A method for producing biomass liquefied gas from biomass raw materials

Abstract

A method for producing biomass-derived liquefied biogas from biomass feedstock. The biomass feedstock, after hydrogenation, 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 liquefied gas and a C5+ fraction. The entire C5+ 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 a biomass-derived liquefied biogas product, or separate propane and butane 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 liquefied gas from biomass raw materials. Background Art

[0002] 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).

[0003] 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.

[0004] CN106318433A and CN106256880A disclose a method for hydrotreating animal and vegetable oils. In a hydrotreating reactor, hydrogen and a control gas, vegetable oil or animal and vegetable oil, and a sulfided hydrotreating catalyst are contacted and reacted under hydrotreating reaction conditions to obtain a hydrotreating reaction product. The control gas is CO or CO2, the partial pressure of hydrogen in the reaction zone is 1.0-20.0 MPa, and the ratio of the hydrogen partial pressure to the control gas partial pressure is 2-300. The hydrotreating reaction product undergoes 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.

[0005] 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%. Before entering the hydrotreating reactor, one or more of hydrogen sulfide, CS2, dimethyl disulfide, methyl sulfide, n-butyl sulfide, and thiophene is mixed with the animal or vegetable fat.

[0006] CN106190286A discloses a method for preparing jet fuel, in which 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.

[0007] 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.

[0008] As can be seen from the existing technologies listed above, hydrogenation technology in existing biomass feedstock processing technologies primarily focuses on using biomass feedstock to produce products such as low-aromatic solvent oil, jet fuel, and diesel. Therefore, it is imperative to develop hydrogenation methods that can directly produce liquefied petroleum gas (LPG) or propane from biomass feedstock, thereby achieving green, clean, and low-carbon production of civilian LPG or commercial LPG (propane and butane) feedstock. Summary of the Invention

[0009] The present invention aims to solve the technical problem of how to effectively utilize biomass raw materials to produce biomass liquefied gas.

[0010] The present invention provides a method for producing biomass liquefied gas from biomass raw materials. The biomass raw materials after hydrogenation treatment are mixed with hydrogen and then enter a first hydrocracking reaction zone to contact with an optional hydrogenation protective agent and a reduced hydrocracking catalyst I for reaction. The reaction effluent directly enters a second hydrocracking reaction zone to contact with a reduced hydrocracking catalyst II for reaction. The acidic component of the hydrocracking catalyst I is amorphous silica-alumina, and the acidic component of the hydrocracking catalyst II is a medium-pore molecular sieve. The reaction temperature of the first hydrocracking reaction zone and the second hydrocracking reaction zone is 350-420°C.

[0011] The reaction effluent from the second hydrocracking reaction zone enters the supplementary refining reaction zone after heat exchange and contacts with a non-precious metal hydrorefining catalyst at a reaction temperature of 200-340°C for supplementary refining reaction.

[0012] The hydrogenation reaction effluent obtained in the supplementary refining reaction zone is separated to obtain liquefied gas and a C5+ fraction. The entire C5+ fraction is recycled to the first hydrocracking reaction zone to continue the reaction. The C5+ fraction circulation ratio is controlled to be 19% to 63%. The C5+ fraction circulation ratio refers to the ratio of the total weight of the recycled C5+ fraction to the weight of the fresh feedstock, based on the weight of the fresh feedstock.

[0013] The final distillation point of the biomass raw material after hydrogenation treatment is less than 340°C.

[0014] In the present invention, the liquefied gas obtained mainly includes propane and butane, which is prepared from biomass raw materials and is a biomass liquefied gas product. If further separated, propane and butane products can be obtained separately.

[0015] In the present invention, the biomass raw materials include one or more of various animal and plant oils and fats, and waste cooking oils.

[0016] 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.

[0017] 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.

[0018] 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.

[0019] 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.

[0020] In a preferred embodiment of the present invention, the final distillation point of the biomass feedstock after hydrotreatment is preferably less than 310°C.

[0021] To ensure stable operation and increase liquefied gas yield, the present invention utilizes two hydrocracking catalysts with different acidic compositions. This grading not only controls cracking activity, but also allows for the production of either propane or butane. In one embodiment of the present invention, the catalyst loading volume ratio between the first hydrocracking reaction zone and the second hydrocracking reaction zone is 3:1 to 1:3.

