A method for producing biomass low-carbon alkanes from biomass raw materials

Abstract

A method for producing biomass-derived light alkanes from biomass feedstock. The biomass feedstock, after being hydrotreated, is sequentially passed through a first hydrocracking reaction zone, a second hydrocracking reaction zone, and a supplemental refining reaction zone for reaction. The resulting hydrogenation reaction effluent is separated to produce light alkanes and a C4+ fraction. The entire C4+ fraction is recycled back to the first hydrocracking reaction zone for further reaction. The method of the present invention can convert the biomass feedstock into biomass-derived light alkanes, or separate methane, ethane, and propane products.

Description

Technical Field

[0001] The present invention relates to the field of biomass raw material processing, and in particular to a method for producing biomass low-carbon alkanes from biomass raw materials. Background Art

[0002] Ethane and propane dehydrogenation processes offer advantages such as high olefin product yields, short plant processes, and low investment costs. In recent years, they have experienced rapid development in North America and the Middle East, regions rich in ethane and propane resources. Traditional ethane and propane resources primarily come from oil or gas fields, with some derived from light hydrocarbons as refinery by-products. However, proven oil and gas fossil energy reserves are limited, and the petroleum refining process is energy-intensive and emits significant carbon emissions. Therefore, developing clean, renewable resources to produce products such as ethane and propane as a supplement or alternative to fossil energy is of great practical significance for achieving a green and clean feedstock for ethylene and propylene production.

[0003] Biomass renewable oils and fats are animal and plant oils or waste cooking oils and fats. They are a renewable resource with abundant sources and large usage. Existing technologies have disclosed relevant technologies for using renewable oils and fats to produce transportation fuels (such as jet fuel and diesel).

[0004] CN108441260A discloses a hydrogen countercurrent oil hydrogenation process, in which biorenewable oil is contacted with a hydrotreating catalyst in the presence of hydrogen to undergo a first-stage hydrotreating to obtain refined oil. The refined oil is then contacted with an isomerization catalyst in the presence of hydrogen to undergo a second-stage hydroisomerization pour point decompression treatment, and then separated by distillation to obtain biomass fuel. Fresh hydrogen is introduced into the second hydrodecompression stage, and the hydrogen-rich gas after the reaction is circulated to the first-stage hydrotreating. The hydrogen-rich gas obtained after the first-stage hydrotreating is subjected to impurity removal treatment and then enters the second-stage hydrotreating together with the fresh hydrogen. An amine washing device is used for gas purification.

[0005] CN106318433A and CN106256880A disclose a method for hydrogenating animal and vegetable oils and fats. In a hydrogenation reactor, hydrogen and a control gas, vegetable oil or animal and vegetable oils and fats are contacted with a sulfided hydrogenation catalyst and reacted under hydrogenation reaction conditions to obtain a hydrogenation reaction product. The control gas is CO or CO2, the partial pressure of hydrogen in the reaction zone is 1.0 to 20.0 MPa, and the ratio of the hydrogen partial pressure to the control gas partial pressure is 2 to 300. The hydrogenation reaction product is subjected to gas-liquid separation to obtain a liquid hydrocarbon stream / gas stream and water. The liquid stream is C8-C 24 Normal alkanes. By controlling the partial pressure of the gas, the reaction is directed toward deoxygenation, reducing the decarbonization of the raw material and increasing the deoxygenation reaction, thereby reducing the CO and CO2 content of the reaction products and retaining carbon atoms in the liquid hydrocarbon products.

[0006] CN106281729A discloses a method for hydrotreating renewable raw materials. In the presence of hydrogen and water, vegetable oil or animal or vegetable fat is contacted with a sulfurized hydrotreating catalyst and reacted under hydrotreating reaction conditions to obtain a hydrotreating reaction product. The mass of water entering the hydrotreating reactor accounts for 1% to 20% of the mass of the vegetable oil or animal or vegetable fat, preferably 3% to 14%. One or more of hydrogen sulfide, CS2, dimethyl disulfide, methyl sulfide, n-butyl sulfide, and thiophene are mixed into the animal or vegetable fat before entering the hydrotreating reactor.

[0007] CN106190286A discloses a method for preparing jet fuel. Hydrogen and animal and vegetable oils are first subjected to a hydrodeoxygenation reaction to produce normal alkanes. The hydrodeoxygenation reaction product and hydrogen are then subjected to an isomerization-depression reaction in an isomerization-depression reaction zone provided with zones A and B, so that the mass fraction of isoalkanes in the product reaches 50% or more while the normal alkanes and isoalkanes are 100%. The isomerization reaction product is refined in a hydrorefining reaction zone, and the reaction effluent is separated and fractionated to obtain jet fuel.

