the present invention and do not constitute an improper limitation of the present invention. In the drawings:

[0061] Figure 1 The process flow chart of the method for preparing saturated oil and fat with by-product alkanes by two-stage low-pressure hydrogenation of the present application is shown.

[0062] The following are the descriptions of the reference numerals:

[0063] 1-unstable inferior oil buffer tank, 2-unstable inferior oil booster pump, 3-circulating oil heating furnace, 4-liquid catalyst storage tank, 5-liquid catalyst filling pump, 6-reactor A, 7-esterification oil intermediate tank, 8-esterification oil booster pump, 9-reactor B, 10-reactor C, 11-hot high-pressure separator, 12-circulating pump, 13-hot low-pressure separator, 14-vacuum fractionating tower, 15-normal pressure fractionating tower, 16-oil residue separator, 17-boosting pump, 18-isomerized feed heating furnace, 19-reactor D, 20-cooler, 21-cold high-pressure separator, 22-cold low-pressure separator, 23-stripping stabilization tower, 24-circulating hydrogen compressor, 25-fresh hydrogen compressor, 26-oil residue separation tank. DETAILED DESCRIPTION

[0064] In order to more clearly illustrate the overall concept of the present invention, a detailed description is given below in an exemplary manner in conjunction with the accompanying drawings.

[0065] In the following description, many specific details are set forth to facilitate a full understanding of the present invention. However, the present invention may also be implemented in other ways different from those described herein. Therefore, the protection scope of the present invention is not limited to the specific embodiments disclosed below.

[0066] Example 1

[0067] refer to Figure 1 This embodiment discloses a method for preparing saturated oil and fat and by-product alkanes by two-stage low-pressure hydrogenation, comprising the following steps:

[0068] (1) The unstable inferior grease buffer tank 1 pumps the unstable inferior grease raw material into the reactor A6 without solid catalyst loading through the unstable inferior grease pressure pump 2. Before entering the reactor A6, the unstable inferior grease raw material is mixed with the circulating oil and heated to 240-300°C by the circulating oil, and mixed with 200-2000ppm of liquid catalyst a. The liquid catalyst a is mixed with the raw material through the liquid catalyst storage tank 4 and the liquid catalyst filling pump 5.

[0069] Specifically, unstable low-quality oils are biological oils containing unsaturated bonds, and the unstable low-quality oils include any one or more of kitchen waste oil, sewage oil, acidified oil or palm oil. The unstable low-quality oil raw materials are first preheated to improve their fluidity, filtered to remove mechanical impurities, and the pressure is increased by a pressure pump. Then, they are mixed with the circulating material from this system to even out the temperature, and the heat brought by the circulating material is used to increase the temperature of the cold feed.

[0070] Afterwards, hydrogen is mixed in to form a mixed hydrogen feed with a hydrogen-oil volume ratio of 400-1000 in the reactor A6. Before the mixed hydrogen raw material enters the reactor, 200-2000ppm of liquid catalyst a is added by a catalyst filling pump to become a mixed hydrogen feed. Preferably, the reactor A6 is a one-stage homogeneous hydrogenation reactor. The liquid catalyst a includes a catalyst body and an auxiliary agent, and the catalyst body is a nickel-based room temperature fluid catalyst, and the nickel-based room temperature fluid catalyst includes any one of a nickel-molybdenum fluid catalyst, a nickel-cobalt fluid catalyst, and a nickel-molybdenum-cobalt fluid catalyst. Among them, the nickel-molybdenum fluid catalyst has a nickel: molybdenum molar ratio of 1: (0.1-0.5), the nickel-cobalt fluid catalyst has a nickel: cobalt molar ratio of 1: (0.2-0.6), and the nickel-molybdenum-cobalt fluid catalyst has a nickel: molybdenum: cobalt molar ratio of 1: (0.03-0.1): (0.04-0.08). The weight of the auxiliary agent is 80-200% of the weight of the catalyst body, and the auxiliary agent is selected from a chelating agent, an organic acid, a carboxylate, and 1,2,4-triazole in a weight ratio of 1:(0.1-0.2):(0.1-0.2):(0.1-0.2).