[0022] To further improve the quality of the liquefied gas, the present invention incorporates a supplementary refining reaction zone downstream of the second hydrocracking reaction zone, employing a non-precious metal hydrorefining catalyst for supplementary refining. The purpose of supplementary refining is to remove the small amount of olefins, such as ethylene and propylene, produced in the hydrocracking reaction zone, thereby reducing the olefin content in the liquefied gas product. This not only ensures that propane and butane concentrations meet quality standards, but also improves the storage and operational stability of the liquefied gas. 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.

[0023] 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.

[0024] 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 50~300m 2 / g.

[0025] In one embodiment of the present invention, the hydrocracking catalyst I comprises a support and a hydrogenation-active metal component. The support comprises 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.

[0026] 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, calculated as oxide, is 0.1 to 10 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.

[0027] 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.

[0028] 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.

[0029] 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.

[0030] 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°C to 400°C, liquid hourly volume space velocity of 0.5 h -1 ~6h -1 , the hydrogen-to-oil volume ratio is 300~2000.

[0031] 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 320°C, liquid hourly volume space velocity of 0.5 h -1~15h -1 , the hydrogen-to-oil volume ratio is 300~2000.

[0032] In one embodiment of the present invention, the hydrogenation reaction effluent obtained in the supplementary refining reaction zone enters the high-pressure separator and the low-pressure separator in sequence for gas-liquid separation, the liquid phase stream obtained by the low-pressure separator enters the debutanizer from the lower inlet of the debutanizer, and the gas phase stream obtained by the low-pressure separator enters the debutanizer from the upper inlet of the debutanizer, and they are separated together. The bottom stream of the debutanizer is circulated and enters the first hydrocracking reaction zone together with the biomass feedstock after hydrogenation treatment.

[0033] In a preferred embodiment of the present invention, the hydrogen is produced by utilizing one or more of wind power, hydropower, or solar energy.

[0034] Features of the present invention:

[0035] (1) When processing low-sulfur and low-nitrogen biomass feedstock that has been hydrotreated, the prior art uses a sulfurized hydrocracking catalyst for hydrocracking. During the operation of the device, in order to maintain the stability of the activity of the sulfurized catalyst, it is necessary to add a sulfurizing agent during operation. The present invention uses a reduced hydrocracking catalyst, which does not require sulfur addition.

[0036] (2) The present invention utilizes amorphous silica-alumina hydrocracking catalyst I in the first hydrocracking reaction zone to achieve feedstock lightweighting. In the second hydrocracking reaction zone, a medium-pore size-selective molecular sieve hydrocracking catalyst II is utilized to further convert the unconverted normal paraffin products and small molecular weight isoparaffins in the first hydrocracking unit, thereby producing a large proportion of propane and butane products. Furthermore, the two hydrocracking catalysts with different acidic centers employed in the present invention are graded, resulting in suitable cracking activity and good operational stability.

[0037] (3) The main product of the present invention is biomass liquefied gas. The degree of cracking of the raw materials is deep and the temperature rise of the cracking reaction is large. By controlling a certain C5+ fraction circulation ratio, the present invention can not only increase the mass flow rate of the liquid phase material in the reaction system and effectively reduce the reaction temperature rise, but also inhibit the reduction of product selectivity caused by excessive cracking of the raw hydrocarbon molecules, thereby improving the selectivity of the target product propane or butane. BRIEF DESCRIPTION OF THE DRAWINGS

[0038] Figure 1 It is a schematic diagram of one embodiment of the method for producing biomass liquefied gas from biomass raw materials provided by the present invention. DETAILED DESCRIPTION

[0039] The present invention will be further described below with reference to the accompanying drawings, but the present invention is not limited thereto.