[0008] CN102206502A discloses a method for co-refining animal and plant oils and oxygen-containing compounds to produce aromatic hydrocarbons and light olefins. The animal and plant oils and oxygen-containing compounds are introduced into a catalytic cracking reactor for reaction. The reaction products are separated to obtain liquefied gas, gasoline and other fractions. The liquefied gas is subjected to gas separation to obtain light olefins, and the gasoline is subjected to aromatic extraction to obtain aromatic products.

[0009] As can be seen from the aforementioned prior art, existing biomass feedstock processing technologies, particularly hydrogenation technologies, primarily focus on producing low-aromatic solvent oil, jet fuel, diesel, and other products from biomass feedstocks. Therefore, it is imperative to develop hydrogenation methods that can directly produce low-carbon alkanes from biomass feedstocks, thereby achieving green, clean, and low-carbon production of ethylene and propylene feedstocks. Summary of the Invention

[0010] The present invention aims to solve the technical problem of how to effectively utilize biomass raw materials to produce biomass low-carbon alkanes.

[0011] The method provided by the present invention for producing biomass low-carbon alkanes from biomass raw materials comprises: mixing the biomass raw materials after hydrogenation treatment with hydrogen and sequentially passing through a first hydrocracking reaction zone, a second hydrocracking reaction zone, and a supplementary refining reaction zone for reaction; cooling the obtained hydrogenation reaction effluent and entering a cold high-pressure separator for separation; the liquid phase stream obtained by the cold high-pressure separator is separated and then enters a cold low-pressure separator for further separation; the hydrogen-rich high-fraction gas obtained by the cold high-pressure separator is separated and enters a membrane separation unit for hydrogen recovery; the tail gas of the membrane separation unit and the stream obtained by the cold low-pressure separator are both entered into a depropanizer for separation; the C4+ fraction obtained by separation at the bottom of the depropanizer is recycled back to the first hydrocracking reaction zone for further reaction; the top stream of the depropanizer enters a deethanizer, propane is separated from the bottom of the deethanizer, and methane and ethane are separated from the top gas of the deethanizer after hydrogen recovery through a PSA.

[0012] The first hydrocracking reaction zone is filled with an optional hydrogenation protective agent and a reduced hydrocracking catalyst I, the acidic component of which is amorphous silica-alumina. The second hydrocracking reaction zone is filled with a hydrocracking catalyst II, the acidic component of which is a medium-pore molecular sieve. The supplementary refining reaction zone is filled with a non-precious metal hydrorefining catalyst.

[0013] In the present invention, the obtained low-carbon alkanes are methane, ethane, and propane, which are prepared from biomass raw materials and belong to biomass low-carbon alkane products.

[0014] In the present invention, the biomass raw materials include one or more of various animal and plant oils and fats, and waste cooking oils.

[0015] The plant and animal 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 from vegetable oils and / or animal fats through an ester exchange process. The glycerides include triglycerides, diglycerides, and monoglycerides. The plant oils include, but are not limited to, one or more of soybean oil, rapeseed oil, cottonseed oil, corn oil, rice bran oil, sunflower oil, peanut oil, castor oil, sesame oil, prickly ash seed oil, tea oil, coconut oil, olive oil, Pistacia chinensis oil, palm oil, tung oil, jatropha oil, rosin oil, and Chinese tallow tree oil.

[0016] Waste cooking oil refers to waste oil that is no longer suitable for consumption and is generated during the processing and consumption of animal and plant oils. This includes fatty acids and acidified oils produced during the production of edible oils from oilseeds; frying oil, kitchen waste oil, slop oil, and other types of waste cooking oil generated by households, hotels, restaurants, and food production companies; animal fat produced as a byproduct of meat production and processing; and edible oil that has exceeded its shelf life.

[0017] Because the biomass feedstock contains certain impurities such as sulfides, nitrides, and oxides, it requires hydrotreatment to remove these impurities before entering the first hydrocracking reaction zone of the present invention. However, the present invention is not limited to the specific hydrotreatment process; any hydrotreatment process that can reduce the impurity content of the biomass feedstock to a certain level is suitable for the present invention.

[0018] In a preferred embodiment of the present invention, the sulfur content, nitrogen content and oxygen content in the biomass feedstock after hydrogenation are all less than 20 μg / g.

[0019] In a preferred embodiment of the present invention, the final distillation point of the biomass feedstock after hydrogenation is preferably less than 300°C.