[0071] The unstable inferior oil raw material mixed with the circulating oil enters the reactor A6 and reacts at a temperature of 240-300°C, a hydrogen-oil volume ratio of 400-1000, a pressure of 1-6MPa and a liquid catalyst concentration of 200-2000ppm, preferably a pressure of 2-5MPa. Under these conditions, the reactor A6 performs a pretreatment of the colloid, sulfur, chlorine and other components in the feed with homogeneous catalysis as the characteristic.

[0072] (2) The effluent from reactor A6 is heated to 180-240°C and enters reactor B9 which is not loaded with solid catalyst.

[0073] Specifically, the total chlorine content in the effluent of reactor A6 is <80ppm, and the sulfur content is <60ppm. Before the effluent of reactor A6 enters reactor B9, it is mixed with the esterified oil separated by the vacuum distillation tower 14 and the oil residue separator 16 from the rear end of the process through the esterified oil intermediate tank 7 and the esterified oil booster pump 8, so that the temperature is 180-240°C, and becomes the feed of reactor B9. Preferably, reactor B9 is a one-stage two-stage homogeneous hydrogenation reactor. Due to the high efficiency of the liquid catalyst, reactor B9 can achieve deep pretreatment of the raw materials without adding any catalyst again. The hydrogen-oil volume ratio in reactor B9 is 200-800, and the pressure is generally 1-6MPa, preferably, the pressure is 2-5MPa.

[0074] (3) The effluent from the reactor B9 enters the reactor C10 filled with the solid catalyst b.

[0075] Specifically, the total chlorine content in the effluent of reactor B9 is <10ppm, and the sulfur content is <15ppm. Preferably, reactor C10 is a fixed bed reactor, and the feed is subjected to mild hydrogenation of unsaturated double bonds. The hydrogen-to-oil volume ratio in reactor C10 is 300-1000, and the volume space velocity is 0.6-6h -1 , the pressure is 1-6MPa, preferably, the pressure is 2-5MPa. The solid catalyst b is a molybdenum-nickel-based alumina catalyst, and the molybdenum-nickel-based alumina catalyst is any one of a molybdenum-nickel alumina catalyst and a cobalt-molybdenum-nickel alumina catalyst; wherein the molybdenum:nickel molar ratio in the molybdenum-nickel alumina catalyst is 0.3-(0.6:1), and the cobalt-molybdenum-nickel alumina catalyst has a cobalt:molybdenum:nickel molar ratio of 0.1-(0.3:0.3)-(0.6:1). Under the action of the catalyst and hydrogenation conditions, the saturation index iodine value of the final product saturated fatty acid glyceride is not greater than 5g / 100g or the bromine value is not greater than 5g / 100g.

[0076] (4) The effluent from the reactor C10 enters the hot high-pressure separator 11. The gas phase flowing out of the hot high-pressure separator 11 is cooled by the cooler 20 and then passes through the cold high-pressure separator 21 and the cold low-pressure separator 22 in sequence to enter the gas stabilization tower 23. Acid gas flows out from the top of the gas stabilization tower 23, and the effluent from the bottom enters the atmospheric pressure fractionation tower 15 to obtain saturated fatty acids, saturated fatty acid glycerides and by-product alkanes;

[0077] The liquid phase flowing out of the hot high-pressure separator 11 is divided into two parts. One part passes through the circulating pump 12 as circulating oil to the circulating oil heating furnace 3, and then enters the reactor A6 again after mixing with the liquid catalyst a. The rest passes through the hot low-pressure separator 13 and then enters the vacuum distillation tower 14.

[0078] Specifically, the outflow of the reactor C10 passes through the hot high-pressure separator 11 to separate the gas phase from the liquid phase. The liquid phase material separated by the hot high-pressure separator 11 is divided into two parts, one part is used as circulating oil, pressurized by the circulating pump 12, and then heated by the circulating oil heating furnace 3, and returned to the feed end of the system to mix with the unstable and inferior oil and fat of the raw material to achieve uniform temperature; the other part is used as the feed of the vacuum fractionation tower 14 through the hot high-pressure separator for component cutting, and through the vacuum fractionation operation, two products of saturated fatty acids and fatty acid glycerides are cut.

[0079] The remaining components at the bottom of the vacuum fractionation tower 14 are esterified oils enriched with various impurities. This part of the bottom oil is centrifuged or filtered by the oil residue separator 16. After the impurities are separated and removed, the light components enter the esterified oil intermediate tank 7 for storage, and then return to the feed end through the esterified oil booster pump 8 to mix with the effluent of the reactor A6, and then reprocess it as the feed component of the reactor B9. This can significantly improve the utilization rate of raw materials, and also has the effect of diluting the reactor feed, which increases the fluidity of the reaction materials and improves the yield of saturated oil.

