A process for the preparation of n-pentene from n-pentane

Abstract

The present application belongs to the technical field of alkane processing, and specifically relates to a process for preparing n-pentene from n-pentane, which comprises an n-pentane dehydrogenation device, an n-pentene separation device, an isomerization reaction and a circulating separation device; the n-pentene separation device separates a hydrogen stream G-H2 and an unreacted n-pentane stream M1 from a dehydrogenated gas stream from the n-pentane dehydrogenation device and cooled, separates an n-pentene stream P-2 and sends the n-pentene stream P-2 to the isomerization reaction and the circulating separation device for treatment; the isomerization reaction and the circulating separation device comprise a pentene separation tower, a 2-pentene heavy component removal tower and a 2-pentene isomerization reactor group; the isomerization reactor group is provided with 2-3 fixed bed reactors VI filled with 2-pentene isomerization catalyst in parallel; the operating conditions of each reactor VI include: a temperature of 375-390 DEG C, a pressure of 0.05-0.15 MPa, a 2-pentene mass space velocity of 1-2 h ‑1 The process has certain application prospect.

Description

Technical Field

[0001] This invention belongs to the field of alkane processing technology, specifically relating to a process for preparing n-pentene from n-pentane. Background Technology

[0002] n-Pentene (1-pentene) is an α-olefin with an odd number of carbon atoms and is an important raw material for the synthesis of various specialty polymers and many fine chemical products, such as 1,2-pentanediol. Typically, n-pentene is obtained by separating raffinate C5 atoms, but the source and quantity of raffinate C5 atoms are problematic, and the separation process is complex, thus limiting its yield.

[0003] With the development and widespread application of dehydrogenation processes for producing olefins from low-carbon alkanes and isomerization processes for low-carbon olefins, a combined process for producing n-pentene from n-pentane dehydrogenation and olefin isomerization has the potential for development and application. n-Pentane is a basic product from the separation of C5 byproducts from catalytic cracking and naphtha cracking for ethylene production, as well as natural gas condensate. There is an ample supply of n-pentane feedstock with a purity of over 95%, even 99%. Under normal conditions, 1-pentene and 2-pentene are produced simultaneously during n-pentane dehydrogenation, but 2-pentene is more stable and thus produced in a higher proportion. Since 1-pentene (n-pentene) has a higher price and application value, separating the feedstock rich in 2-pentene first, and then converting 2-pentene to 1-pentene through an olefin isomerization process to finally obtain the n-pentene product, is a feasible process.

[0004] The dehydrogenation of isobutane and isopentane to produce isobutene and isopentene is a related process for the dehydrogenation of n-pentane to produce n-pentene. For example, in the process for producing isobutene from isobutane by dehydrogenation disclosed in CN103232312A, hydrogen and water vapor are used as diluents and heat carriers. The dehydrogenation reaction of isobutane is carried out in a multi-stage fixed-bed adiabatic reactor packed with a Pt-Sn-K / alumina type platinum-based dehydrogenation catalyst and under inlet gas flow temperatures of 540-600℃. This achieves the target reaction effect of isobutane single-pass conversion of over 30% or even over 35% and isobutene selectivity of over 97%. After the dehydrogenation catalyst gradually accumulates coke and the reaction efficiency decreases, in-situ regeneration within the reactor is required, with a regeneration cycle of over 400 hours or even over 750 hours.

[0005] CN120887769A discloses a process for preparing isopentene by dehydrogenation of isopentane. This process is developed based on the process for producing isobutene by dehydrogenation of isobutane in CN103232312A. It utilizes a zirconium-containing spherical or strip-shaped catalyst H, prepared with improvements based on the proportions and methods in CN103212411A, and a newly developed magnesium-containing spherical or strip-shaped catalyst J. The process includes an isopentane dehydrogenation unit and an isopentene separation unit. The isopentane dehydrogenation unit includes a dehydrogenation reactor group, where the reactor is a fixed-bed adiabatic reactor with an axial or radial reaction structure, and is packed with Pt-Sn-K / alumina. The catalysts are H or J, which are alkane dehydrogenation catalysts. Catalyst H or J exhibits hydromethanization performance under CO2-containing dehydrogenation reaction conditions. Catalyst H, under the setup and operation conditions of dehydrogenation reactor group modes A, B, and C, can achieve a long-term operating effect of isopentane single-pass conversion rate of over 30%, isopentene selectivity of over 90%, and catalyst regeneration cycle of over 300 hours. Catalyst J, under the setup and operation conditions of dehydrogenation reactor group modes A, B, and C, can achieve a long-term operating effect of isopentane single-pass conversion rate of over 30% or even over 35%, isopentene selectivity of over 90%, and catalyst regeneration cycle of over 400 hours or even over 600 hours. The setup and operation of the dehydrogenation reactor group includes any of the following modes: Mode A: 2-4 reactors are set up, each with the same catalyst loading volume, and a heater is installed before the gas flow inlet of each reactor; the reactors are connected in parallel for isopentane dehydrogenation; during the stable operation of isopentane dehydrogenation, the inlet gas flow conditions of each reactor include: temperature 520-550℃, pressure ≤0.2MPa (gauge pressure), and isopentane mass hourly space velocity 1-2h. -1 The molar ratio of isopentane to H2 is 1:(0.5-1.5), and the molar content of CO2 is 1.5-2.5%. The maximum temperature of the catalyst bed in each reactor is controlled below 550℃. The outlet gas flow temperature or the outlet temperature of the catalyst bed in each reactor is controlled above 530℃. Mode B: Set up 2-3 reactor series in parallel. Each reactor series includes two reactors I and II connected in series, where reactor I is the first reactor. The catalyst loading volume ratio in reactors I and II is 1:(0.8-1.2). A heater is set up before the gas flow inlet of each reactor. During the stable operation of isopentane dehydrogenation, the inlet gas flow conditions of reactor I include: temperature 520-550℃, pressure ≤0.2MPa, and isopentane mass hourly space velocity 1.5-3h. -1The molar ratio of isopentane to H2 is 1:(0.5-1), and the molar content of CO2 is 1-2%. The inlet gas flow conditions for reactor II include: temperature 510-540℃, and CO2 molar content 1-2%. The maximum temperature of the catalyst bed in each reactor is controlled below 550℃. The outlet gas flow temperature or the outlet temperature of the catalyst bed in reactor II is controlled above 530℃. Mode C: Set up 2-3 parallel reactor series, each reactor series includes three reactors III, IV, and V connected in series, where reactor III is the first reactor. The catalyst loading volume ratio in reactors III, IV, and V is 1:(0.8-1):(0.8-1). A heater is set up before the gas flow inlet of each reactor. During the stable operation of isopentane dehydrogenation, the inlet gas flow conditions for reactor III include: temperature 550-560℃, pressure ≤0.2MPa, and isopentane mass hourly space velocity 1.5-2.5h. -1 The molar ratio of isopentane to H2 and water is 1:(0.3-1):(1.5-3); the inlet gas flow temperature of reactors IV and V is controlled at 540-550℃; the outlet gas flow temperature of reactor V or the outlet temperature of the catalyst bed is controlled at above 530℃. The isopentene separation unit includes a cooling separation system and a pressure swing adsorption (PSA) separation system. The PSA system includes a PSA-H2 unit for separating hydrogen and one or more PSA-CH units for separating hydrocarbons. In addition to isopentene and hydrogen products with a purity of 98% or even 99%, the main byproduct is n-pentane (n-pentane accounts for about 80% of the isopentane byproducts, calculated by carbon).

[0006] However, due to the differences in properties between n-pentane and isobutane, as well as the differences in properties and compositional distribution among the dehydrogenation products, the separation processes are significantly different. Therefore, the above-mentioned processes for preparing isobutene by dehydrogenation and separation of isobutane and isopentane cannot be simply adapted to the process for preparing n-pentene by dehydrogenation and separation of n-pentane.

[0007] On the other hand, in the existing technology, the process of isomerizing 1-pentene from the 2-pentene-rich stream separated from the n-pentane dehydrogenation stream does not yet meet the economic requirements.

[0008] For example, CN104557407A discloses a method for isomerizing 2-pentene to 1-pentene, which mainly includes the following steps: 1) The residual C5 feedstock enters a precision distillation column (first distillation column), and a mixed feed stream of 1-pentene and isopentane is discharged from the top of the column, while a mixed fraction containing 2-pentene and n-pentane is obtained from the bottom of the column; 2) The bottom feed containing 2-pentene and n-pentane obtained in step 1) enters a second distillation column, n-pentane is distilled off from the top of the column, and a feed stream rich in 2-pentene is obtained from the bottom of the column; 3) The bottom feed containing 2-pentene obtained in step 2) passes through a fixed-bed isomerization reactor, and part of the 2-pentene isomerizes to 1-pentene; 4) The isomerization reaction liquid obtained in step 3) is returned to step 1) and enters the first distillation column, and the 1-pentene generated by the isomerization reaction is separated and enters the mixed feed stream of 1-pentene and isopentane discharged from the top of the column. In step 3), the catalyst packed in the isomerization reactor is crystalline aluminosilicate (selected from ZSM series molecular sieves) with a silicon-to-aluminum molar ratio of SiO2 / Al2O3 of (100-500):1. The aluminosilicate crystal powder has the following specifications: surface area 400-500 m² / g. 2 / g, the main axis of the elliptical pores is 0.6-0.9nm, and the minor axis is 0.5nm; the preferred isomerization reaction conditions are: the mass hourly space velocity of the feed to the isomerization reactor is 0.2-1.2h. -1 The reaction temperature is 370-400℃. The isomerization reaction effect of this method includes: 1-pentene process yield of 68.7-82.4%, and single-pass reaction yield of 10.3-11.7%. However, this method is difficult to apply because: (1) the second distillation column requires a number of trays much higher than 100 and a top reflux ratio of more than 100, resulting in extremely high operating energy consumption and difficulty in stable control; (2) the top of the first distillation column yields a mixture of 1-pentene and isopentane, with a 1-pentene content of about 30wt%, which is not the high-purity product required by the market with a 1-pentene content of more than 95wt% or even more than 98wt%; (3) the byproduct of its isomerization reaction, isopentene (mainly 2-methyl-2-butene), circulates and accumulates in the system due to the lack of a separation outlet; (4) the total content of isopentane and n-pentane in the raw material is 71.7wt%, which leads to very high and ineffective energy consumption in the double-tower distillation process. The atmospheric boiling points of the relevant components are as follows: isopentane 27.9℃, 1-pentene 30.0℃, n-pentane 36.1℃, trans-2-pentene 36.4℃, cis-2-pentene 37.0℃, and 2-methyl-2-butene 38.4℃. The yield of 1-pentene defined in paragraphs 0018-0021 of the specification CN104557407A still has issues.