[0040] Figure 1 This is a schematic diagram of one embodiment of the method for producing biomass liquefied gas 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, where they react with a hydrogenation protective agent and hydrocracking catalyst I. The reaction effluent then enters the second hydrocracking reaction zone 4, where it reacts with hydrocracking catalyst II. The reaction effluent then enters the supplemental refining reaction zone 5, where it contacts a non-precious metal hydrorefining catalyst for supplemental refining. The hydrogenation effluent from the supplemental refining reaction zone is cooled by heat exchange with the outlet stream of the circulating hydrogen compressor and then enters a cold high-pressure separator 6 for gas-liquid separation. The high-fraction gas 7 separated from the top of the cold high-pressure separator is recovered, mixed with supplemental hydrogen 8, and then recirculated to the circulating hydrogen compressor. The liquid stream from the cold high-pressure separator enters a cold low-pressure separator 9 for further gas-liquid separation. The gaseous material from the top of the cold low-pressure separator 9 enters a debutanizer 10 via the upper inlet of the debutanizer, while the liquid material from the bottom of the cold low-pressure separator 9 enters the debutanizer 10 via the lower inlet of the debutanizer, where both are separated. The bottom stream 11 (C5+ fraction) of the debutanizer is circulated, mixed with the biomass feedstock oil 1 after hydrogenation, and then enters the first hydrocracking reaction zone. The top stream of the debutanizer enters the deethanizer 12 for separation. The top material 13 of the deethanizer is dry gas, and the bottom material 14 of the deethanizer is liquefied gas (propane and butane), which can be directly used as a product delivery device or further separated to obtain propane product and butane product.

[0041] The present invention will be further described below with reference to the examples, but the present invention is not limited thereto.

[0042] In the examples and comparative examples, the component analysis data of the liquefied petroleum gas (propane and butane) samples were obtained according to SH / T 0230 “Determination of the composition of liquefied petroleum gas (chromatographic method)”.

[0043] Table 1 lists the main properties of biomass raw oil after hydrotreatment.

[0044] 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.

[0045] 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:

[0046] 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 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.

[0047] 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.

[0048] Tables 3, 4 and 5 list the data of the embodiments and comparative examples of the present invention, and Table 6 lists the product standard of liquefied petroleum gas GB 11174-2011 "Liquefied Petroleum Gas".

[0049] Example 1

[0050] After being hydrotreated, the biomass feedstock 1 is mixed with hydrogen and then enters the first hydrocracking reaction zone, where it contacts a hydrogenation protectant and a reduced hydrocracking catalyst I for a reaction. The reaction effluent then directly enters the second hydrocracking reaction zone, where it contacts a reduced hydrocracking catalyst II for a reaction. 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. In this embodiment, the C5+ fraction recycle ratio is controlled to be 63.0%. The C5+ fraction recycle ratio refers to the ratio of the total weight of the recycled C5+ fraction to the weight of the fresh feedstock, based on the weight of the fresh feedstock.

[0051] The hydrogenation effluent from the supplemental refining reaction zone enters the high-pressure separator and the low-pressure separator sequentially for gas-liquid separation. The liquid stream from the low-pressure separator enters the debutanizer through the lower inlet of the debutanizer, while the gas stream from the low-pressure separator enters the debutanizer through the upper inlet of the debutanizer for separation. The debutanizer bottoms stream, the C5+ fraction, is recycled and fed into the first hydrocracking reaction zone along with the hydrogenated biomass feedstock. The overhead stream from the debutanizer enters the deethanizer for separation into dry gas (methane and ethyl acetate) and liquefied gas (propane and butane). Specific catalyst loading ratios, reaction conditions, and product yields are shown in Table 3.

[0052] Example 2

[0053] This example uses the same raw materials, process flow, and catalyst loading scheme as Example 1, except that the C5+ fraction recycle ratio is controlled at 54.0%. Specific catalyst loading ratios, reaction conditions, and product yields are shown in Table 3.

[0054] Example 3

[0055] This example uses the same raw materials, process flow, and catalyst loading scheme as Example 1, except that the C5+ fraction recycle ratio is controlled at 42.19%. Specific catalyst loading ratios, reaction conditions, and product yields are shown in Table 3.

[0056] As can be seen from Table 3, the product propane yields of Examples 1, 2, and 3 were 41.70%, 42.31%, and 44.06%, respectively; the product butane yields were 20.98%, 17.26%, and 16.80%, respectively; and the volume fractions of hydrocarbon components in the deethanizer bottom liquefied gas (C3+C4) were 98.06%, 98.04%, and 98.0%, respectively, all of which met the quality index requirements of commercial propane and butane liquefied gas in GB 11174-2011 "Liquefied Petroleum Gas."

[0057] Comparative Example 1

[0058] 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 C5+ fraction recycle ratio to 87.0%. The specific catalyst loading ratio, reaction conditions, and product yield are shown in Table 3.