[0080] After being cooled by the cooler 20, the hot high-pressure separator 11 enters the cold high-pressure separator 21 to separate the material into a gas phase and a liquid phase. The gas phase component is mainly excess residual hydrogen. This part of hydrogen is pressurized by the circulating hydrogen compressor 24 and returned to the feed end to be mixed with the unstable and inferior oil of the raw material for recycling; the liquid phase component is relatively light oil. This part of light oil is first passed through the cold low-pressure separator 22 and then stripped with water vapor in the stripping stabilization tower 23 to remove non-ideal components dissolved in the liquid phase material, that is, gaseous compounds converted from sulfur, chlorine, oxygen, nitrogen and other elements in the raw material during the hydrogenation process, as well as short-chain hydrocarbons that affect the stability of the product oil, etc. The material after gas stripping is then heated and heated to enter the atmospheric distillation tower 15, where saturated fatty acids and fatty acid glycerides can be distilled out.

[0081] (5) The bottom oil of the vacuum distillation tower 14 is centrifuged and filtered through an oil residue separator 16 to obtain a light component which is circulated back to the esterification oil intermediate tank 7, and then directly mixed with the outflow of the reactor A6 without heat exchange, heated to 180-240° C., and then enters the reactor B9 without solid catalyst loading. The outflow of the vacuum distillation tower 14 obtains saturated fatty acids and saturated fatty acid glycerides;

[0082] (6) The bottom oil of the atmospheric distillation tower 15 is pressurized by the booster pump 17 and heated by the isomerization feed heating furnace 18, then mixed with fresh hydrogen and enters the reactor D19 filled with solid catalyst c at 160-240°C. The effluent of the reactor D (19) is mixed with the gas phase of the hot high-pressure separator 11 and cooled before entering the cold high-pressure separator 21.

[0083] Specifically, the reactor D19 is a fixed bed isomerization reactor, the hydrogen to oil volume ratio in the reactor D19 is 300-700, and the volume space velocity is 0.5-5h -1 , the pressure is 1-5MPa, preferably, the pressure is 2-4MPa. The solid catalyst c is a reduced nickel-alumina catalyst. The bottom oil of the atmospheric distillation tower 15 is pressurized by a pressure pump, fresh hydrogen from a fresh hydrogen compressor 25 is added, and then heated by an isomerization feed heating furnace 18 and enters the reactor D19. Under the selective action of the catalyst, the isomerization of straight-chain alkanes is achieved.

[0084] Specifically, the material isomerized by the reactor D19 is combined with the gas phase material at the outlet of the hot high-pressure separator 11, and passes through the same path as the gas phase material, and finally enters the atmospheric distillation tower 15 as the feed of the atmospheric distillation tower 15. After the isomerization reaction and distillation process, the process by-product alkane can be achieved, wherein the C15-C18 alkane content (mass fraction) is greater than 75%, the fatty acid methyl ester content (mass fraction) is less than 1.5%, and the cetane number is greater than 65.

[0085] The bottom effluent of the reactor B9 and the reactor C10 is also separated from the oil residue discharge system through the oil residue separation tank 26. In this step, the chelating agent and the metal elements removed therefrom can be discharged out of the system along with the oil residue without affecting the subsequent separation and treatment of intermediates and products.

[0086] The following Example 2 and Comparative Example all adopt the method of Example 1 for production and processing, and the specific production parameters are different as follows:

[0087] Example 2

[0088] The unstable and low-quality oil raw material used is waste cooking oil:

[0089] (1) The waste cooking oil is heated to 240°C and enters a reactor A without solid catalyst loading. The hydrogen-to-oil volume ratio in the reactor A is 400 and the pressure is 1 MPa. 200 ppm of liquid catalyst a is mixed before entering the reactor A. The liquid catalyst a includes a catalyst body and an auxiliary agent. The catalyst body is a nickel-molybdenum fluid catalyst (nickel-molybdenum molar ratio is 1:0.1). The weight of the auxiliary agent is 80% of the weight of the catalyst body. The auxiliary agent is composed of a chelating agent, salicylic acid, zinc acetate and 1,2,4-triazole in a weight ratio of 1:0.1:0.2:0.1.