[0009] Therefore, it is necessary to develop a process for preparing n-pentene from n-pentane, which meets the overall economic requirements through a relatively efficient n-pentane dehydrogenation and 2-pentene double bond transfer isomerization process, as well as a corresponding efficient separation process; the specific requirements include: (1) during the n-pentane dehydrogenation process, the single-pass yield of n-pentene exceeds 25%, and the selectivity of n-pentene exceeds 90%; (2) a high-concentration 2-pentene stream is separated from the n-pentane dehydrogenation product stream, the 2-pentene stream is subjected to double bond transfer isomerization, and then 1-pentene products with a purity of more than 95 wt% or even more than 98 wt% are separated under the condition of a high concentration of n-pentene; (3) during the isomerization reaction, the single-pass yield of 1-pentene converted from 2-pentene exceeds 20%, and the selectivity of 1-pentene exceeds 90%, preferably more than 95%. Summary of the Invention

[0010] To solve the above-mentioned technical problems, the present invention provides a process for preparing n-pentene from n-pentane, including an n-pentane dehydrogenation unit, an n-pentene separation unit, an isomerization reaction and a recycling separation unit;

[0011] The n-pentene separation unit includes a cooling separation system and a pressure swing adsorption separation system. From the dehydrogenated gas stream from the n-pentane dehydrogenation unit after cooling, hydrogen gas stream G-H2 is separated and reused in the n-pentane dehydrogenation unit and sold externally. Unreacted n-pentane gas stream M1 is separated and reused in the n-pentane dehydrogenation unit. n-pentene gas stream P-2 is separated and sent to the isomerization reaction and recycling separation unit for processing.

[0012] The n-pentane dehydrogenation unit includes a fixed-bed adiabatic reactor with a Pt-Sn-K / alumina type alkane dehydrogenation catalyst and an axial or radial reaction structure. The alkane dehydrogenation catalyst has hydromethanation performance under CO2-containing dehydrogenation reaction conditions.

[0013] The isomerization reaction and recycling separation unit includes a pentene separation tower, a 2-pentene de-heavy tower, and a 2-pentene isomerization reactor group. A stream of n-pentene gas, P-2, from the n-pentene separation unit is mixed with a cooled isomerized gas stream from the 2-pentene isomerization reactor group and then enters the pentene separation tower. A 1-pentene liquid product is separated after the top condenser of the pentene separation tower, and the 2-pentene-rich liquid stream obtained at the bottom of the tower enters the 2-pentene de-heavy tower. A heavy hydrocarbon liquid stream, P-4, is separated from the bottom of the 2-pentene de-heavy tower as a byproduct. The product is sold by separating a 2-pentene gas stream before the top condenser and sending it to the isomerization reactor group. The isomerization reactor group consists of 2-3 fixed-bed reactors VI, each packed with a 2-pentene isomerization catalyst and having an axial or radial reaction structure. Each reactor VI has an olefin heater installed before its gas stream inlet. During stable operation, the reaction conditions for each reactor VI include: temperature 375-390℃, pressure 0.05-0.15MPa (gauge pressure, the same below), and 2-pentene mass hourly space velocity (MHV) 1-2h⁻¹.-1 .

[0014] This invention discloses a process for preparing n-pentene from n-pentane. In reactor VI of the isomerization reaction and recycling separation unit, the 2-pentene isomerization catalyst is a zinc-yttrium modified ZSM-5 molecular sieve-alumina, containing 1.4-2.6 wt% zinc (based on ZnO) and 3-5 wt% yttrium (based on Y₂O₃), and is prepared through the following steps:

[0015] (1) A product with a relative crystallinity of 95% or higher, a silicon-to-aluminum ratio (SiO2:Al2O3, mol ratio) of 300-350, and a surface area of ​​500-600 m² 2 / g of Na-type ZSM-5 molecular sieve powder with an average crystal size of 0.3-0.5μm is exchanged 4-5 times with a sufficient amount of NH4NO3 aqueous solution of 1-2mol / L at 80-90℃, washed with water, dried, and calcined at 450-500℃ to transform into H-type HZSM-5 molecular sieve powder.

[0016] (2) Place the HZSM-5 molecular sieve powder obtained in step (1) into a mixer. During stirring, apply a zinc acetate aqueous solution of 95-110% of its inner pore volume (determined by low-temperature nitrogen adsorption method, the pore volume of HZSM-5 molecular sieve powder, multiplied by the amount of feed) and the required concentration by atomization spraying. After homogenization treatment for more than 10 hours, discharge and dry. After drying, calcine the powder at 500-550℃ for 2-4 hours, cool, and crush to obtain zinc-containing modified ZSM-5 molecular sieve powder.

[0017] (3) Add 200-600 mesh boehmite powder to a kneader (anhydrous), start the kneader, and uniformly apply 20-25 wt% acetic acid aqueous solution through atomized spraying. The weight ratio of boehmite powder to acetic acid aqueous solution is 100:(10-15). After the acetic acid aqueous solution is added, stop the kneader 30-60 minutes later and cover and seal for more than 10 hours. Start the kneader again, add modified ZSM-5 molecular sieve powder and guar gum powder, mix the powders evenly, and then spray through atomized spraying. Add water in liquid form, knead for 30-60 minutes, then extrude into strips. Dry the extruded strips, calcine at 500-550℃ for 2-4 hours, and cool to obtain strip-shaped, zinc-modified ZSM-5 molecular sieves with a diameter of 1.2-1.5 mm. In the formulation, the weight ratio of modified ZSM-5 molecular sieve powder, pseudoboehmite powder (calculated as Al2O3), guar gum powder, and water (total water content in acetic acid aqueous solution and water added later) is 100:(30-40):(3-5):(100-130).

[0018] (4) The strip-shaped, zinc-modified ZSM-5 molecular sieve is added to the spray drum, and yttrium nitrate aqueous solution of 95-100% of its saturated water absorption volume and the required concentration is applied by atomization spraying for 10-20 min. After sealing and homogenization treatment for 10-15 h, the material is discharged and dried, calcined at 540-560℃ for 2-4 h, cooled, and shaped to obtain an isomerized catalyst of Φ(1.2-1.5)×(3-6) mm.

[0019] In the process for preparing n-pentene from n-pentane dehydrogenation of this invention, the n-pentane dehydrogenation unit is further equipped with a heat exchanger T1. A mixture of n-pentane liquid flow and / or n-pentane-rich gas flow, hydrogen flow and / or hydrogen-rich gas flow, and carbon dioxide flow X1, along with a post-reaction gas flow X3 from the dehydrogenation reactor group, undergoes countercurrent heat exchange in heat exchanger T1 to obtain a pre-reaction gas flow X2 and a cooled gas flow X4. The pre-reaction gas flow X2 enters the dehydrogenation reactor series, and the cooled gas flow X4 goes to the n-pentene separation unit. The impurities in the n-pentane-rich gas flow and the hydrogen-rich gas flow that are detrimental to the n-pentane dehydrogenation process and the performance of the dehydrogenation catalyst should be controlled below a certain content.

[0020] The isomerization reaction and circulation separation device is also equipped with a heat exchanger T2; the 2-pentene gas stream separated before the top condenser of the 2-pentene deweighting tower exchanges heat with the outlet gas stream of each reactor VI in the heat exchanger T2 before going to the isomerization reactor group.

[0021] The number of trays in the pentene separation tower and the 2-pentene deweighting tower are 110-130 and 25-30, respectively, and the top pressure of both towers is 0.1-0.2 MPa. The top reflux ratio of the pentene separation tower is 10-20, and the top reflux ratio of the 2-pentene deweighting tower is 0.5-1.

[0022] The 2-pentene-rich liquid stream at the bottom of the pentene separation tower should have its 1-pentene content controlled to below 1 wt%, preferably below 0.5 wt%, to avoid affecting or reducing the 1-pentene yield in isomerization reactor VI.

[0023] The heavy hydrocarbon stream P-4 separated from the bottom of the 2-pentene de-heavy column has an initial boiling point above 100°C at atmospheric pressure. Its main components should be the dimerization and trimerization products of 2-pentene in isomerization reactor VI. By controlling the operating conditions of isomerization reactor VI, the yield of heavy hydrocarbon stream P-4 can be less than 5% of the 1-pentene yield. An appropriate amount of polymerization inhibitor, such as 0.01-0.03 wt% hydroquinone, can be added to the bottom liquid of the 2-pentene de-heavy column.

[0024] In the process for preparing n-pentene from n-pentane of this invention, the setup and operation of the n-pentane dehydrogenation unit are basically the same as those in CN120887769A, with the main difference being that the inlet gas flow temperature of the dehydrogenation reactor is reduced, including any of the following modes:

[0025] Mode B: Set up a series of 2-3 parallel dehydrogenation reactors. Each series includes two dehydrogenation reactors, I and II, connected in series sequentially, with reactor I being the first reactor. The volume ratio of the dehydrogenation catalyst in reactors I and II is 1:(0.8-1.2). An alkane heater is installed before the gas flow inlet of each dehydrogenation reactor. During stable operation, the inlet gas flow conditions for reactor I include: temperature 490-510℃, pressure ≤0.2MPa, and n-pentane mass hourly space velocity 1.5-2.5h⁻¹. -1 The molar ratio of n-pentane to H2 is 1:(0.5-1), and the molar content of CO2 is 1.5-2.5%. The inlet gas flow conditions for reactor II include: temperature 480-500℃, and molar content of CO2 1.5-2%. The maximum temperature of the dehydrogenation catalyst bed in each reactor is controlled below 520℃. The outlet gas flow temperature of reactor II or the outlet section temperature of the dehydrogenation catalyst bed is controlled above 500℃.

[0026] Mode C: A series of 2-3 parallel dehydrogenation reactors are set up. Each series includes three dehydrogenation reactors (III, IV, and V) connected in series sequentially, with reactor III being the first reactor. The volume ratio of the dehydrogenation catalyst in reactors III, IV, and V is 1:(0.8-1):(0.8-1). An alkane heater is installed before the gas flow inlet of each dehydrogenation reactor. During stable operation, the inlet gas flow conditions for reactor III include: temperature 520-540℃, pressure ≤0.2MPa, and n-pentane mass hourly space velocity (MHV) 1.5-2.5h⁻¹. -1 The molar ratio of n-pentane to H2 and water is 1:(0.3-0.8):(2-3); the inlet gas flow temperature of reactors IV and V is controlled at 520-530℃; the outlet gas flow temperature of reactor V or the outlet section temperature of the dehydrogenation catalyst bed is controlled at above 490℃.

[0027] In modes B and C, when the reaction performance of a certain dehydrogenation reactor series drops to a specified range, the dehydrogenation operation is switched off and the catalyst is regenerated in situ. After the regeneration is completed, the dehydrogenation operation is switched back on.