[0020] In order to obtain small molecular ethane and propane products, the raw materials need to be cracked multiple times. The cracking reaction releases a large amount of heat and increases the temperature. For the safe and stable operation of the device, the reaction process needs to reasonably control the activity of the catalyst and the distribution of reaction heat. The present invention uses two hydrocracking catalysts with different acidic components for grading. On the one hand, the cracking activity of the two acidic materials is suitable, and the cracking reaction temperature can be operated stably. Secondly, the two acidic catalysts have different cracking functions and can play a role in distributing the cracking reaction heat. In one embodiment of the present invention, the catalyst filling volume ratio of the first hydrocracking reaction zone to the second hydrocracking reaction zone is 3:1 to 1:3.

[0021] To remove small olefins produced by the hydrocracking reaction unit and improve the concentration and stability of products such as ethane and propane, the present invention provides a supplementary refining reaction zone downstream of the second hydrocracking reaction zone. This zone uses a non-precious metal hydrorefining catalyst for supplementary refining, removing small amounts of byproduct olefins generated by the hydrocracking reaction unit. In one embodiment of the present invention, the catalyst loading volume ratio between the second hydrocracking reaction zone and the supplementary refining reaction zone is 2:1 to 15:1.

[0022] In one embodiment of the present invention, based on the overall catalyst in the first hydrocracking reaction zone, the volume fractions of the hydrogenation protectant and the hydrocracking catalyst I are: 0% to 20%; 80% to 100%, respectively.

[0023] In one embodiment of the present invention, the hydrogenation protective agent comprises a carrier and an active metal component supported on the carrier, the carrier is selected from one or more of aluminum oxide, silicon oxide and titanium oxide, the active metal component is selected from one or more of Group VIB metals and Group VIII non-precious metals, and the active metal component is 0.1 to 15% by weight, calculated as oxide, based on the weight of the hydrogenation protective agent. The particle size of the hydrogenation protective agent is 0.5 to 50.0 mm, and the bulk density is 0.3 to 1.2 g / cm 3 , with a specific surface area of 50 to 300 m 2 / g.

[0024] In one embodiment of the present invention, the hydrocracking catalyst I contains a support and a hydrogenation-active metal component. The support contains alumina or silica-alumina. Based on the support, the content of alumina is 20 to 80% by weight, and the content of silica-alumina is 80 to 20% by weight. The hydrogenation-active metal component is selected from any one or more of Mo, Co, Ni, W, V, Zn, Pt and Pd, and preferably contains Pt and / or Pd. Based on the total weight of the hydrocracking catalyst I as 100%, the content of the hydrogenation-active metal component, calculated as oxide, is 0.1 to 20% by weight.

[0025] In one embodiment of the present invention, the hydrocracking catalyst II contains a support and a hydrogenation-active metal component, the support contains a mesoporous molecular sieve and alumina, the hydrogenation-active metal component is selected from any one or more of Mo, Co, Ni, W, V, Zn, Pt and Pd, preferably contains Pt and / or Pd, based on the total weight of the hydrocracking catalyst II, the content of the mesoporous molecular sieve is 20 to 80 weight%, the content of alumina is 15 to 75 weight%, and the content of the hydrogenation-active metal component is 0.1 to 10 weight% calculated as oxide; the mesoporous molecular sieve is selected from one or more of ZSM-5, ZSM-11, ZSM-12, ZSM-22, ZSM-23, ZSM-35, ZSM-38, SAPO-11 and SAPO-41.

[0026] In one embodiment of the present invention, the non-precious metal hydrorefining catalyst is a supported catalyst, the carrier is alumina and / or silica-alumina, the hydrogenation active metal component is at least one metal selected from Group VIB and at least one metal selected from Group VIII, the Group VIII metal is selected from nickel and / or cobalt, and the Group VIB metal is selected from molybdenum and / or tungsten. Based on the total weight of the non-precious metal hydrorefining catalyst, the content of the Group VIII metal is 1 to 15% by weight, and the content of the Group VIB metal is 5 to 40% by weight, calculated as oxide.

[0027] In the present invention, the hydrocracking catalyst I and the hydrocracking catalyst II need to be reduced before use. The present invention does not limit the reduction process, and the technical solutions of single reduction or double reduction are applicable to the present invention.

[0028] In one embodiment of the present invention, hydrocracking catalyst I and hydrocracking catalyst II are reduced in the presence of hydrogen to obtain the reduced hydrocracking catalyst I and reduced hydrocracking catalyst II, with a reduction temperature of 300 to 420° C., a reduction time of 8 to 32 hours, and a reduction operating pressure of 0.3 to 3.0 MPa.