[0059] As can be seen from Table 3, the methane yield in the product of Comparative Example 1 is 8.97%, the ethane yield is 21.21%, the propane yield is 38.4%, and the butane yield is 26.2%. Although the volume fraction of the hydrocarbon components in the liquefied gas (C3+C4) at the bottom of the deethanizer is 98.01%, which also meets the quality index requirements of commercial propane and butane liquefaction in GB 11174-2011 "Liquefied Petroleum Gas", the low primary conversion rate of fresh feedstock and the large circulation volume increase the operating costs of the device. Secondly, when the primary conversion rate of fresh feedstock is reduced, the product methane and ethane yields are higher, and the selectivity of the product propane and butane is correspondingly reduced, and the process selectivity is not reasonable enough.

[0060] Examples 4, 5 and 6

[0061] Examples 4, 5, and 6 utilize hydrotreated biomass feedstock 2. This feedstock is mixed with hydrogen and then fed into the first hydrocracking reaction zone, where it reacts with a hydrogenation protectant and a reduced hydrocracking catalyst I. The reaction effluent then directly enters the second hydrocracking reaction zone, where it reacts with a reduced hydrocracking catalyst II. The reaction effluent from the second hydrocracking reaction zone, after heat exchange, enters a supplementary refining reaction zone, where it contacts a non-precious metal hydrorefining catalyst for a supplementary refining reaction. The hydrorefining effluent from the supplementary refining reaction zone sequentially enters a high-pressure separator and a low-pressure separator for gas-liquid separation. The liquid stream from the low-pressure separator enters the debutanizer through the lower inlet of the debutanizer, while the gas stream from the low-pressure separator enters the debutanizer through the upper inlet of the debutanizer, where both are separated. The debutanizer bottoms stream, comprising the C5+ fraction, is recycled and fed into the first hydrocracking reaction zone along with the hydrotreated biomass feedstock. The top stream of the debutanizer enters the deethanizer and is separated into dry gas (methane and ethane) and liquefied gas (propane and butane).

[0062] In Example 4, the C5+ fraction recycle ratio was controlled at 35.34%; in Example 5, the C5+ fraction recycle ratio was controlled at 22.82%; and in Example 6, the C5+ fraction recycle ratio was controlled at 36.13%. Specific catalyst loading ratios, reaction conditions, and product yields are shown in Table 4.

[0063] As can be seen from Table 4, the product propane yields of Examples 4, 5, and 6 were 60.10%, 56.91%, and 58.30%, respectively; the volume fractions of the C3 hydrocarbon components in the liquefied gas (propane) at the bottom of the deethanizer were 99.81%, 99.47%, and 99.93%, respectively, all meeting the commercial propane quality index requirements of GB 11174-2011 "Liquefied Petroleum Gas".

[0064] Comparative Example 2

[0065] Comparative Example 2 also used the same raw materials, process flow, catalyst, and loading scheme as Example 4. Unlike Example 4, the C5+ fraction recycle ratio was controlled at 12.13%. The specific catalyst loading ratio, reaction conditions, and product yield are shown in Table 4.

[0066] As can be seen from Table 4, the methane yield in the product of Comparative Example 2 is 13.83%, the ethane yield is 39.10%, and the propane yield is 46.80%. Although the volume fraction of the C3 hydrocarbon component in the propane sample at the bottom of the deethanizer is 95.37%, which also meets the quality index requirements of commercial propane liquefaction in GB11174-2011 "Liquefied Petroleum Gas", the one-time conversion rate of the fresh raw material is too high, the product methane and ethane yields are high, and the product propane selectivity is correspondingly reduced, and the process is not economical and reasonable.

[0067] Comparative Example 3

[0068] 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 (CAT-2) into the first hydrocracking reaction zone. The C5+ fraction recycle ratio was controlled at 50.02%. Specific catalyst loading ratios, reaction conditions, and product yields are shown in Table 5.

[0069] As can be seen from Table 5, the methane yield of the product in Comparative Example 3 is 6.34%, the ethane yield is 15.43%, the propane yield is 28.53%, and the butane yield is 14.64%. Since the hydrocracking catalyst II in the first hydrocracking reaction unit is difficult to achieve light weighting of isoparaffins and cyclic hydrocarbons in the feed, and the hydrocracking catalyst II in the second hydrocracking reaction unit is also difficult to convert the isoparaffins and cyclic hydrocarbon components in the feed, the cracking depth of the feed is insufficient. Therefore, under the condition of a circulation ratio of 50.02%, it is necessary to discard about 35.06% of the unconverted C5+ fraction rich in isoparaffins and cyclic hydrocarbon components. It can be seen that using only one hydrocracking catalyst without adopting the two-graded hydrocracking catalyst technical solution of the present invention, it is difficult to fully convert the feed into propane and butane products.