[0028] The alkane dehydrogenation catalysts packed in reactors I-V can be spherical or strip-shaped. Pt is uniformly distributed in the catalyst particles (i.e., uniform cross-sectional color) and is highly dispersed at the beginning of each operating cycle. The spherical catalyst size can be Φ2.5-3.0 mm, and the strip catalyst size can be Φ(1.2-2)×(2-10) mm. Catalyst H or catalyst J prepared and used in CN120887769A can be used. Catalyst H uses a high-temperature zirconium-aluminum composite oxide containing 0.8-1.5 wt% ZrO2 as a support, and the supported components are Pt 0.2-0.4 wt%, Sn 0.5-1 wt%, K 0.65-1 wt% (based on elemental composition). Preferably, catalyst HA is used, which contains 0.35 wt% Pt, 0.7 wt% Sn, and 0.83 wt% K, and the support is a high-temperature zirconium-aluminum composite oxide containing 1.1 wt% ZrO2. Catalyst J uses MgO... The support consists of 1-2 wt% magnesium-modified alumina, with the supported components being Pt 0.2-0.4 wt%, Sn 0.5-1 wt%, and K 0.65-1 wt% by element. The catalyst JB containing Pt 0.3 wt%, Sn 0.7 wt%, and K 0.83 wt% is preferred, and the support is magnesium-modified alumina containing 1.5 wt% MgO. It was found that catalyst H, especially catalyst HA, and catalyst J, especially catalyst JB, under fixed-bed adiabatic reactor and hydrogen-containing conditions, such as the setup and operation conditions of mode B, all exhibit good reaction performance for the dehydrogenation of n-pentane to n-pentene and the matching hydrogenation of CO2 to methane, with CO2 conversion easily exceeding 90%. Under the setup and operation conditions of mode C (without CO2), they also exhibit good reaction performance for the dehydrogenation of n-pentane to n-pentene, and their catalytic activity is relatively stable with slow deactivation. Both catalysts can achieve long-term operation with a single-pass conversion of n-pentane of over 25%, a selectivity of over 90% for n-pentene, and a catalyst regeneration cycle of over 300 hours or even over 450 hours. In CN120887769A, when using Mode A for the dehydrogenation process of n-pentane, even with inlet gas flow temperature and gas composition conditions similar to Mode B, it is difficult to achieve long-term operating results with a single-pass conversion rate of over 25% for n-pentane, a selectivity of over 90% for n-pentene, and a catalyst regeneration cycle of over 220 hours. This is because the reaction capacity of a single reactor is insufficient under low temperature conditions, the controllability of operating conditions is too narrow, and increasing the CO2 content will lead to an increase in side reactions in the dehydrogenation of isopentane and a decrease in the selectivity of n-pentene to below 90%.

[0029] In the process of preparing n-pentene from n-pentane in this invention, the cooling separation system of the n-pentene separation device first further cools and pressurizes the n-pentane dehydrogenated gas stream X4 after it has been cooled by heat exchanger T1, such as cooling it to (-20 to -10)℃ and pressurizing it to 0.5-1MPa, separating water and obtaining a gas stream G0 mainly containing H2, CH4 and a small amount of CO2, ethane and ethylene, and a hydrocarbon-containing liquid stream L0; the hydrocarbon-containing liquid stream L0 enters the light hydrocarbon removal tower (50-60 trays), and after the condenser at the top of the tower, a light hydrocarbon liquid stream P-1 containing C4 components and a small amount of isopentane is separated, as well as a small amount of light hydrocarbon gas stream G1 containing C4 and below components, and a hydrocarbon-containing liquid stream L is separated at the bottom of the tower. The pressure swing adsorption (PSA) separation system includes a PSA-H2 unit for separating hydrogen from gas flow G0, and an alkane-alkene separation unit PSA-CH. The PSA-H2 unit employs conventional PSA hydrogen extraction adsorbents, equipment configuration, and separation operation processes, while the PSA-CH unit employs known alkane-alkene separation adsorbents, equipment configuration, and separation operation processes. Gas flow G0 is separated in the PSA-H2 unit to obtain a hydrogen gas flow G-H2 with a molar purity of 98% or even 99% and almost no unsaturated hydrocarbons. Part of this hydrogen is recycled to the n-pentane dehydrogenation unit, and the remainder is sold. A low-hydrogen gas flow G2 containing CH4, ethane, ethylene, and CO2 is also obtained and used as fuel gas in a heater. The alkane-olefin separation unit PSA-CH separates the vaporized hydrocarbon-containing liquid stream L into three streams: n-pentane stream M1, n-pentene stream P-2, and heteroolefin stream P-3, each with the desired purity. Stream P-2 can be controlled to have a n-pentene molar content of over 99%. Small amounts of dienes and other heteroolefins are primarily distributed into heteroolefin stream P-3 (in very small quantities), which can be used as fuel in a heating furnace or collected for sale after liquefaction. The n-pentane molar purity of stream M1 is easily controlled to over 98%. The light hydrocarbon liquid stream P-1, separated after the light hydrocarbon stripping tower top condenser, is collected and sold. The light hydrocarbon stream G1 is also used as fuel in alkane and olefin heating furnaces.

[0030] The process of the cooling and separation system may include: cooling gas flow X4 to 20-60℃ (e.g., through a heat exchanger with circulating water), pressurizing to 0.5-1MPa, cooling to 20-60℃, then successively cooling to 1-10℃ and (-20 to -10)℃ while performing water separation and de-icing operations; after cooling to (-20 to -10)℃, gas-liquid separation is performed to obtain gas flow G0 and hydrocarbon-containing liquid flow L0; the two-step cooling of 1-10℃ and (-20 to -10)℃ uses a cooling heat exchanger with water separation function, and condensate is separated when cooling to 1-10℃ (in mode B, this is the water produced by the hydromethane reaction of CO2), and (-20 to -10)℃... During the cooling process at 0℃, multiple cooling heat exchangers can be used in turn to perform de-icing operations at regular intervals (short-term shutdown of the circulation of the cooling medium) to melt and remove the ice on the heat exchange surface, thereby maintaining long-term heat exchange capacity and deeply removing moisture from the material flow. In mode C, since a large amount of water vapor is introduced into the inlet gas flow of reactor III, the gas flow X4 is continuously separated into condensate by the oil-water separator after two cooling processes at 20-60℃ before being pressurized to 0.5-1MPa and afterward. The gas-liquid mixture flow cooled to (-20 to -10)℃ is then separated by the gas-liquid separator into gas flow G0, which mainly contains H2, CH4 and a small amount of CO2, ethane, and ethylene, and liquid flow L0 containing hydrocarbons.

[0031] The adsorbents used in each adsorption tower of the PSA-CH unit should have an adsorption selectivity of 4.5 or higher, preferably 6 or higher, for n-pentene / n-pentane in the vaporized stream of hydrocarbon-containing liquid flow L under their operating conditions (e.g., 40-100℃, absolute pressure 30-1000kPa), meaning they are more likely to adsorb n-pentene.

[0032] In the process of preparing n-pentene from n-pentane of this invention, the reduction method performed after loading new catalyst H or catalyst J into the n-pentane dehydrogenation reactor and before the dehydrogenation operation, and the in-situ regeneration method performed when the reaction performance indicators drop to a specified range at the end of the dehydrogenation operation, are basically the same as those used in CN120887769A, which can achieve better regeneration effect and maintain stable dehydrogenation reaction performance; the in-situ regeneration of the alkane dehydrogenation catalysts H and J both adopt the more effective chlorine regeneration method.

[0033] In a reactor VI loaded with 2-pentene isomerization catalyst, after the reaction performance indicators drop to the specified range, the isomerization operation is simultaneously shut off from the upstream olefin heater and purged with nitrogen. The in-situ char regeneration process is as follows: An oxygen-containing stream is obtained by mixing a nitrogen stream (N2 molar content above 99%) with an air stream, with an O2 molar content of 0.5-1.2%, and a space velocity of 500-2000 h⁻¹. -1The inlet temperature is 450-490℃ (preheated by an olefin heater or other heater), the pressure is 0-0.1MPa, the initial bed temperature is 360-390℃, and the maximum bed temperature is controlled not to exceed 530℃. The O2 molar content difference between the outlet and inlet gas streams is reduced to below 0.1%. Then, the O2 molar content in the oxygen-containing gas stream is increased to 5-10%, and the maximum bed temperature is controlled not to exceed 550℃. Treatment lasts for 4-8 hours. Afterwards, purging is performed successively with ordinary nitrogen gas at 370-375℃ and high-purity nitrogen gas at a space velocity of 500-2000 h⁻¹. -1 Once the O2 molar content in the outlet gas stream is reduced to below 0.1% and the maximum bed temperature is reduced to below 390℃, the isomerization operation is then initiated.

[0034] The beneficial effects of the present invention include:

[0035] 1. The process for preparing n-pentene by dehydrogenation of n-pentane, when the n-pentane dehydrogenation unit is set and operated under the conditions of modes B and C, can achieve a single-pass conversion rate of over 25% for n-pentane, a selectivity of over 90% for n-pentene, and a catalyst regeneration cycle of over 300 hours or even over 500 hours for long-term operation. The n-pentene separation unit, isomerization reaction, and circulation separation unit are set up with high efficiency and reasonableness. During the isomerization reaction, the single-pass yield of 2-pentene to 1-pentene exceeds 20%, and the selectivity of 1-pentene exceeds 95%, which can finally produce n-pentene with a molar purity of over 98% or even over 99%, and hydrogen. The regeneration cycle of the dehydrogenation reactor series and the 2-pentene isomerization catalyst reactor VI both exceed 300 hours.

[0036] 2. This invention provides an effective process for preparing n-pentene from n-pentane, expanding the production sources of n-pentene and offering a potentially better option for the application of n-pentane; n-pentene has a high market price. As a complete process, this process is superior to existing technologies in terms of overall technology and economy, and has certain application prospects.

[0037] Comprehensive analysis suggests that the main principles underlying the process for preparing n-pentene from n-pentane via dehydrogenation in this invention, which achieves the aforementioned beneficial effects, include:

[0038] 3. The feed to isomerization reactor VI is 2-pentene with a molar purity of 98% or even 99%, achieving a single-pass yield of over 20% for the conversion to 1-pentene, with a selectivity of over 95%. The primary reactions are hydrogen bond transfer and oligomerization, with almost no detectable cracking or skeletal isomerization products. Furthermore, the regeneration cycle exceeds 300 hours, which is unexpected and not achieved in existing technologies such as CN104557407A. The heavy hydrocarbon stream P-4 separated from the bottom of the 2-pentene de-heavy column has an initial boiling point above 100℃ at atmospheric pressure. Its main components are the dimerization and trimerization products of 2-pentene in isomerization reactor VI, with a yield less than 5% of the 1-pentene yield. This effect indicates that the prepared 2-pentene isomerization catalyst has a relative crystallinity of over 95%, a silicon-to-aluminum ratio (SiO2:Al2O3, molar ratio) of 300-350, and a surface area of ​​500-600 m². 2 The hydrogen-form ZSM-5 molecular sieve with an average grain size of 0.3-0.5 μm and a g / g, after being modified with zinc and yttrium, exhibits both moderate acidity (L-acidity) and weak alkalinity (L-baseity). The acidic sites of the molecular sieve are successively replaced by zinc and yttrium. In step (2), the substitution of zinc gives the molecular sieve moderate acidity, and in step (4), the substitution of yttrium gives the molecular sieve weak alkalinity, thus eliminating strong acid sites. In step (4), after the molding strip is immersed in yttrium nitrate aqueous solution and calcined, the strong acid sites on the inner and outer surfaces of the active alumina grains are also replaced by yttrium and transformed into weak alkaline sites. During the isomerization reaction, the catalyst achieved a 2-pentene conversion rate and 1-pentene selectivity that are much higher than those of the prior art, indicating that the modification effect of zinc and yttrium is significant.