[0090] The preparation method of the chelating agent is:

[0091] S10: dissolving 2,4,6-trivinyl-2,4,6-trimethylcyclotrisiloxane and methyl allyl alcohol polyoxyethylene ether in a molar ratio of 1:1 in a mixed solvent of tetrahydrofuran, water and ethanol (the volume ratio of tetrahydrofuran, water and ethanol is 3:1:1), adding azobisisobutyronitrile and reacting at 80° C. for 4 hours to obtain a solution containing intermediate A, wherein the amount of azobisisobutyronitrile accounts for 0.5wt% of the amount of methyl allyl alcohol polyoxyethylene ether;

[0092] S20: adding aminosulfonic acid and BF3 to the solution containing intermediate A, wherein the molar ratio of aminosulfonic acid to methyl allyl alcohol polyoxyethylene ether is 0.8:1, and the amount of BF3 accounts for 1wt% of the amount of aminosulfonic acid; reacting at 125°C for 4h, then cooling to 100°C, adding ethanolamine for neutralization to obtain a solution containing intermediate B, and the pH after neutralization with ethanolamine is 6;

[0093] S30: adding a chloroform solution of polystyrene and BF3 to a solution containing intermediate B, wherein the molar ratio of polystyrene to 2,4,6-trivinyl-2,4,6-trimethylcyclotrisiloxane in the chloroform solution of polystyrene is 1:1800, the molecular weight of polystyrene is 200,000, and the amount of BF3 is 2wt% of the polystyrene, and reacting at 80°C to obtain a crude product;

[0094] S40: The crude product is washed with tetrahydrofuran, purified by soaking in chloroform, and dried to obtain a chelating agent.

[0095] (2) The effluent from reactor A is heated to 180°C and enters reactor B without solid catalyst loading. The hydrogen-to-oil volume ratio in reactor B is 200 and the pressure is 1 MPa;

[0096] (3) The effluent from reactor B enters reactor C filled with molybdenum-nickel alumina catalyst (molybdenum-nickel molar ratio is 0.3). The hydrogen-oil volume ratio in reactor C is 300, and the volume space velocity is 3h -1 , pressure is 6MPa;

[0097] (4) Same as Example 1;

[0098] (5) The esterified oil is directly mixed with the effluent of reactor A without heat exchange, heated to 180°C, and then enters reactor B without solid catalyst loading;

[0099] (6) The bottom oil of the atmospheric distillation tower (15) is pressurized by a pressure pump (17) and heated by an isomerization feed heating furnace (18), and then mixed with fresh hydrogen and enters a fixed bed reactor D filled with a reduced nickel-alumina solid catalyst at 160°C. The hydrogen-to-oil volume ratio of the reactor D is 300, and the volume space velocity is 0.5h -1 , the pressure is 5MPa.

[0100] Example 3

[0101] The unstable and inferior grease raw material used is sewage oil:

[0102] (1) The sewage oil is heated to 270° C. and enters a reactor A without solid catalyst loading. Before entering the reactor A, 1000 ppm of liquid catalyst a is mixed. The liquid catalyst a comprises a catalyst body and an auxiliary agent. The catalyst body is a nickel-cobalt fluid catalyst (nickel-cobalt molar ratio is 1:0.2). The auxiliary agent is the same as that in Example 2. The hydrogen-oil volume ratio in the reactor A is 700 and the pressure is 4 MPa.

[0103] (2) The effluent from reactor A is heated to 210°C and enters reactor B without solid catalyst loading. The hydrogen-to-oil volume ratio in reactor B is 500 and the pressure is 4 MPa;

[0104] (3) The effluent from reactor B enters reactor C filled with molybdenum-nickel alumina catalyst (molybdenum-nickel molar ratio is 0.6). The hydrogen-oil volume ratio in reactor C is 1000 and the volume space velocity is 6 h -1 , pressure is 1MPa;

[0105] (4) Same as Example 1;

[0106] (5) The esterified oil is directly mixed with the effluent of reactor A without heat exchange, heated to 210°C, and then enters reactor B without solid catalyst loading;

[0107] (6) The bottom oil of the atmospheric distillation tower (15) is pressurized by a pressure pump (17) and heated by an isomerization feed heating furnace (18), and then mixed with fresh hydrogen and enters a reactor D filled with a reduced nickel-alumina solid catalyst at 240°C. The hydrogen-to-oil volume ratio of the reactor D is 700, and the volume space velocity is 5h -1 , the pressure is 1MPa.