[0029] In one embodiment of the present invention, the reaction conditions of the first hydrocracking reaction zone and the second hydrocracking reaction zone are: hydrogen partial pressure of 1.0 MPa to 20.0 MPa, reaction temperature of 355°C to 405°C, preferably 360 to 400°C, liquid hourly volume space velocity of 0.5 h -1 ~6h -1 , the hydrogen-to-oil volume ratio is 300-2000.

[0030] In one embodiment of the present invention, the reaction conditions of the supplementary refining reaction zone are: hydrogen partial pressure of 1.0 MPa to 20.0 MPa, reaction temperature of 230°C to 300°C, liquid hourly volume space velocity of 0.5 h -1 ~15h -1 , the hydrogen-to-oil volume ratio is 300-2000.

[0031] In one embodiment of the present invention, the operating pressure of the cold high-pressure separator is 1.0 MPa to 20.0 MPa, and the operating temperature is 20 to 50°C.

[0032] In one embodiment of the present invention, the membrane separation unit comprises a membrane tube filled with a hollow fiber membrane, and one or more combinations selected from a cooler, a buffer tank, a gas-liquid separator, a cyclone separator, a filter, a coalescer, and a dryer.

[0033] In the present invention, the equipment in the membrane separation unit can be connected in a variety of combinations. One configuration involves cooling the high-fraction gas before entering a cyclone separator for gas-liquid separation. The gas phase is then filtered and dried before entering the membrane tube for separation to produce hydrogen and tail gas. Another configuration involves cooling the high-fraction gas before entering a gas-liquid separator for preliminary separation. The separated gas phase is then processed in a coalescer before entering a membrane separator for separation to produce hydrogen and tail gas.

[0034] In one embodiment of the present invention, the membrane separation unit is controlled to have an inlet pressure of 2.4 MPa to 3.4 MPa, a membrane inlet temperature of 30°C to 80°C, and an inlet pressure difference of 1.0 MPa to 2.5 MPa. The inlet pressure difference refers to the difference between the pressure on the feed side (feed gas) and the pressure on the permeate side (product hydrogen) of the membrane separation unit.

[0035] In one embodiment of the present invention, the liquid stream obtained by the low-pressure separator enters the depropanizer from the lower inlet of the depropanizer, and the gas stream obtained by the low-pressure separator and the tail gas from the membrane separation unit enter the depropanizer from the upper inlet of the depropanizer to be separated together. The bottom stream of the depropanizer is all recycled and enters the first hydrocracking reaction zone together with the biomass feedstock after hydrogenation.

[0036] In a preferred embodiment of the present invention, the C4+ fraction circulation ratio is controlled to be 15% to 40%, where the C4+ fraction circulation ratio refers to the ratio of the total weight of the circulating C4+ fraction to the weight of the fresh raw material based on the weight of the fresh raw material.

[0037] In a preferred embodiment of the present invention, the hydrogen is produced by utilizing one or more of wind power, hydropower, or solar energy.

[0038] Features of the present invention:

[0039] (1) When treating low-sulfur and low-nitrogen biomass feedstock that has been hydrotreated, the prior art uses a sulfurized hydrocracking catalyst for hydrocracking. During operation, in order to maintain the stability of the sulfurized catalyst activity, it is necessary to add a sulfurizing agent during operation. The present invention uses a reduced hydrocracking catalyst, which eliminates the need for sulfur addition.

[0040] (2) The hydrocracking catalysts with two different acidic centers used in the present invention are graded, have suitable cracking activity, good operational stability, high biomass feedstock conversion rate, and high ethane and propane yields.

[0041] (3) The products such as ethane and propane obtained by the present invention require the raw oil to undergo multiple cracking reactions, which release a large amount of heat and increase the temperature rise. The present invention controls the circulation ratio of a certain C4+ fraction, increases the materials in the reaction system, and thus carries more reaction heat, thereby reducing the temperature rise of the cracking reaction, which is beneficial to the stable operation of the device. In addition, reducing the temperature rise of the cracking reaction is also beneficial to reducing the hot spot temperature of the hydrocracking catalyst bed, which is beneficial to improving the yield of the product ethane and propane.

[0042] (4) The membrane separation unit used in the present invention can be operated at room temperature, and no phase change occurs during the separation process, resulting in low separation energy consumption. Secondly, the membrane separation unit can be used to recover high-purity hydrogen from the waste hydrogen discharged during the reaction process, and at the same time, obtain tail gas rich in light hydrocarbons. The tail gas can be further separated to obtain products such as ethane and propane, thereby improving the utilization rate of hydrogen and light hydrocarbon materials. BRIEF DESCRIPTION OF THE DRAWINGS

[0043] Figure 1 This is a schematic diagram of one embodiment of the method for producing biomass low-carbon alkanes from biomass raw materials provided by the present invention. DETAILED DESCRIPTION

[0044] The present invention will be further described below with reference to the accompanying drawings, but the present invention is not limited thereto.