[0039] 4. The n-pentene separation device uses the aforementioned alkene separation unit PSA-CH, which separates n-pentene gas stream P-2 with a purity of over 99% and essentially free of pentane and diene. This simplifies the equipment configuration of the isomerization reaction and the circulating separation device, and significantly improves the efficiency and operational stability of separating and generating 1-pentene. Therefore, the process of this invention has a combined effect. Attached Figure Description

[0040] Appendix Figure 1 The following is a schematic diagram of the process flow for preparing n-pentene from n-pentane in Example 1; wherein, the n-pentane dehydrogenation unit adopts Mode B, and only the outline of the material inflow and outflow relationship is shown for each PSA unit. Detailed Implementation

[0041] The technical solution of the present invention will be specifically described and explained below with reference to the embodiments.

[0042] Example 1

[0043] Based on the performance evaluation of the dehydrogenation catalyst HA (containing 0.35wt% Pt, 0.7wt% Sn, and 0.83wt% K, with a spherical catalyst HA supported by a high-temperature zirconium-aluminum composite oxide containing 1.1wt% ZrO2) in Catalyst Evaluation Example 1, the performance of the dehydrogenation catalyst JB (containing 0.3wt% Pt, 0.7wt% Sn, and 0.83wt% K, with a magnesium-modified alumina supported by 1.5wt% MgO) in Catalyst Evaluation Example 2, and the performance of the 2-pentene isomerization catalyst prepared in Example 4 in Catalyst Evaluation Example 3, the process for preparing n-pentene from n-pentane in this example was designed, as detailed in the appendix. Figure 1 The configuration and operating effects shown are based on simulations using the Aspen Plus software.

[0044] The process for preparing n-pentene from n-pentane in Example 1 includes a n-pentane dehydrogenation unit, a n-pentene separation unit, an isomerization reaction and a recycling separation unit;

[0045] The n-pentane dehydrogenation unit comprises a fixed-bed adiabatic reactor packed with a Pt-Sn-K / alumina type alkane dehydrogenation catalyst and having an axial reaction structure. The alkane dehydrogenation catalyst exhibits hydromethanation performance under CO2-containing dehydrogenation reaction conditions. The setup and operation of the n-pentane dehydrogenation unit include any of the following modes:

[0046] The setup and operation of the n-pentane dehydrogenation unit includes any of the following modes:

[0047] Mode B: Set up a series of 2-3 parallel dehydrogenation reactors. Each series includes two dehydrogenation reactors, I and II, connected in series sequentially, with reactor I being the first reactor. The volume ratio of the dehydrogenation catalyst in reactors I and II is 1:(0.8-1.2). An alkane heater is installed before the gas flow inlet of each dehydrogenation reactor. During stable operation, the inlet gas flow conditions for reactor I include: temperature 490-510℃, pressure ≤0.2MPa, and n-pentane mass hourly space velocity 1.5-2.5h⁻¹. -1 The molar ratio of n-pentane to H2 is 1:(0.5-1), and the molar content of CO2 is 1.5-2.5%. The inlet gas flow conditions for reactor II include: temperature 480-500℃, and molar content of CO2 1.5-2%. The maximum temperature of the dehydrogenation catalyst bed in each reactor is controlled below 520℃. The outlet gas flow temperature of reactor II or the outlet section temperature of the dehydrogenation catalyst bed is controlled above 500℃.

[0048] Mode C: A series of 2-3 parallel dehydrogenation reactors are set up. Each series includes three dehydrogenation reactors (III, IV, and V) connected in series sequentially, with reactor III being the first reactor. The volume ratio of the dehydrogenation catalyst in reactors III, IV, and V is 1:(0.8-1):(0.8-1). An alkane heater is installed before the gas flow inlet of each dehydrogenation reactor. During stable operation, the inlet gas flow conditions for reactor III include: temperature 520-540℃, pressure ≤0.2MPa, and n-pentane mass hourly space velocity (MHV) 1.5-2.5h⁻¹. -1 The molar ratio of n-pentane to H2 and water is 1:(0.3-0.8):(2-3); the inlet gas flow temperature of reactors IV and V is controlled at 520-530℃; the outlet gas flow temperature of reactor V or the outlet section temperature of the dehydrogenation catalyst bed is controlled at above 490℃.

[0049] The n-pentane dehydrogenation unit is also equipped with a heat exchanger T1; the n-pentane liquid flow and / or the mixed flow of n-pentane-rich gas, hydrogen gas and carbon dioxide gas X1, and the reaction gas X3 from the dehydrogenation reactor group, exchange heat in countercurrent with the heat exchanger T1 to obtain the pre-reaction gas X2 and the cooled gas X4. The pre-reaction gas X2 enters the dehydrogenation reactor series, and the cooled gas X4 goes to the n-pentene separation unit.

[0050] The n-pentene separation unit includes a cooling separation system and a pressure swing adsorption separation system. From the dehydrogenated gas stream from the n-pentane dehydrogenation unit after cooling, hydrogen gas stream G-H2 is separated and reused in the n-pentane dehydrogenation unit and sold externally. Unreacted n-pentane gas stream M1 is separated and reused in the n-pentane dehydrogenation unit. n-pentene gas stream P-2 is separated and sent to the isomerization reaction and recycling separation unit for processing.

[0051] The cooling separation system of the n-pentene separation unit first further cools the n-pentane dehydrogenated gas stream X4, which has been cooled by heat exchanger T1, by passing it through a cooling heat exchanger and pressurizing it by a compressor. The stream is cooled to (-20 to -10)℃ and pressurized to 0.5-1 MPa, separating water and obtaining a stream G0 mainly containing H2, CH4, and small amounts of CO2, ethane, and ethylene, and a hydrocarbon-containing liquid stream L0. The hydrocarbon-containing liquid stream L0 enters a light-light removal tower (50-60 trays), and after condensation at the top of the tower, a C4-containing component is separated. The system consists of a light hydrocarbon liquid stream P-1 containing a small amount of isopentane and a light hydrocarbon gas stream G1 containing C4 or less components. The bottom of the tower separates a hydrocarbon-containing liquid stream L. The pressure swing adsorption separation system includes a PSA-H2 unit for separating hydrogen from gas stream G0 and a PSA-CH unit for separating alkanes and alkenes. The PSA-H2 unit uses conventional pressure swing adsorption hydrogen extraction adsorbents, equipment configurations, and separation operation processes, while the PSA-CH unit uses known alkane and alkene separation adsorbents, equipment configurations, and separation operation processes. Gas flow G0 is separated in the PSA-H2 unit to obtain a hydrogen gas flow G-H2 with a molar purity of over 98% or even over 99% and almost no unsaturated hydrocarbons; a low-hydrogen gas flow G2 containing CH4, ethane, propane, propylene, and CO2 is also obtained and used as fuel gas in the heater. The alkane-olefin separation unit PSA-CH separates the hydrocarbon-containing liquid flow L after gasification into three parts: n-pentane gas flow M1, n-pentene gas flow P-2, and heteroene gas flow P-3, all of the required purity. The molar content of n-pentene in gas flow P-2 is controlled to be over 99%, and a small amount of heteroenes such as olefins with more than 9 Cs are basically sent to heteroene gas flow P-3, liquefied, collected, and sold. The molar purity of n-pentane in gas flow M1 is controlled to be over 98%. The light hydrocarbon liquid flow P-1 separated after the light tower top condenser is collected and sold, and the separated light hydrocarbon gas flow G1 is also used as fuel gas in the alkane heater and olefin heater. The operating pressure of the light tower, as well as the maximum operating pressure of the PSA-H2 and PSA-CH units, are slightly lower than the outlet gas flow pressure of the compressor.

[0052] The process of the cooling and separation system includes: cooling gas flow X4 to 20-60℃ (e.g., using a heat exchanger with circulating water for cooling), pressurizing to 0.5-1MPa, cooling to 20-60℃, then successively cooling to 1-10℃ and (-20 to -10)℃ while performing water separation and de-icing operations; after cooling to (-20 to -10)℃, gas-liquid separation yields gas flow G0 and hydrocarbon-containing liquid flow L0; the 1-10℃ and (-20 to -10)℃ two-step cooling uses a cooling heat exchanger with water separation function, separating condensate (in mode B, this is the water produced by the hydromethanization reaction of CO2) when cooling to 1-10℃, and (-20 to -10)℃... During the ℃ cooling process, multiple cooling heat exchangers can be used in turn to perform de-icing operations (short-term shutdown of the circulation of cooling medium) to melt and remove the ice on the heat exchange surface, thereby maintaining long-term heat exchange capacity and deeply removing moisture from the material flow. In mode C, since a large amount of water vapor is introduced into the inlet gas flow of reactor III, the gas flow X4 is continuously separated into condensate by the oil-water separator after two cooling processes at 20-60℃ before being pressurized to 0.5-1MPa and afterward. The gas-liquid mixture flow cooled to (-20 to -10)℃ is then separated by the gas-liquid separator into gas flow G0, which mainly contains H2, CH4 and a small amount of CO2, ethane, and ethylene, and liquid flow L0 containing hydrocarbons.

[0053] The PSA-CH unit employs a three-tower adsorption process. Under operating conditions of 40-100℃ and 30-1000kPa absolute pressure, the alkane adsorbents used in each adsorption tower exhibit an adsorption selectivity of approximately 6.3 for n-pentene / n-pentane in the vaporized stream containing hydrocarbon liquid flow L. The operating sequence of each adsorption tower includes high-pressure adsorption, displacement, vacuum desorption, and pressurization processes, and is cyclically operated. The opening and closing sequence of each exhaust valve, i.e., the cutting and reuse sequence of the exhaust gas flow, is adjusted according to the composition and purity requirements of each separated product.

[0054] The isomerization reaction and recycling separation unit includes a pentene separation tower, a 2-pentene de-heavy tower, and a 2-pentene isomerization reactor group. A stream of n-pentene gas, P-2, from the n-pentene separation unit is mixed with a cooled isomerized gas stream from the 2-pentene isomerization reactor group and then fed into the pentene separation tower. A 1-pentene liquid product is separated after the top condenser of the pentene separation tower, and the 2-pentene-rich liquid stream obtained at the bottom of the tower enters the 2-pentene de-heavy tower. A heavy hydrocarbon liquid stream, P-4, is separated from the bottom of the 2-pentene de-heavy tower as a byproduct. The product is sold externally. A 2-pentene gas stream is separated before the condenser at the top of the column and sent to the isomerization reactor group. The isomerization reactor group consists of 2-3 fixed-bed adiabatic reactors VI, each packed with 2-pentene isomerization catalyst and having an axial reaction structure. Each reactor VI has an olefin heater installed before its gas stream inlet. During stable operation, the reaction conditions for each reactor VI include: temperature 375-390℃, pressure 0.05-0.15MPa (gauge pressure), and 2-pentene mass hourly space velocity (MHV) 1-2h⁻¹. -1 ;

[0055] The isomerization reaction and circulation separation device is also equipped with a heat exchanger T2; the 2-pentene gas stream separated before the top condenser of the 2-pentene deweighting tower exchanges heat with the outlet gas stream of each reactor VI in the heat exchanger T2 before going to the isomerization reactor group.