[0108] Example 4

[0109] The unstable and inferior oil raw material used is acidified oil:

[0110] (1) The acidified oil is heated to 300° C. and enters a reactor A without solid catalyst loading. The volume ratio of hydrogen to oil in the reactor A is 1000 and the pressure is 6 MPa. 2000 ppm of liquid catalyst a is mixed before entering the reactor A. The liquid catalyst a comprises a catalyst body and an auxiliary agent. The catalyst body is a nickel-molybdenum-cobalt fluid catalyst (nickel-molybdenum-cobalt molar ratio is 1:0.03:0.04). The weight of the auxiliary agent is 200% of the weight of the catalyst body. The auxiliary agent is composed of a chelating agent, glycolic acid, sodium tartrate and 1,2,4-triazole in a weight ratio of 1:0.2:0.1:0.2.

[0111] The preparation method of the chelating agent is:

[0112] S10: dissolving 2,4,6-trivinyl-2,4,6-trimethylcyclotrisiloxane and methyl allyl alcohol polyoxyethylene ether in a molar ratio of 2:1 in a mixed solvent of tetrahydrofuran, water and ethanol (the volume ratio of tetrahydrofuran, water and ethanol is 3:1:1), adding azobisisobutyronitrile and reacting at 70°C for 6h to obtain a solution containing intermediate A, wherein the amount of azobisisobutyronitrile accounts for 2wt% of the amount of methyl allyl alcohol polyoxyethylene ether;

[0113] S20: adding aminosulfonic acid and BF3 to the solution containing intermediate A, wherein the molar ratio of aminosulfonic acid to methyl allyl alcohol polyoxyethylene ether is 1.2:1, and the amount of BF3 accounts for 2wt% of the amount of aminosulfonic acid; reacting at 120°C for 5h, then cooling to 100°C, adding ethanolamine for neutralization to obtain a solution containing intermediate B, and the pH after neutralization with ethanolamine is 8;

[0114] S30: adding a chloroform solution of polystyrene and BF3 to a solution containing intermediate B, wherein the molar ratio of polystyrene to 2,4,6-trivinyl-2,4,6-trimethylcyclotrisiloxane in the chloroform solution of polystyrene is 1:1500, the molecular weight of polystyrene is 150,000, and the amount of BF3 is 5wt% of the polystyrene, and reacting at 100° C. to obtain a crude product;

[0115] S40: The crude product is washed with tetrahydrofuran, purified by soaking in chloroform, and dried to obtain a chelating agent.

[0116] (2) The effluent from reactor A is heated to 240°C and enters reactor B without solid catalyst loading. The hydrogen-to-oil volume ratio in reactor B is 800 and the pressure is 6 MPa;

[0117] (3) The effluent from reactor B enters reactor C filled with cobalt-molybdenum-nickel alumina catalyst (cobalt-molybdenum-nickel molar ratio of 0.1:0.3:1). The hydrogen-to-oil volume ratio in reactor C is 600, and the volumetric space velocity is 0.6 h -1 , pressure is 2MPa;

[0118] (4) Same as Example 1;

[0119] (5) The esterified oil is directly mixed with the effluent of reactor A without heat exchange, heated to 240°C, and then enters reactor B without solid catalyst loading;

[0120] (6) The bottom oil of the atmospheric distillation tower (15) is pressurized by a pressure pump (17) and heated by an isomerization feed heating furnace (18), and then mixed with fresh hydrogen and enters a reactor D filled with a reduced nickel-alumina solid catalyst at 220°C. The hydrogen-to-oil volume ratio of the reactor D is 600, and the volume space velocity is 0.5h-1 , the pressure is 2MPa.

[0121] Example 5

[0122] The unstable and low-quality oil raw material used is palm oil:

[0123] (1) Palm oil was heated to 260° C. and entered into a reactor A without solid catalyst loading. 1500 ppm of liquid catalyst a was mixed into the reactor before entering the reactor. The liquid catalyst a included a catalyst body and an auxiliary agent. The catalyst body was a nickel-molybdenum-cobalt fluid catalyst (nickel-molybdenum molar ratio was 1:0.5). The hydrogen-oil volume ratio in the reactor A was 800, the pressure was 5 MPa, and the auxiliary agent was the same as that in Example 4.