[0070] Comparative Example 4

[0071] Comparative Example 4 used the same raw materials and process flow as Example 1. Unlike Example 1, this comparative example used RHC-220 hydrocracking catalyst in both the first and second hydrocracking reaction zones. The acidic component of this catalyst was a Y-type molecular sieve, and the reaction was carried out in a sulfurized state. Under the processing conditions of the raw materials provided by this invention, due to the high cracking activity of the Y-type molecular sieve catalyst and its lack of a clear linear variation in cracking activity in the absence of ammonia, the test achieved near-complete conversion of the raw material at a cracking reaction temperature of only 325°C, with a controlled C5+ fraction recycle ratio of 12.03%. Specific catalyst loading ratios, reaction conditions, and product yields are shown in Table 5.

[0072] As shown in Table 5, Comparative Example 4 uses the same raw materials and process flow as Example 1, but uses a conventional Y-type molecular sieve cracking catalyst. Although the product has a high light hydrocarbon yield and meets the commercial liquefied gas (propane and butane) index requirements of GB 11174-2011 "Liquefied Petroleum Gas", there are problems such as the need to supplement the sulfiding agent with the sulfided catalyst, and the strong cracking reaction activity making it difficult to operate smoothly and control the reaction, which is not conducive to the stable operation of the device.

[0073] Table 1

[0074]

[0075] Table 2

[0076]

[0077] Table 3

[0078]

[0079] Table 4

[0080]

[0081] Table 5

[0082]

[0083] Table 6 GB 11174-2011 Specification requirements for liquefied petroleum gas

[0084]

Claims

1. A method for producing biomass liquefied gas from biomass feedstock, comprising: mixing the hydrogenated biomass feedstock with hydrogen and then entering a first hydrocracking reaction zone, where the feedstock is contacted with an optional hydrogenation protective agent and a reduced hydrocracking catalyst I for reaction; and directly entering a second hydrocracking reaction zone where the feedstock is contacted with a reduced hydrocracking catalyst II for reaction; wherein the acidic component of the hydrocracking catalyst I is amorphous silica-alumina, and the acidic component of the hydrocracking catalyst II is a medium-pore molecular sieve; and wherein the catalyst loading volume ratio of the first hydrocracking reaction zone to the second hydrocracking reaction zone is 3:1 to 1:3; and the reaction temperature of the first hydrocracking reaction zone and the second hydrocracking reaction zone is 350° C. to 420° C. The reaction effluent of the second hydrocracking reaction zone enters the supplementary refining reaction zone after heat exchange and contacts with a non-precious metal hydrorefining catalyst at a reaction temperature of 200-340°C for supplementary refining reaction. The catalyst loading volume ratio of the second hydrocracking reaction zone to the supplementary refining reaction zone is 2:1-15:

1. The hydrogenation reaction effluent obtained in the supplementary refining reaction zone is separated to obtain liquefied gas and a C5+ fraction. The entire C5+ fraction is recycled to the first hydrocracking reaction zone to continue the reaction. The C5+ fraction circulation ratio is controlled to be 19% to 63%. The C5+ fraction circulation ratio refers to the ratio of the total weight of the recycled C5+ fraction to the weight of the fresh feedstock, based on the weight of the fresh feedstock. The biomass raw materials include one or more of various animal and plant oils and catering waste oils, and the final distillation point of the biomass raw materials after hydrogenation treatment is less than 340°C.

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 310°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 1, wherein 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℃~320℃, 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 hydrogenation reaction effluent obtained in the supplementary refining reaction zone enters the high-pressure separator and the low-pressure separator in sequence for gas-liquid separation. The liquid phase flow obtained by the low-pressure separator enters the debutanizer from the lower inlet of the debutanizer, and the gas phase flow obtained by the low-pressure separator enters the debutanizer from the upper inlet of the debutanizer, and is separated together. The bottom flow of the debutanizer is circulated and enters the first hydrocracking reaction zone together with the biomass feedstock after hydrogenation treatment.

Images