[0056] The number of trays in the pentene separation tower and the 2-pentene deweighting tower are 110-130 and 25-30, respectively, and the top pressure of both towers is 0.1-0.2 MPa. The top reflux ratio of the pentene separation tower is 10-20, and the top reflux ratio of the 2-pentene deweighting tower is 0.5-1. The 2-pentene-rich liquid flow at the bottom of the pentene separation tower should have the 1-pentene content controlled to below 1 wt%, preferably below 0.5 wt%. In the bottom liquid of the 2-pentene deweighting tower, 0.01-0.03 wt% of hydroquinone is added and maintained to inhibit polymerization.

[0057] In this embodiment, the process for preparing n-pentene from n-pentane involves a reduction operation after the n-pentane dehydrogenation reactor is filled with new catalyst HA or catalyst JB and before the dehydrogenation operation begins. The final step in the regeneration process is also a reduction operation. The reduction operation conditions are: a hydrogen gas flow with an H2 molar content higher than 98% and an olefin molar content lower than 0.1%, and a space velocity of 500-1000 h⁻¹. -1 The pressure is 0-0.2 MPa, the reduction time at temperatures above 480℃ is 3-6 hours, and the maximum bed temperature does not exceed 520℃.

[0058] In a series of dehydrogenation reactors in a n-pentane dehydrogenation unit, after the overall reaction performance indicators drop to a specified range during operation, the dehydrogenation operation is stopped and the reactor is purged with nitrogen or steam, followed by an in-situ regeneration process including carbonization, oxychlorination, and reduction. During the carbonization operation, an oxygen-containing air stream with an O2 molar content of 0.5-0.8% is prepared by mixing a general nitrogen stream (N2 molar content above 99%) with an air stream, at a space velocity of 500-2000 h⁻¹. -1 The inlet temperature is 450-490℃ (preheated by an alkane heater or refrigerant heater), the pressure is 0-0.1MPa, the initial bed temperature is 450-500℃, and the maximum bed temperature does not exceed 520℃. The process continues for 2-4 hours until the difference in O2 molar content between the outlet and inlet gas streams decreases to below 0.1%. For oxychlorination, an oxygen-containing gas stream with an O2 molar content of 4-15%, obtained by mixing nitrogen gas with air, is used, with a space velocity of 500-2000 h⁻¹. -1The pressure is 0.02-0.2 MPa, and the initial bed temperature is 480-520℃. First, water or steam, as well as chlorides such as dichloroethane, trichloroethane, carbon tetrachloride, or tetrachloroethylene, are added to the oxygen-containing stream. Alternatively, a 40-50% aqueous solution of chloroacetic acid (containing water and organic chlorine) can be directly added before the alkane heater or heater. The inlet temperature of the gas flow is controlled at 500-520℃, and the amount of chloride added is 0.4-0.6 wt% of the catalyst in the reactor, calculated as chlorine. The water-chlorine molar ratio is controlled at 5-8, and the addition time is 3-6 hours. Then, the inlet temperature of the oxygen-containing stream (without adding water and chloride) is controlled at 520-540℃, and the treatment is carried out for 4-6 hours, with the reactor outlet gas flow temperature above 515℃. After the oxychlorination operation was completed, the reactor was purged successively with a stream of ordinary nitrogen at a temperature of 480-500℃ and a stream of high-purity nitrogen (N2 molar content above 99.9%), with a space velocity of 500-2000 h⁻¹. -1 Once the O2 molar content in the outlet gas stream is reduced to below 0.2%, catalyst HA or catalyst JB is reduced again under the aforementioned reduction operating conditions to complete the regeneration process and initiate the dehydrogenation operation. After in-situ regeneration of catalyst HA or catalyst JB including the aforementioned oxychlorination operation, the coke deposits are largely burned off, the Pt particles are redispersed, and the performance of n-pentane dehydrogenation and CO2 hydromethanization is basically restored. Catalyst HA can basically recover to the performance level of the first operating cycle, while the overall performance of catalyst JB is slightly lower than the performance level of the first operating cycle without chlorine.

[0059] In a reactor VI loaded with 2-pentene isomerization catalyst, after the reaction performance indicators drop to the specified range, the isomerization operation is simultaneously shut off from the upstream olefin heater and purged with nitrogen. The in-situ char regeneration process is as follows: an oxygen-containing air stream with an O2 molar content of 0.5-1.2% is obtained by mixing a nitrogen stream (N2 molar content above 99%) with an air stream, and a space velocity of 500-2000 h⁻¹. -1 The inlet temperature is 450-490℃ (preheated by an olefin heater or other heater), the pressure is 0-0.1MPa, the initial bed temperature is 360-390℃, and the maximum bed temperature is controlled not to exceed 530℃. The O2 molar content difference between the outlet and inlet gas streams is reduced to below 0.1%. Then, the O2 molar content in the oxygen-containing gas stream is increased to 5-10%, and the maximum bed temperature is controlled not to exceed 550℃. Treatment lasts for 4-8 hours. Afterwards, purging is performed successively with nitrogen gas at 370-375℃ and high-purity nitrogen gas at a space velocity of 500-2000 h⁻¹. -1 Once the O2 molar content in the outlet gas stream is reduced to below 0.1% and the maximum bed temperature is reduced to below 390℃, the isomerization operation is then initiated.

[0060] This embodiment 1 describes a process for preparing n-pentene from n-pentane via dehydrogenation. Under the settings and operating conditions of modes B and C, the n-pentane dehydrogenation unit achieves a single-pass conversion rate of over 25% for n-pentane, a selectivity of over 90% for n-pentene, and a catalyst regeneration cycle of over 300 hours, even exceeding 500 hours, for long-term operation. The n-pentene separation unit, isomerization reaction, and circulating separation unit are set up efficiently and rationally. During the isomerization reaction, the single-pass yield of 1-pentene from 2-pentene exceeds 20%, and the selectivity of 1-pentene exceeds 95%. The regeneration cycle of isomerization reactor VI exceeds 300 hours, ultimately yielding n-pentene with a molar purity of over 98%, even exceeding 99%, and hydrogen. The service life of dehydrogenation catalysts HA, JB, and the 2-pentene isomerization catalyst should all be over 3 years. As a complete process, this process is superior to existing technologies in terms of overall technology and economy, and has certain application prospects.

[0061] Example 2

[0062] This Example 2 directly uses the Pt-Sn-K / alumina type alkane dehydrogenation catalyst HA prepared in Example 3 of CN120887769A (see paragraphs 0069-0074 of its specification): a high-temperature zirconium-aluminum composite oxide containing 1.1wt% ZrO2 as a support, with the supported components being Pt 0.35wt%, Sn 0.7wt%, and K 0.83wt% (calculated content in the catalyst), and spherical with an average outer diameter of Φ2.8mm.

[0063] Example 3

[0064] This Example 3 directly uses the Pt-Sn-K / alumina type alkane dehydrogenation catalyst JB prepared in Example 4 of CN120887769A (see paragraphs 0075-0080 of its specification): Magnesium-modified alumina containing 1.5wt% MgO is used as the support, and the supported components are Pt 0.3wt%, Sn 0.7wt%, and K 0.83wt% (calculated content in the catalyst), and are spherical with an average outer diameter of Φ2.8mm.

[0065] Example 4

[0066] In Example 4, a 2-pentene isomerization catalyst was prepared, and the operation steps are as follows:

[0067] (1) A product with a relative crystallinity of over 97%, a silicon-to-aluminum ratio (SiO2:Al2O3, mol ratio) of 320, and a surface area of ​​537 m² 23000g of Na-type ZSM-5 molecular sieve powder with an average grain size of 0.41μm was placed in a jacketed 15L stirred tank. 10L of 1.7mol / L NH4NO3 aqueous solution was added each time to slurry the powder, and the temperature was controlled at 80-90℃ for 2h. This process was repeated 5 times. Between two exchanges, the slurry was discharged and filtered, and the filter cake was returned to the stirred tank. After the exchange was completed, the slurry was discharged and filtered, the filter cake was washed with water, and the washed filter cake was placed on a tray and dried in an oven at 120℃ for 6h. The dried material was passed through an 80-mesh sieve and calcined in a muffle furnace at 480℃ for 3h. When the temperature was lowered to 300℃, it was immediately transferred to a sealed aluminum container for cooling to obtain HZSM-5 molecular sieve powder. The pore volume was determined to be 0.20mL / g by low-temperature nitrogen adsorption method.

[0068] (2) Take 2000g of HZSM-5 molecular sieve powder obtained in step (1) and place it in a mixer. During stirring, apply 400mL (100% of its inner pore volume of 0.20mL / g*2000g) of zinc acetate aqueous solution with a concentration of 1.30mol / L by atomizing and spraying. After homogenization treatment for 12h, discharge the material, pack it into a tray and dry it in an oven at 120℃ for 6h. After drying, calcine the powder in a muffle furnace at 530℃ for 3h, cool it, and pass it through a 200-mesh sieve to obtain 2040g of modified ZSM-5 molecular sieve powder containing 2.1wt% zinc (calculated as ZnO).

[0069] (3) Add 1087g of boehmite powder (containing 65.7wt% Al2O3) to a 20L kneader (anhydrous). Start the kneader and uniformly apply 130g of 22wt% acetic acid aqueous solution by atomization spraying. The weight ratio of boehmite powder to acetic acid aqueous solution is 100:12. After adding the acetic acid aqueous solution, stop the kneader for 45 minutes and cover and seal for 15 hours. Start the kneader again and add 2040g of all the ZSM-5 molecular sieve powder and 81.6g of guar gum powder prepared in step (2). After 60 minutes, mix the powder evenly and add water by atomization spraying. 2245g of material was kneaded for 45 minutes and then extruded into strips. The extruded strips were placed in trays and dried in an oven at 120℃ for 4 hours. The dried strips were then calcined in a muffle furnace at 530℃ for 3 hours. When the temperature was lowered to 300℃, the strips were immediately transferred to a sealed aluminum container for cooling. This yielded strip-shaped, zinc-modified ZSM-5 molecular sieves with a diameter of 1.3-1.4mm. The pore volume was determined to be 0.32mL / g by the water absorption method. In the formulation, the weight ratio of modified ZSM-5 molecular sieve powder, pseudoboehmite powder (calculated as Al2O3), guar gum powder, and water (the total amount of water in the acetic acid aqueous solution and the water added later) was 100:35:4:115.

[0070] (4) Take 960g of the strip-shaped, zinc-modified ZSM-5 molecular sieve from step (3) and add it to a small, sealable drum (12L volume). Apply 307mL of yttrium nitrate aqueous solution with a concentration of 1.15mol / L (accounting for 100% of the saturated water absorption volume of the strip-shaped, zinc-modified ZSM-5 molecular sieve) by atomization spraying for 15min. After sealing and homogenization treatment for 8h, discharge and dry the material. Calcine at 550℃ for 3h. When the temperature drops to 300℃, immediately transfer it to a sealed aluminum bucket for cooling and shaping to obtain an isomerized catalyst with a size of Φ(1.3-1.4)×(3-6)mm, containing 2wt% zinc (calculated as ZnO) and 4wt% yttrium (calculated as Y2O3).