[0124] (2) The effluent from reactor A is heated to 230°C and enters reactor B without solid catalyst loading. The hydrogen-to-oil volume ratio in reactor B is 600 and the pressure is 2 MPa;

[0125] (3) The effluent from reactor B enters reactor C filled with cobalt-molybdenum-nickel alumina catalyst (cobalt-molybdenum-nickel molar ratio of 0.3:0.6:1). The hydrogen-to-oil volume ratio in reactor C is 800, and the volumetric space velocity is 2 h -1 , pressure is 5MPa;

[0126] (4) Same as Example 1;

[0127] (5) The esterified oil is directly mixed with the effluent of reactor A without heat exchange, heated to 190°C, and enters reactor B without solid catalyst loading;

[0128] (6) The bottom oil of the atmospheric distillation tower (15) is pressurized by a pressure pump (17) and heated by an isomerization feed heating furnace (18). After mixing with fresh hydrogen, the oil enters the reactor D filled with a reduced nickel-alumina solid catalyst at 200°C. The hydrogen-to-oil volume ratio of the reactor D is 500, and the volume space velocity is 2h -1 , the pressure is 4MPa.

[0129] Example 6

[0130] The unstable and low-quality oil raw material used is waste cooking oil:

[0131] (1) The waste cooking oil was heated to 250° C. and entered a reactor A without solid catalyst loading. Before entering the reactor, 1200 ppm of liquid catalyst a was mixed in. The liquid catalyst a included a catalyst body and an additive. The catalyst body was a nickel-cobalt fluid catalyst (nickel-molybdenum molar ratio was 1:0.6). The hydrogen-oil volume ratio in the reactor A was 600, the pressure was 4 MPa, and the additive was the same as that in Example 4.

[0132] (2) The effluent from reactor A is heated to 210°C and enters reactor B without solid catalyst loading. The hydrogen-to-oil volume ratio in reactor B is 800 and the pressure is 3 MPa;

[0133] (3) The effluent from reactor B enters reactor C filled with cobalt-molybdenum-nickel alumina catalyst (cobalt-molybdenum-nickel molar ratio of 0.2:0.5:1). The hydrogen-to-oil volume ratio in reactor C is 500, and the volumetric space velocity is 4 h -1 , pressure is 4MPa;

[0134] (4) Same as Example 1;

[0135] (5) The esterified oil is directly mixed with the effluent of reactor A without heat exchange, heated to 230°C, and then enters reactor B without solid catalyst loading;

[0136] (6) The bottom oil of the atmospheric distillation tower (15) is pressurized by a pressure pump (17) and heated by an isomerization feed heating furnace (18), and then mixed with fresh hydrogen and enters a reactor D filled with a reduced nickel-alumina solid catalyst at 190°C. The hydrogen-to-oil volume ratio of the reactor D is 400, and the volume space velocity is 4h -1 , the pressure is 3MPa.

[0137] Example 7

[0138] The unstable and low-quality oil raw material used is palm oil:

[0139] (1) Palm oil was heated to 260° C. and entered into a reactor A without solid catalyst loading. 1000 ppm of liquid catalyst a was mixed into the reactor before entering the reactor. The liquid catalyst a included a catalyst body and an auxiliary agent. The catalyst body was a nickel-molybdenum-cobalt fluid catalyst (nickel-molybdenum-cobalt molar ratio was 1:0.03:0.08). The hydrogen-oil volume ratio in the reactor A was 8000, the pressure was 6 MPa, and the auxiliary agent was the same as that in Example 4.

[0140] (2) The effluent from reactor A is heated to 240°C and enters reactor B without solid catalyst loading. The hydrogen-to-oil volume ratio in reactor B is 300 and the pressure is 5 MPa;

[0141] (3) The effluent from reactor B enters reactor C filled with cobalt-molybdenum-nickel alumina catalyst (cobalt-molybdenum-nickel molar ratio of 0.3:0.4:1). The hydrogen-to-oil volume ratio in reactor C is 400, and the volumetric space velocity is 5 h -1 , pressure is 3MPa;

[0142] (4) Same as Example 1;

[0143] (5) The esterified oil is directly mixed with the effluent of reactor A without heat exchange, heated to 220°C, and then enters reactor B without solid catalyst loading;

[0144] (6) The bottom oil of the atmospheric distillation tower (15) is pressurized by a pressure pump (17) and heated by an isomerization feed heating furnace (18), and then mixed with fresh hydrogen and enters a reactor D filled with a reduced nickel-alumina solid catalyst at 180°C. The hydrogen-to-oil volume ratio of the reactor D is 400, and the volume space velocity is 2h -1 , the pressure is 2MPa.