[0045] Figure 1 This is a schematic diagram of one embodiment of the method for producing biomass light alkanes from biomass raw materials provided by the present invention. Figure 1 As shown, the hydrotreated biomass feedstock 1 and hydrogen 2 enter the first hydrocracking reaction zone 3 together, where they react with a hydrogenation protective agent and hydrocracking catalyst I. The reaction effluent enters the second hydrocracking reaction zone 4, where it reacts with a hydrocracking catalyst II. The reaction effluent then enters the supplementary refining reaction zone 5, where it contacts a non-precious metal hydrorefining catalyst for a supplementary refining reaction. The hydrogenation reaction effluent from the supplementary refining reaction zone is cooled by heat exchange with the outlet stream of the hydrogen compressor 9 and then enters a cold high-pressure separator 6 for gas-liquid separation. The hydrogen-rich high-fraction gas separated from the top of the cold high-pressure separator 6 is depressurized and cooled, and then sent to a cyclone separator 7 to remove a small amount of C4+ fraction and impurities such as water. After removal, the gaseous stream enters a dryer 8 and a filter module 10 for dehydration and filtration of C4+ fraction impurities. It is then sent to a membrane tube 11 equipped with a hollow fiber membrane for separation. Hydrogen 12 is separated on the permeate side, and tail gas 14 is obtained on the retentate side. The liquid stream from the cold high-pressure separator 6 enters the cold low-pressure separator 13 for further gas-liquid separation. The liquid stream from the cold low-pressure separator 13 enters the depropanizer 15 through its lower inlet. The gas stream from the cold low-pressure separator 13 and the tail gas 14 from the membrane separation unit enter the depropanizer 15 through its upper inlet for separation. The bottom stream 16 of the depropanizer 15, comprising the C4+ fraction, is recycled and, along with the hydrotreated biomass feedstock, enters the first hydrocracking reaction zone for further reaction. The overhead stream from the depropanizer is cooled and enters the deethanizer 17 for separation. Propane product 18 is separated from the bottom of the deethanizer 17. The overhead gas from the deethanizer 17 is fed to a PSA unit to recover hydrogen 19 and produce a methane and ethane mixed stream 20, which can be further separated into methane and ethane products.

[0046] The present invention will be further described below with reference to the examples, but the present invention is not limited thereto.

[0047] In the examples and comparative examples, the component analysis data of the liquefied gas (propane and butane) samples were obtained according to SH / T0230 “Determination of the composition of liquefied petroleum gas (chromatographic method)”.

[0048] Table 1 lists the main properties of biomass raw oil after hydrotreatment.

[0049] Table 2 lists the physical and chemical properties of the catalysts used in the examples and comparative examples of the present invention. The catalysts with commercial brands are all produced by Sinopec Catalyst Branch.

[0050] The preparation method and composition of the hydrocracking catalyst II (CAT-2) used in the second hydrocracking reaction zone of the present invention are as follows:

[0051] A ZSM-22 molecular sieve (supplied by Changling Catalyst Factory, with a silicon-aluminum ratio of 56) was uniformly mixed with pseudo-boehmite and sesbania powder. A nitric acid aqueous 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.

[0052] Before use, hydrocracking catalyst I and hydrocracking catalyst II were reduced in a hydrogen atmosphere at a reduction temperature of 350° C., a reduction time of 4 hours, and a hydrogen pressure of 0.1 MPa.

[0053] In the embodiments and comparative examples, the high-fraction gas enters the membrane separation unit, and the membrane separation unit is provided with a raw gas feed pipeline, a permeate gas (hydrogen) discharge pipeline and a membrane separation unit tail gas delivery pipeline. The pipelines are connected to a component of a material filled with a hollow fiber membrane to form a membrane separation unit, wherein the hollow fiber membrane is a polysulfone hollow fiber-based membrane coated with a casting liquid formed by a catalyst including polydimethylsiloxane, ethyl orthosilane and dibutyltin dilaurate. The membrane inlet pressure is controlled to be 2.8MPa, the membrane inlet temperature is 45°C, and the membrane inlet pressure difference is 1.45MPa.

[0054] In the examples and comparative examples, the operating pressure of the cold high-pressure separator was the pressure of the reaction unit, and the operating temperature was 45°C.

[0055] Tables 3 and 4 list the data of the embodiments and comparative examples of the present invention, and Table 5 lists the product standard of liquefied petroleum gas GB 11174-2011 "Liquefied Petroleum Gas".