[0071] Example 5

[0072] The 2-pentene isomerization catalyst of this example is prepared by basically following steps (2)-(4) of Example 4, with the main difference being:

[0073] In step (2), a zinc acetate aqueous solution with a concentration of 1.57 mol / L was used instead. After drying, the calcination temperature of the powder in the muffle furnace was changed to 550℃ to obtain modified ZSM-5 molecular sieve powder containing 2.5 wt% zinc (calculated as ZnO).

[0074] In step (4), the amount of strip-shaped, zinc-modified ZSM-5 molecular sieve was changed to 970g, and yttrium nitrate aqueous solution with a concentration of 0.86mol / L was used to obtain an isomerized catalyst with a size of Φ(1.3-1.4)×(3-6)mm, containing 2.4wt% zinc (calculated as ZnO) and 3wt% yttrium (calculated as Y2O3).

[0075] Example 6

[0076] The 2-pentene isomerization catalyst of this example is prepared by basically following steps (1)-(4) of Example 4, with the main difference being:

[0077] In step (2), a zinc acetate aqueous solution with a concentration of 0.94 mol / L was used instead. After drying, the calcination temperature of the powder in the muffle furnace was changed to 550℃ to obtain modified ZSM-5 molecular sieve powder containing 1.5 wt% zinc (calculated as ZnO).

[0078] In step (4), the amount of strip-shaped, zinc-modified ZSM-5 molecular sieve was changed to 950g, and an isomerization catalyst with a concentration of 1.53mol / L yttrium nitrate aqueous solution was used to obtain a catalyst with a size of Φ(1.3-1.4)×(3-6)mm, containing 1.4wt% zinc (calculated as ZnO) and 5wt% yttrium (calculated as Y2O3).

[0079] Comparative Example 1

[0080] The catalyst of this comparative example was prepared by basically following steps (1) and (3)-(4) of Example 4. The main difference is that in step (3), the modified ZSM-5 molecular sieve powder containing 2.1 wt% zinc (calculated as ZnO) was replaced with the HZSM-5 molecular sieve powder from step (1). The catalyst prepared in this comparative example does not contain zinc, that is, the ZSM-5 molecular sieve is not modified with zinc.

[0081] Comparative Example 2

[0082] The partially strip-shaped, zinc-modified ZSM-5 molecular sieve prepared in step (3) of Example 4 was used as the catalyst in this comparative example. The catalyst does not contain yttrium, that is, the ZSM-5 molecular sieve is not modified by yttrium.

[0083] Comparative Example 3

[0084] The catalyst of this comparative example was prepared by basically following steps (1) and (3) of Example 4. The main difference is that in step (3), the modified ZSM-5 molecular sieve powder containing 2.1 wt% zinc (calculated as ZnO) was replaced with the HZSM-5 molecular sieve powder from step (1). The catalyst prepared in this comparative example does not contain zinc or yttrium, that is, the ZSM-5 molecular sieve has not been modified with zinc and yttrium.

[0085] Comparative Example 4

[0086] The catalyst of this comparative example was prepared by basically following steps (1)-(4) of Example 4, with the main difference being that 400 mL of 1.30 mol / L zinc nitrate aqueous solution was used in step (2).

[0087] Comparative Example 5

[0088] The catalyst of this comparative example was prepared by basically following steps (1)-(4) of Example 4. Its zinc yttrium content is the same as that of the catalyst in Example 4. The main difference is that 400 mL of 0.883 mol / L yttrium nitrate aqueous solution was used in step (2), and 1.69 mol / L zinc acetate aqueous solution was used in step (4).

[0089] Comparative Example 6

[0090] Dissolution experiments revealed that a high-concentration aqueous solution containing 1.30 mol / L zinc acetate and 0.6-0.83 mol / L yttrium nitrate could not be prepared at temperatures below 50°C. Therefore, it was impossible to prepare a catalyst with the same zinc yttrium content as in Example 4 through steps (1)-(3).

[0091] Catalyst Evaluation Example 1

[0092] In a dehydrogenation pilot plant (the same plant used in Catalyst Evaluation Example 1 of CN120887769A, located at the production site of an isobutane dehydrogenation unit), the performance evaluation test of the alkane dehydrogenation catalyst HA of Example 2 was conducted under fixed-bed and quasi-adiabatic reactor conditions for n-pentane dehydrogenation. The evaluation method and process of the dehydrogenation catalyst (including the catalyst reduction process and the chlorinated regeneration process) were basically the same as the isopentane dehydrogenation evaluation process in CN120887769A (Catalyst Evaluation Example 1, paragraphs 0081-0109 of the specification), with the main difference being that the inlet gas flow temperature of the reactor was reduced.

[0093] During the evaluation process, the hydrogen gas stream used came from the hydrogen pipeline network of the n-butane dehydrogenation production unit, with a molar purity of over 98% and a hydrocarbon content of less than 350 mg / Nm³. 3 Unsaturated hydrocarbon content is less than 100 mg / Nm 3 The dew point is below -18℃; the carbon dioxide gas stream is generated by the vaporization of liquid carbon dioxide in the steel cylinder, with a molar purity of over 99.98% and a hydrocarbon content of less than 3 mg / Nm3; the main composition of the n-pentane used is: n-pentane 99.5 wt%, isopentane 0.26 wt%, components below C4 0.11 wt%, components above C6 0.09 wt%, total olefins less than 0.04 wt%, and total sulfur 0.6 mg / kg.

[0094] After evaluating the conditions of catalyst HA, 1000 mL of catalyst HA was reloaded into each dehydrogenation reactor, and process tests were conducted under the operating conditions listed in Table 1, in Mode B. The inlet gas flow temperature of reactors I and II was controlled within the range of 507-510℃. In Table 1, the n-pentane mass hourly space velocity is 2 h⁻¹. -1 That is, a n-pentane liquid flow rate of 2000 g / h is equivalent to a n-pentane flow rate of 620.9 NL / h after vaporization. The CO2 molar content is a calculated value. The CO2 molar content in the inlet gas stream of reactor I is adjusted by adjusting the flow rate of carbon dioxide gas stream I. The CO2 molar content in the inlet gas stream of reactor II is adjusted by adjusting the flow rate of carbon dioxide gas stream II based on the estimated flow rate and CO2 content of the outlet gas stream of reactor I. The control of the CO2 molar content in the two inlet gas streams is also based on the principle of ensuring that the temperature of the upper, middle and lower sections of the catalyst bed in reactors I and II does not exceed 520℃, while taking into account the n-pentene yield.

[0095] Table 1. Evaluation of operating conditions for the n-pentane dehydrogenation reaction during the first operating cycle of catalyst HA.

[0096]

[0097] Under the operating conditions listed in Table 1, the temperature difference between the inlet gas flow and the oven temperature of both reactors was less than 3℃. Specific effects of the n-pentane dehydrogenation reaction included: during the evaluation test, the maximum temperature of the catalyst bed in both reactors could be controlled below 520℃; the temperature of the catalyst bed outlet section in reactor I could be controlled above 470℃, and the temperature of the catalyst bed outlet section in reactor II could be controlled above 500℃; a stable overall reaction effect of over 25% total n-pentane conversion (operating time of over 28% n-pentane conversion was 133h), over 90% selectivity for n-pentene, and over 80% total CO2 methanation rate could be achieved. Reactor I could stably achieve a reaction effect of over 14% n-pentane conversion, over 92% selectivity for n-pentene, and over 82% CO2 methanation rate. By the 455th hour of evaluation, the total n-pentane conversion rate had decreased to below 25%, and the system pressure drop, i.e., the bed pressure drop of both reactors, had increased to 0.012MPa, indicating significant coking had occurred.

[0098] When the evaluation reaches the 458th hour, shut off the n-pentane and carbon dioxide feed, shut off the cooling medium circulation of condenser B (the circulating water of condenser A continues), fully open the back pressure valve, end the first operating cycle, and perform the first in-situ regeneration according to the method in paragraph 0101 of the CN120887769A instruction manual, including carbonization, oxychlorination, and reduction operations.

[0099] The second operating cycle of catalyst HA was then evaluated using the n-pentane dehydrogenation reaction under the operating conditions described in Table 1. The results showed a near-repeat of the first operating cycle's performance, consistently achieving a total n-pentane conversion of over 25% (with a total n-pentane conversion of over 28% achieved within 115 hours), a n-pentene selectivity of over 90%, and a total CO2 methanation rate of over 80%. However, by the 443rd hour of evaluation, the total n-pentane conversion had dropped below 25%, and the system pressure drop had increased to 0.010 MPa, indicating significant coking had occurred.

[0100] Following the first in-situ regeneration method described above, the catalysts in both reactors underwent a second in-situ regeneration. Subsequently, an evaluation test of the n-pentane dehydrogenation reaction during the third operating cycle of catalyst HA was conducted to further investigate the operational effectiveness of Mode B. Specific conditions included: inlet gas flow temperature of 540℃, pressure of 0.05 MPa, and n-pentane mass hourly space velocity of 2.5 h⁻¹. -1The molar ratio of n-pentane to H2 was 1:0.5; the inlet gas temperature of reactor II was 490℃; by adjusting the flow rates of carbon dioxide gas streams I and II, the overall temperature of the catalyst beds in reactors I and II was controlled to a moderately high level, but the maximum temperature never exceeded 520℃. Results included: the outlet temperatures of the catalyst beds in reactors I and II could be maintained above 470℃ and above 500℃, respectively; a stable overall reaction effect of over 25% total n-pentane conversion (operating time above 28% n-pentane conversion was 146h), over 90% n-pentene selectivity, and over 80% total CO2 methanation rate could be achieved, with reactor I consistently achieving over 15% n-pentane conversion, over 90% n-pentene selectivity, and over 82% CO2 methanation rate. At 431h, the total n-pentane conversion rate dropped below 25%, and the system pressure drop increased to 0.015MPa, indicating significant coking had occurred.

[0101] Following the first in-situ regeneration method described above, the catalysts in both reactors underwent a third in-situ regeneration. Before reduction, heater III and reactor III were connected (other connections were appropriately modified, and reactor III was filled with new catalyst HA for reduction). The dehydrogenation reaction of n-pentane in mode C and the fourth operating cycle was evaluated. The carbon dioxide inlet was shut off, and water was injected through the n-pentane pipeline (pressurized and flow controlled by a plunger flow pump). Specific conditions included: inlet gas temperature of reactor I 530℃, pressure 0.03MPa, and n-pentane mass hourly space velocity 2h⁻¹. -1 The molar ratio of n-pentane to H2 and water was 1:0.5:2.5 (without CO2 added); the inlet gas flow temperature of reactors II and III was 525℃. Results showed that the outlet temperatures of the catalyst beds in reactors I, II, and III could be maintained above 460℃, 470℃, and 490℃, respectively; a stable overall reaction effect of over 25% total n-pentane conversion (operating time above 28% n-pentane conversion was 85 hours) and over 92% n-pentene selectivity could be achieved, with reactor I consistently achieving over 12% n-pentane conversion and over 92% n-pentene selectivity. By the 503rd hour of evaluation, the total n-pentane conversion had dropped below 25%, and the system pressure drop (i.e., the bed pressure drop of the two reactors) had increased to 0.017 MPa, indicating significant coking had occurred.