[0145] Example 8

[0146] Compared with Example 5, the difference is that the pressure and hydrogen-to-oil ratio of reactor A are different, the pressure is 7 MPa, and the hydrogen-to-oil ratio is 350.

[0147] Example 9

[0148] Compared with Example 5, the difference is that the pressure and hydrogen-to-oil ratio of reactor B are different, the pressure is 7 MPa, and the hydrogen-to-oil ratio is 150.

[0149] Example 10

[0150] Compared with Example 5, the difference is that the liquid catalyst a only includes the catalyst body but does not include the additive.

[0151] Embodiment 11

[0152] Compared with Example 5, the difference is that the auxiliary agent does not contain a chelating agent.

[0153] The product indicators obtained in Examples 2-11 are shown in Tables 1 and 2, where the metal content refers to the total content of iron, chromium, copper, lead and vanadium metals.

[0154] Table 1 Saturated fatty acid glycerides

[0155]

[0156] Table 2 By-product alkanes

[0157]

[0158] As shown in Tables 1 and 2, in the products obtained in Examples 2-7, the yield of saturated fatty acid glyceride is between 85% and 90%, and the yield of saturated fatty acid glyceride in Example 5 is as high as 90%. The iodine value of saturated fatty acid glyceride is between 3.5 and 4.8, the bromine value of saturated fatty acid glyceride is between 3.1 and 4.5, and the sulfur content, chlorine content and metal content of the product obtained after treatment are all low, which proves that the production process of the present application can not only improve the yield of saturated products, but also improve the quality of products.

[0159] In addition, the method of the above embodiment does not cause coking during continuous production, which means that the production method can be continuously produced and improve production efficiency. The acid value of the product prepared in the above embodiment is also low, which proves that the quality of the product is high and the shelf life is extended.

[0160] The above description is only an embodiment of the present invention and is not intended to limit the present invention. For those skilled in the art, the present invention may have various modifications and variations. Any modification, equivalent substitution, improvement, etc. made within the spirit and principle of the present invention should be included in the scope of the claims of the present invention.

Claims

1. A method for preparing saturated oils and by-product alkanes by two-stage low-pressure hydrogenation, characterized in that: The method comprises the following steps: (1) The unstable low-quality oil raw material is heated to 240-300° C. and enters a reactor A (6) without solid catalyst loading, and 200-2000 ppm of liquid catalyst a is mixed into the reactor A (6); (2) The effluent from reactor A (6) is heated to 180-240°C and enters reactor B (9) without solid catalyst loading; (3) The effluent from reactor B (9) enters reactor C (10) filled with solid catalyst b; (4) The effluent from the reactor C (10) enters the hot high-pressure separator (11), and the gas phase flowing out of the hot high-pressure separator (11) is cooled and sequentially passes through the cold high-pressure separator (21) and the cold low-pressure separator (22) and enters the stripping stabilization tower (23). Acid gas flows out from the top of the stripping stabilization tower (23), and the effluent from the bottom enters the atmospheric pressure fractionation tower (15) to obtain saturated fatty acids, saturated fatty acid glycerides and by-product alkanes; The liquid phase flowing out of the hot high-pressure separator (11) is divided into two parts, one part is used as circulating oil to go to the circulating oil heating furnace (3), mixed with the liquid catalyst a and then enters the reactor A (6) again, and the rest passes through the hot low-pressure separator (13) and then enters the vacuum distillation tower (14); (5) The bottom oil of the vacuum distillation tower (14) is centrifuged and filtered to obtain a light component which is recycled to the esterification oil intermediate tank (7), and then directly mixed with the effluent of the reactor A (6) without heat exchange, heated to 180-240° C., and then enters the reactor B (9) without solid catalyst loading. The side stream of the vacuum distillation tower (14) is saturated fatty acids and saturated fatty acid glycerides; (6) The bottom oil of the atmospheric distillation tower (15) is pressurized by a booster pump (17) and heated by an isomerization feed heating furnace (18), then mixed with fresh hydrogen and enters a reactor D (19) filled with a solid catalyst c at 160-240°C. The effluent of the reactor D (19) is mixed with the gas phase of the hot high-pressure separator (11), cooled, and then enters a cold high-pressure separator (21).