[0056] Examples 1-3

[0057] In Examples 1-3: the biomass raw material 2 after hydrogenation treatment is mixed with hydrogen and then reacted in sequence through the first hydrocracking reaction zone, the second hydrocracking reaction zone, and the supplementary refining reaction zone. The resulting hydrogenation reaction effluent is cooled and then enters a cold high-pressure separator for separation. The liquid phase stream obtained by the cold high-pressure separator enters a cold low-pressure separator for further separation. The high-fraction gas rich in hydrogen obtained by the cold high-pressure separator enters the membrane separation unit for hydrogen recovery. The tail gas of the membrane separation unit and the stream separated by the cold low-pressure separator enter the depropanizer for separation together. The bottom stream of the depropanizer (C4+ fraction) is recycled to the first hydrocracking reaction zone to continue the reaction. The top stream of the depropanizer enters the deethanizer. The bottom stream of the deethanizer is mainly propane. The top gas of the deethanizer is separated into methane and ethane after hydrogen is recovered by PSA. The specific catalyst loading ratio, reaction conditions and product yield are shown in Table 3.

[0058] In Example 1, the C4+ fraction circulation ratio is controlled to be 40.0%, in Example 2, the C4+ fraction circulation ratio is controlled to be 24.0%, and in Example 3, the C4+ fraction circulation ratio is controlled to be 20.0%. The C4+ fraction circulation ratio refers to the ratio of the total weight of the circulating C4+ fraction to the weight of the fresh raw material based on the weight of the fresh raw material.

[0059] As shown in Table 3, the propane yields of Examples 1, 2, and 3 were 42.50%, 50.36%, and 71.07%, respectively. The volume fractions of C3 hydrocarbons in the deethanizer bottoms stream were 97.06%, 98.06%, and 99.44%, respectively, meeting the quality requirements for commercial propane liquefaction in GB 11174-2011, "Liquefied Petroleum Gas." After hydrogen recovery via PSA, the deethanizer overhead gas had high ethane volume fractions of 51.1%, 56.5%, and 55.2%, respectively, demonstrating good utilization value.

[0060] Comparative Example 1

[0061] Comparative Example 1 used the same raw materials, process flow, and catalyst loading scheme as Example 1. Unlike Example 1, this comparative example controlled the C4+ fraction recycle ratio to 62.66%. The specific catalyst loading ratio, reaction conditions, and product yield are shown in Table 3.

[0062] As can be seen from Table 3, the methane yield in the product dry gas in Comparative Example 1 is 18.37%, the ethane yield is 32.88%, and the propane yield is 39.08%. Due to the high volume fractions of the product dry gas and ethane, the separation effect of the deethanizer is reduced, so that the volume fraction of the hydrocarbon component of the propane sample (C3) at the bottom of the deethanizer is only 90.63%, which cannot meet the quality index requirements of commercial propane liquefaction in GB 11174-2011 "Liquefied Petroleum Gas". In addition, under the reaction conditions, the methane product yield is too high, which reduces the selectivity of the target ethane and propane products.

[0063] Examples 4, 5 and 6

[0064] Examples 4, 5 and 6 use the biomass raw oil 1 after hydrogenation treatment.

[0065] After being mixed with hydrogen, the feedstock enters the first hydrocracking reaction zone, where it reacts with a hydrogenation protectant and a reduced hydrocracking catalyst I. The reaction effluent then enters the second hydrocracking reaction zone, where it reacts with a reduced hydrocracking catalyst II. After heat exchange, the reaction effluent from the second hydrocracking reaction zone enters a supplementary refining reaction zone, where it contacts a non-precious metal hydrorefining catalyst for a supplementary refining reaction. The hydrogenation reaction effluent from the supplementary refining reaction zone then enters a cold high-pressure separator, a cold low-pressure separator, a depropanizer, and a deethanizer for separation. The high-fraction gas rich in hydrogen obtained by the cold high-pressure separator enters the membrane separation unit for hydrogen recovery. The tail gas of the membrane separation unit and the stream separated by the cold low-pressure separator enter the depropanizer for separation. The bottom stream of the depropanizer (C4+ fraction) is recycled to the first hydrocracking reaction zone to continue the reaction. The top stream of the depropanizer enters the deethanizer. The bottom stream of the deethanizer is mainly propane. The top gas of the deethanizer recovers hydrogen through PSA and then separates methane and ethane.

[0066] In Example 4, the C4+ fraction recycle ratio was controlled at 37.7%; in Example 5, the C4+ fraction recycle ratio was controlled at 24.32%; and in Example 6, the C4+ fraction recycle ratio was controlled at 21.72%. Specific catalyst loading ratios, reaction conditions, and product yields are shown in Table 4.