[0102] In the four operating cycles described in Example 1, the main dehydrogenation products and side reactions of n-pentane included: among the main dehydrogenation product n-pentene, 1-pentene accounted for 15-20% of the molar percentage, cis-2-pentene accounted for 17-32% of the molar percentage, and the remainder was mainly trans-2-pentene; the main byproducts were low-carbon hydrocarbons of C4 and below (accounting for more than 85% of the byproducts by carbon), and minor byproducts included isopentane and dienes, but in very small quantities.

[0103] As can be seen from the overall reaction effect of this evaluation example 1, under the operating conditions, catalyst HA has both high performance in the n-pentane dehydrogenation reaction and moderate performance in the CO2 hydromethanization reaction. The exothermic CO2 methanation reaction significantly improves the conversion rate of n-pentane in the dehydrogenation reaction, while basically not increasing side reactions. In the above evaluation process, a n-pentane conversion rate of more than 25% was stably obtained at the lower reaction temperature, the n-pentene selectivity was stably maintained at more than 90%, and the regeneration cycle was more than 400 hours. Controlling the catalyst bed temperature to the lower level is the key condition for obtaining a longer regeneration cycle. The in-situ chlorine regeneration can basically restore the reaction performance of catalyst HA. The conditions of each reactor are relatively easy to achieve and have a certain degree of controllability.

[0104] Catalyst Evaluation Example 2

[0105] In another dehydrogenation pilot plant identical to that used in Catalyst Evaluation Example 1, the Pt-Sn-K / alumina type alkane dehydrogenation catalyst JB from Example 3 was used to conduct a process test for the dehydrogenation of n-pentane to n-pentene, as described in Example 2, under fixed-bed and quasi-adiabatic reactor conditions. The test methods, the source and purity of the n-pentane, carbon dioxide, and hydrogen used were exactly the same as in Catalyst Evaluation Example 1.

[0106] In this evaluation example 2, the evaluation results of the n-pentane dehydrogenation of catalyst JB include: the temperature at the catalyst bed outlet section of each reactor is exactly the same as in evaluation example 1; in the first operating cycle, the operating time with a total n-pentane conversion rate of 25% or higher is 571 h (the operating time with a conversion rate of 28% or higher is 181 h); in the second, third, and fourth operating cycles (the catalyst underwent three chlorine regenerations), the operating times with a total n-pentane conversion rate of 25% or higher are 568 h (the operating time with a conversion rate of 28% or higher is 193 h), 543 h (the operating time with a conversion rate of 28% or higher is 172 h), and 615 h (the operating time with a conversion rate of 28% or higher is 194 h), respectively.

[0107] In the four operating cycles described in this evaluation example 2, the main dehydrogenation products and side reactions of n-pentane included: among the main dehydrogenation product n-pentene, 1-pentene accounted for 15-20% of the molar percentage, cis-2-pentene accounted for 20-35% of the molar percentage, and the remainder was mainly trans-2-pentene; the main byproducts were low-carbon hydrocarbons of C4 and below (accounting for more than 87% of the byproducts by carbon), and minor byproducts included isopentane and dienes, but in very small quantities.

[0108] The overall reaction performance of Example 2 shows that under the operating conditions described, catalyst JB exhibits both high performance in the n-pentane dehydrogenation reaction and moderate performance in the CO2 hydromethanization reaction. The exothermic CO2 methanation reaction significantly improves the conversion rate of n-pentane in the dehydrogenation reaction without significantly increasing side reactions. During the evaluation process, a n-pentane conversion rate of over 25% was stably achieved at the relatively low reaction temperature, the n-pentene selectivity was stably maintained at over 91%, and the regeneration cycle was over 400 hours. After in-situ chlorine regeneration, the reaction performance of catalyst JB was basically restored, and the conditions of each reactor were relatively easy to achieve and had a certain degree of controllability.

[0109] Catalyst Evaluation Example 3

[0110] In an isomerization catalyst evaluation apparatus, the isomerization catalysts prepared in Examples 4-6 were subjected to performance evaluation tests on 2-pentene isomerization under fixed bed and quasi-isothermal reactor conditions.

[0111] The isomerization evaluation device mainly includes a vertical fixed-bed quasi-isothermal reactor with an axial reaction structure and a post-condenser.

[0112] The reactor body is a cylindrical stainless steel tube (Φ36×2mm), which is densely packed from bottom to top with 240mL of alumina ceramic balls (Φ2mm) and 80mL of isomerization catalyst (catalyst cut to 1-1.5mm, bed height approximately 105mm, height-to-diameter ratio 3.3). A thermocouple sheath (Φ6×1mm) extends along the centerline of the reactor tube from the lower end, through the isomerization catalyst section, to the middle of the alumina ceramic ball section. Three K-type armored thermocouples with an outer diameter of 1mm are inserted into the thermocouple sheath. Two of the thermocouples are fixed and measure the bottom temperature of the alumina ceramic ball section and the middle temperature of the catalyst bed, respectively. The third thermocouple can move up and down in the isomerization catalyst section to measure the temperature of the catalyst bed in the vertical direction. A 10mm thick brass homogenizing sleeve is embedded outside the stainless steel reactor tube, which is then installed in an electric heating furnace. The temperature of the middle of the catalyst bed is used as the control temperature of the electric heating furnace and is set as needed. The 2-pentene liquid flow from the plunger flow pump enters from the top of the reactor, is vaporized and preheated in the alumina ceramic ball section, and flows downward through the catalyst bed. The 2-pentene feed bottle (500mL glass bottle) from the plunger flow pump is weighed and replenished at regular intervals.

[0113] The condenser is made of stainless steel tubing with a diameter of 6×1×1000mm and is placed in a cold water bath tank at 0-5℃ (placed in a small freezer with temperature control at 0℃). The hot gas flow discharged from the lower end of the isomerization reactor is cooled by the condenser and the liquid product is separated. After being depressurized by the back pressure valve, it is then sent to a humidified gas flow meter (the medium is water). The separated liquid product is temporarily stored in a 500mL stainless steel intermediate tank. It is discharged and sampled for chemical composition testing at regular intervals (once every 1 hour) through a needle valve. The liquid product discharged from the intermediate tank at regular intervals is sent through a polyethylene hose to a 5L plastic storage tank (also placed in the small freezer with temperature control at 0℃, and weighed, emptied, or replaced at regular intervals).

[0114] In the isomerization catalyst evaluation apparatus, the catalysts of Examples 4-6 were sequentially evaluated for their initial activity in the 2-pentene isomerization reaction. First, the catalyst of Example 4 was evaluated using exploratory conditions, and the suitable temperature range was determined to be 375-390℃. Then, a new catalyst from Example 4 was loaded, and a formal evaluation was conducted. Specific reaction conditions and effects are shown in Table 2. During the isomerization evaluation under the conditions in Table 2, the temperature difference between the bottom of the alumina ceramic spheres and the middle of the catalyst bed could be controlled within 3℃. The main composition of the 2-pentene used was: 99.5 wt% 2-pentene (27.9 wt% cis-2-pentene, the remainder being trans-2-pentene), 0.38 wt% 1-pentene, and 0.3 mg / kg total sulfur.

[0115] Table 2. Initial activity evaluation of the catalyst in the 2-pentene isomerization reaction of Example 4: reaction conditions and effects.

[0116]

[0117] Subsequently, the catalysts of Examples 5 and 6 were evaluated for their initial activity in the 2-pentene isomerization reaction. The targets for both were a 2-pentene conversion exceeding 20% ​​and a 1-pentene selectivity exceeding 95%. The temperature, pressure, and 2-pentene mass hourly space velocity in the catalyst bed were adjusted, and the evaluation time was 50 hours for each. It was found that the evaluation results and reaction patterns were essentially the same as those of the catalyst in Example 4 listed in Table 2. The catalysts were unloaded after evaluation; all three unloaded catalyst samples were light gray (new catalyst is white), indicating that only slight coking occurred.

[0118] Initial activity evaluation results show that these three catalysts are effective at temperatures of 375-390℃, pressures of 0.1-0.15 MPa, and 2-pentene mass hourly space velocity (HHSV) of 1-2 h⁻¹. -1 Under the given reaction conditions, all catalysts exhibited good performance in converting 2-pentene to 1-pentene, with the catalyst in Example 4 showing the best overall performance in terms of both reaction performance and stability.

[0119] In the isomerization catalyst evaluation device, the unloaded catalyst from Example 4 was reloaded, and the activity evaluation and cycle test of the 2-pentene isomerization reaction were continued. With the goal of achieving a 2-pentene conversion of over 20% and a 1-pentene selectivity of over 95%, the reaction was conducted at a temperature of 375-390℃ (starting from 375℃, the reaction temperature was gradually increased as the 2-pentene conversion decreased), a pressure of 0.125 MPa, and a 2-pentene mass hourly space velocity of 1.5 h⁻¹. -1 The reaction was carried out under the specified conditions until the target effects of 2-pentene conversion exceeding 20% ​​and 1-pentene selectivity exceeding 95% could not be achieved simultaneously. The total evaluation time (including the initial activity evaluation time of 50 hours) reached 368 hours.

[0120] After stopping the 2-pentene feed, the reactor was purged with high-purity nitrogen at a flow rate of 1 L / min for 10 min, and then the outlet gas flow was exhausted. In-situ carbonization regeneration was then performed: the set temperature of the electric heater (continuing to use the temperature of the middle part of the catalyst bed as the control temperature of the electric heater) was increased to 460℃. An oxygen-containing gas flow with a flow rate of 1.33 L / min, an O2 molar content of 0.5-1.2%, and a mixture of ordinary nitrogen gas (N2 molar content above 99%) and air was introduced from the top feed pipe of the reactor, with a space velocity of 1000 h⁻¹. -1 The pressure was atmospheric pressure, the initial bed temperature was approximately 389℃, and the maximum bed temperature was controlled to not exceed 530℃. The O2 molar content difference between the outlet and inlet gas streams was reduced to below 0.1%. Then, the O2 molar content in the oxygen-containing stream was increased to 7.5%, and the maximum bed temperature was controlled to not exceed 550℃ for 6 hours. Afterward, the set temperature of the electric heater was lowered to 380℃, and purging was performed successively with a flow rate of 0.67 L / min for ordinary nitrogen and then for high-purity nitrogen, at a space velocity of 500 h⁻¹. -1 The isomerization reaction is carried out only after the O2 molar content in the outlet gas stream is reduced to below 0.1% and the maximum bed temperature is reduced to below 390℃.