2. The method for preparing saturated oils and by-product alkanes by two-stage low-pressure hydrogenation according to claim 1, characterized in that: The unstable low-quality oil is a biological oil containing unsaturated bonds, and the unstable low-quality oil includes any one or more of kitchen waste oil, sewage oil, acidified oil or palm oil.

3. The method for preparing saturated oils and by-product alkanes by two-stage low-pressure hydrogenation according to claim 1, characterized in that: The hydrogen-oil volume ratio in the reactor A (6) is 400-1000, and the pressure is 1-6 MPa, preferably, the pressure is 2-5 MPa.

4. The method for preparing saturated oils and by-product alkanes by two-stage low-pressure hydrogenation according to claim 1, characterized in that: The liquid catalyst a comprises a catalyst body and an additive, wherein the catalyst body is a nickel-based room temperature fluid catalyst, and the nickel-based room temperature fluid catalyst comprises any one of a nickel-molybdenum fluid catalyst, a nickel-cobalt fluid catalyst, and a nickel-molybdenum-cobalt fluid catalyst; The weight of the auxiliary agent is 80-200% of the weight of the catalyst body, and the auxiliary agent is selected from a chelating agent, an organic acid, a carboxylate, and 1,2,4-triazole in a weight ratio of 1:(0.1-0.2):(0.1-0.2):(0.1-0.2).

5. The method for preparing saturated oil and fat and by-product alkanes by two-stage low-pressure hydrogenation according to claim 1, characterized in that: The preparation method of the chelating agent is: S10: dissolving 2,4,6-trivinyl-2,4,6-trimethylcyclotrisiloxane and methyl allyl alcohol polyoxyethylene ether in a solvent, adding an initiator and reacting at 70-80° C. for 4-6 hours to obtain a solution containing intermediate A; S20: adding sulfamic acid and Lewis acid to the solution containing intermediate A, reacting at 120-125° C. for 4-5 hours, then cooling to 100° C., adding ethanolamine for neutralization to obtain a solution containing intermediate B; S30: adding a polystyrene solution and a Lewis acid to the solution containing the intermediate B, and reacting at 80-100° C. to obtain a crude product; S40: washing, purifying and drying the crude product to obtain the chelating agent.

6. The method for preparing saturated oil and fat and by-product alkanes by two-stage low-pressure hydrogenation according to claim 1, characterized in that: The hydrogen-oil volume ratio in the reactor B (9) is 200-800, and the pressure is 1-6 MPa, preferably, the pressure is 2-5 MPa.

7. The method for preparing saturated oils and by-product alkanes by two-stage low-pressure hydrogenation according to claim 1, characterized in that: The hydrogen-oil volume ratio in the reactor C (10) is 300-1000, and the volume space velocity is 0.6-6h -1 , the pressure is 1-6MPa, preferably, the pressure is 2-5MPa.

8. The method for preparing saturated oil and fat with by-product alkanes by two-stage low-pressure hydrogenation according to claim 1, characterized in that: The hydrogen-oil volume ratio in the reactor D (19) is 300-700, and the volume space velocity is 0.5-5h -1 , the pressure is 1-5MPa, preferably, the pressure is 2-4MPa.

9. The method for preparing saturated oils and fats and by-product alkanes by two-stage low-pressure hydrogenation according to claim 1, characterized in that: The solid catalyst b is a molybdenum-nickel-based alumina catalyst, and the molybdenum-nickel-based alumina catalyst is any one of a molybdenum-nickel alumina catalyst and a cobalt-molybdenum-nickel alumina catalyst; The solid catalyst c is a reduced nickel-alumina catalyst.

10. Use of saturated fatty acid glyceride prepared by the method for preparing saturated oils and fats and by-producing alkanes by two-stage low-pressure hydrogenation as claimed in any one of claims 3 to 7 in hydrocarbon-based biodiesel, high-end lubricating oil, emulsifier, etc.

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