[0067] As can be seen from Table 4, the yields of the propane products in Examples 4, 5, and 6 were 55.38%, 60.10%, and 66.09%, respectively; the volume fractions of the C3 hydrocarbon components in the deethanizer bottoms were 97.9%, 99.68%, and 99.92%, respectively, all meeting the commercial propane quality index requirements of GB 11174-2011 "Liquefied Petroleum Gas"; after hydrogen recovery by PSA from the deethanizer overhead gas, the ethane volume fractions in the tail gas were relatively high, namely 53.1%, 49.6%, and 58.3%, respectively, indicating good utilization value.

[0068] Comparative Example 2

[0069] Comparative Example 3 used the same raw materials and process flow as Example 1. Unlike Example 1, this comparative example also loaded hydrocracking catalyst II into the first hydrocracking reaction zone. In this comparative example, the reaction temperature in the first and second hydrocracking reaction zones was controlled at 368°C, and the C4+ fraction recycle ratio was controlled at 41%. However, 30% of the C4+ fraction was still required to achieve equilibrium. Specific catalyst loading ratios, reaction conditions, and product yields are shown in Table 4.

[0070] As can be seen from Table 4, due to the high proportion of C4+ rejection, the volume fractions of the product ethane and propane are relatively low, namely 35.29% and 19.31%, respectively. Therefore, it can be seen that this comparative example only uses a single hydrocracking catalyst and does not adopt the technical solution of the graded hydrocracking catalyst of the present invention, and the ethane and propane yields are relatively low.

[0071] Table 1

[0072]

[0073] Table 2

[0074]

[0075]

[0076] Table 3

[0077]

[0078]

[0079] Table 4

[0080]

[0081] Table 5 GB 11174-2011 Specification requirements for liquefied petroleum gas

[0082]

Claims

1. A method for producing biomass low-carbon alkanes from biomass raw materials, comprising: mixing the biomass raw materials after hydrogenation with hydrogen and reacting in sequence through a first hydrocracking reaction zone, a second hydrocracking reaction zone, and a supplementary refining reaction zone; cooling the resulting hydrogenation reaction effluent and entering a cold high-pressure separator for separation; the liquid phase stream obtained by the cold high-pressure separator entering a cold low-pressure separator for further separation; the hydrogen-rich high-fraction gas obtained by the cold high-pressure separator entering a membrane separation unit for hydrogen recovery; the tail gas from the membrane separation unit and the stream separated by the cold low-pressure separator entering a depropanizer for separation; the C4+ fraction obtained by separation at the bottom of the depropanizer is recycled back to the first hydrocracking reaction zone for further reaction; the overhead stream of the depropanizer entering a deethanizer; propane is separated from the bottom of the deethanizer; the overhead gas of the deethanizer recovers hydrogen through a PSA and then methane and ethane are separated; The first hydrocracking reaction zone is loaded with an optional hydrogenation protective agent and a reduced hydrocracking catalyst I, the acidic component of the hydrocracking catalyst I is amorphous silica-alumina, the second hydrocracking reaction zone is loaded with a hydrocracking catalyst II, the acidic component of the hydrocracking catalyst II is a medium-pore molecular sieve, the catalyst loading volume ratio of the first hydrocracking reaction zone to the second hydrocracking reaction zone is 3:1 to 1:3, the supplementary refining reaction zone is loaded with a non-precious metal hydrorefining catalyst, and the catalyst loading volume ratio of the second hydrocracking reaction zone to the supplementary refining reaction zone is 2:1 to 15:1, Biomass raw materials include one or more of various animal and vegetable oils and catering waste oils.

2. The method according to claim 1, characterized in that The sulfur content, nitrogen content and oxygen content of the biomass raw material after the hydrogenation treatment are all less than 20 μg / g; The final distillation point of the biomass raw material after hydrogenation treatment is less than 300°C.

3. The method according to claim 1, characterized in that Based on the entire catalyst in the first hydrocracking reaction zone, the volume fractions of the hydrogenation protective agent and the hydrocracking catalyst I are: 0%~20%; 80%~100%, respectively.

4. The method according to claim 1, wherein The hydrogenation protective agent comprises a carrier and an active metal component supported on the carrier, wherein the carrier is selected from one or more of aluminum oxide, silicon oxide and titanium oxide, and the active metal component is selected from one or more of Group VIB metals and Group VIII non-precious metals. Based on the weight of the hydrogenation protective agent, the active metal component is 0.1 to 15% by weight in terms of oxide. The particle size of the hydrogenation protective agent is 0.5 to 50.0 mm, and the bulk density is 0.3 to 1.2 g / cm 3 , with a specific surface area of 50~300m 2 / g.

5. The method according to claim 1, wherein The hydrocracking catalyst I contains a carrier and a hydrogenation active metal component. The carrier contains alumina and silica-alumina. Based on the carrier, the content of alumina is 20-80% by weight, and the content of silica-alumina is 80-20% by weight. The hydrogenation active metal component is selected from any one or more of Mo, Co, Ni, W, V, Zn, Pt and Pd. Based on the total weight of the hydrocracking catalyst I being 100%, the content of the hydrogenation active metal component is 0.1-20% by weight, calculated as oxides.

6. The method according to claim 5, wherein the hydrogenation active metal component of the hydrocracking catalyst I is Pt and / or Pd.

7. The method according to claim 1, characterized in that The hydrocracking catalyst II contains a carrier and a hydrogenation-active metal component, wherein the carrier contains a mesoporous molecular sieve and alumina, and the hydrogenation-active metal component is selected from any one or more of Mo, Co, Ni, W, V, Zn, Pt and Pd. Based on the total weight of the hydrocracking catalyst II, the content of the mesoporous molecular sieve is 20-80% by weight, the content of the alumina is 15-75% by weight, and the content of the hydrogenation-active metal component, calculated as oxide, is 0.1-10% by weight. The mesoporous molecular sieve is selected from one or more of ZSM-5, ZSM-11, ZSM-12, ZSM-22, ZSM-23, ZSM-35, ZSM-38, SAPO-11 and SAPO-41.

8. The method according to claim 7, wherein the hydrogenation active metal component of the hydrocracking catalyst II is Pt and / or Pd.

9. The method according to claim 1, characterized in that The non-precious metal hydrorefining catalyst is a supported catalyst, the carrier is alumina or silica-alumina, the hydrogenation active metal component is at least one metal selected from Group VIB and at least one metal selected from Group VIII, the Group VIII metal is selected from nickel and / or cobalt, and the Group VIB metal is selected from molybdenum and / or tungsten. Based on the total weight of the non-precious metal hydrorefining catalyst, the content of the Group VIII metal is 1-15% by weight, and the content of the Group VIB metal is 5-40% by weight, calculated as oxide.

10. The method according to claim 1, characterized in that The reaction conditions of the first hydrocracking reaction zone and the second hydrocracking reaction zone are: hydrogen partial pressure of 1.0 MPa~20.0 MPa, reaction temperature of 355℃~405℃, liquid hourly volume space velocity of 0.5h -1 ~6h -1 , the hydrogen-to-oil volume ratio is 300~2000.

11. The method according to claim 10, characterized in that The reaction temperature of the first hydrocracking reaction zone and the second hydrocracking reaction zone is 360-400°C.

12. The method according to claim 1, characterized in that The reaction conditions of the supplementary refining reaction zone are: hydrogen partial pressure of 1.0MPa~20.0MPa, reaction temperature of 230℃~300℃, liquid hourly volume space velocity of 0.5h -1 ~15h -1 , the hydrogen-to-oil volume ratio is 300~2000.

13. The method according to claim 1, wherein The hydrogen is prepared by utilizing one or more of wind power, hydropower, or solar energy.

14. The method according to claim 1, wherein The operating pressure of the cold high-pressure separator is 1.0MPa~20.0MPa and the operating temperature is 20~50℃.

15. The method according to claim 1, wherein The membrane separation unit comprises a membrane tube filled with a hollow fiber membrane, and one or more combinations selected from a cooler, a buffer tank, a gas-liquid separator, a cyclone separator, a filter, a coalescer, and a dryer; In the membrane separation unit, the membrane inlet pressure is controlled to be 2.4MPa~3.4MPa, the membrane inlet temperature is controlled to be 30℃~80℃, and the membrane inlet pressure difference is controlled to be 1.0MPa~2.5MPa.

16. The method according to claim 1, wherein The liquid phase flow obtained by the low-pressure separator enters the depropanizer from the lower inlet of the depropanizer, and the gas phase flow obtained by the low-pressure separator and the tail gas from the membrane separation unit enter the depropanizer from the upper inlet of the depropanizer and are separated together. The bottom flow of the depropanizer is all recycled and enters the first hydrocracking reaction zone together with the biomass feedstock after hydrogenation treatment.

17. The method according to claim 1, wherein The C4+ fraction circulation ratio is controlled to be 15% to 40%. The C4+ fraction circulation ratio refers to the ratio of the total weight of the circulating C4+ fraction to the weight of the fresh raw material based on the weight of the fresh raw material.

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