[0121] In Example 4, the catalyst was regenerated twice, and a total of three operating cycles of the 2-pentene isomerization reaction evaluation process were carried out. The overall reaction performance of the catalyst did not decrease significantly, and the reaction results of 2-pentene conversion exceeding 20%, 1-pentene selectivity exceeding 95%, and operating time exceeding 350 hours were obtained.

[0122] In the evaluation of the 2-pentene isomerization reaction activity of the catalysts in Examples 4-6 in this Evaluation Example 3, the byproducts were mainly dimerization and trimerization products of 2-pentene, and the yield was less than 5% of the 1-pentene yield. Almost no cracking reaction products or skeletal isomerization reaction products were detected, and there was no gas flow in the wet gas flow meter, indicating that hydrogen bond transfer and oligomerization reactions mainly occurred. Comparative analysis of the reaction effects suggests that the 2-pentene isomerization reaction effect of the catalysts in Examples 4-6 in this Evaluation Example 3 is significantly better than the isomerization reaction effect of the catalyst disclosed in CN104557407A under its operating conditions. In fact, CN104557407A did not fully disclose the specifications of the ZSM molecular sieve used.

[0123] Catalyst Evaluation Example 4

[0124] In the evaluation apparatus for the isomerization catalyst used in Example 3, the catalysts prepared in Comparative Examples 1-5 were evaluated successively, with the target of a 2-pentene conversion exceeding 20% ​​and a 1-pentene selectivity exceeding 95%, under conditions of 375-390℃, 0.125 MPa, and a 2-pentene mass hourly space velocity of 1.5 h⁻¹. -1 With appropriate adjustments to the reaction conditions within the specified range, an evaluation test was conducted to assess the performance of the 2-pentene isomerization reaction.

[0125] During the evaluation of the catalysts in Comparative Examples 1-5, the evaluation was stopped and the catalyst was unloaded when it was confirmed that the 2-pentene conversion rate of more than 18% or the 1-pentene selectivity of more than 90% could not be achieved, or when the 2-pentene conversion rate reached more than 18% or the 1-pentene selectivity reached more than 93% but the performance was not stable enough.

[0126] The evaluation results for each catalyst in Comparative Examples 1-5 include:

[0127] In Comparative Example 1, the 1-pentene selectivity of the catalyst was difficult to control to above 94.2%, and the content of 2-pentene dimerization and trimerization products increased significantly. The catalyst was light gray after discharge (with less coke buildup), indicating that zinc modification has a significant impact on the 1-pentene selectivity of the catalyst.

[0128] The 2-pentene conversion rate of the catalyst in Comparative Example 2 could not be controlled to above 18% after 13 hours of evaluation, and the catalyst was black after discharge (due to extensive carbon buildup), indicating that yttrium modification has a significant impact on the catalyst's reactivity.

[0129] The catalyst in Comparative Example 3 could not achieve a 1-pentene selectivity of over 90%, and the 2-pentene conversion could not be controlled to over 18% after 9 hours of evaluation. The catalyst was black after discharge (due to extensive carbon buildup), indicating that the unmodified catalyst (composed only of HZSM-5 molecular sieve and alumina) was too acidic and had poor reactivity under the condition of 99.5wt% pure 2-pentene feed.

[0130] The catalyst in Comparative Example 4 can achieve a 2-pentene conversion rate of over 20% and a 1-pentene selectivity of over 95%, but the duration of this reaction effect is 247 h, indicating that in the preparation step (2) of the isomerization catalyst, the catalyst with the same concentration of zinc nitrate aqueous solution as the impregnation liquid has lower isomerization reaction performance than zinc acetate aqueous solution.

[0131] The catalyst in Comparative Example 5 achieved a 2-pentene conversion rate exceeding 20% ​​and a 1-pentene selectivity exceeding 95%, but the effect was only maintained for 35 hours. The catalyst was black after removal (due to extensive carbon buildup), indicating that yttrium modification significantly affects the acidity of the inner and outer surfaces of alumina grains in the isomerization catalyst. Different yttrium modification methods have different effects on the catalyst's reaction performance.

Claims

1. A process for preparing n-pentene from n-pentane, comprising a n-pentane dehydrogenation device, a n-pentene separation device, an isomerization reaction and a recycle separation device; the n-pentane dehydrogenation device comprises a fixed bed adiabatic reactor filled with a Pt-Sn-K / alumina type alkane dehydrogenation catalyst, which has a hydromethanation performance under dehydrogenation reaction conditions containing CO2; the n-pentene separation device comprises a temperature reduction separation system and a pressure swing adsorption separation system, from the dehydrogenated gas stream from the n-pentane dehydrogenation device and after temperature reduction, a hydrogen gas stream G-H2 is separated and recycled to the n-pentane dehydrogenation device and sold externally, an unreacted n-pentane gas stream M1 is separated and recycled to the n-pentane dehydrogenation device, and a n-pentene gas stream P-2 is separated and sent to the isomerization reaction and recycle separation device for treatment; the n-pentene molar content of the gas stream P-2 is controlled to be more than 99%; the 2-pentene isomerization catalyst is a ZSM-5 molecular sieve-alumina modified with zinc and yttrium, containing zinc at 1.4-2.6 wt% as ZnO and yttrium at 3-5 wt% as Y2O3, and is prepared by the following steps: (2) the HZSM-5 molecular sieve powder obtained in step (1) is placed in a mixer, and a 95-110% of the internal pore volume of the HZSM-5 molecular sieve powder and a desired concentration of a zinc acetate aqueous solution are uniformly applied by a mist spraying liquid method, and the mixture is placed for homogenization treatment for more than 10 hours and then dried; after drying, the powder is calcined at 500-550℃ for 2-4 hours, cooled, and crushed to obtain a zinc-containing modified ZSM-5 molecular sieve powder; (3) 200-600 mesh pseudo-boehmite powder is added to a kneader, the kneader is started, and a 20-25 wt% concentration of an acetic acid aqueous solution is uniformly applied by a mist spraying liquid method, the weight ratio of the pseudo-boehmite powder to the acetic acid aqueous solution being 100:(10-15), and after the acetic acid aqueous solution is added for 30-60 minutes, the kneader is stopped and placed in a sealed state for more than 10 hours; the kneader is started, the modified ZSM-5 molecular sieve powder and sesbania powder are added, the powder is mixed, and water is added by a mist spraying liquid method, the kneader is operated for 30-60 minutes, and the product is extruded into strips, dried, calcined at 500-550℃ for 2-4 hours, and cooled to obtain a zinc-modified ZSM-5 molecular sieve in the form of strips with a diameter of 1.2-1.5 mm; in the ingredients, the weight ratio of the modified ZSM-5 molecular sieve powder, the pseudo-boehmite powder calculated as Al2O3, the sesbania powder, and the total amount of water in the acetic acid aqueous solution and the additional water is 100:(30-40):(3-5):(100-130); (4) the zinc-modified ZSM-5 molecular sieve in the form of strips obtained in step (3) is added to a spray leaching drum, a yttrium nitrate aqueous solution with a desired concentration is applied by a mist spraying liquid method, the spraying time is 10-20 minutes, the product is placed in a sealed state for homogenization treatment for 10-15 hours, then dried, calcined at 540-560℃ for 2-4 hours, and cooled to obtain an isomerization catalyst in the form of strips with a diameter of 1.2-1.5 mm and a length of 3-6 mm. ​ ​ The isomerization reaction and cyclic separation device comprises a pentene separation tower, a 2-pentene heavy component removal tower and a 2-pentene isomerization reactor group; the normal pentene gas stream P-2 from the normal pentene separation device is mixed with the isomerized gas stream from the 2-pentene isomerization reactor group and cooled, and then fed into the pentene separation tower; the 1-pentene liquid stream product is separated from the pentene separation tower top condenser, and the 2-pentene-rich liquid stream obtained at the bottom is fed into the 2-pentene heavy component removal tower; the heavy hydrocarbon liquid stream P-4 is separated from the 2-pentene heavy component removal tower bottom, and the 2-pentene gas stream is separated from the top condenser and sent to the isomerization reactor group; the isomerization reactor group is provided with 2-3 fixed bed reactors VI filled with 2-pentene isomerization catalyst in parallel, and an olefin heating furnace is arranged in front of the gas stream inlet of each reactor VI; during stable operation, the reaction conditions of each reactor VI include: Temperature 375-390°C, pressure 0.05-0.15 MPa, 2-pentene mass space velocity 1-2 h -1 The 1-pentene content in the 2-pentene-rich liquid stream from the bottom of the pentene separation column is controlled to be below 1 wt%. ​ (1) Na type ZSM-5 molecular sieve powder with relative crystallinity of 95% or more, Si / Al ratio of 300-350, surface area of 500-600 m 2 / g, and average crystal grain size of 0.3-0.5 μm is exchanged with sufficient NH4NO3 aqueous solution with concentration of 1-2 mol / L for 4-5 times at 80-90°C, washed with water, dried, and calcined at 450-500°C to convert into H type HZSM-5 molecular sieve powder; ​ ​ ​ 2. The process for preparing n-pentenes from n-pentane as claimed in claim 1, wherein, Pentene separation column, 2 The number of plates of the pentene deheavy column is 110-130, 25-30 respectively, the reflux ratio at the top is 10-20, 0.5-1 respectively, and the pressure at the top is 0.1-0.2 MPa.

3. The process for making n-pentenes from n-pentane as claimed in claim 1, wherein, The temperature reduction separation system of the normal pentene separation device first cools the normal pentane dehydrogenation gas stream X4 after the heat exchanger T1 to -20℃ to -10℃, increases the pressure to 0.5-1MPa, separates the water, and obtains a gas stream G0 containing H2, CH4, CO2, ethane, and ethylene, and a hydrocarbon-containing liquid stream L0.

4. The process for making n-pentenes from n-pentane according to claim 1, wherein, The pressure swing adsorption separation system includes a PSA-H2 unit for separating hydrogen from the gas stream G0, and an alkene separation unit PSA-CH; the alkene separation unit PSA-CH separates the gasified hydrocarbon-containing liquid stream L to obtain a normal pentane gas stream M1, a normal pentene gas stream P-2, and a heteroalkene gas stream P-3.

5. The process for making n-pentenes from n-pentane according to claim 4, wherein, The adsorbent used in each adsorption tower of the alkene separation unit PSA-CH has an adsorption selectivity of normal pentene / normal pentane in the gasified stream of the hydrocarbon-containing liquid stream L of 6 or more under operating conditions of 40-100℃ and an absolute pressure of 30-1000kPa.

6. The process for making n-pentenes from n-pentane as claimed in claim 1, wherein, The fixed bed reactor VI of the isomerization reactor group, when the reaction performance index drops to a specified range, simultaneously with the preceding olefin heating furnace, is cut off from the isomerization operation and is nitrogen-purged, and then in-situ coke burning regeneration is performed, and the isomerization operation is re-cut in.

7. The process for making n-pentenes from n-pentane as claimed in claim 1, wherein, The isomerization reaction and cyclic separation device is further provided with a heat exchanger T2; 2 The pentene gas stream, after heat exchange between the heat exchanger T2 and the outlet gas stream of each reactor VI, goes to the isomerization reactor group.

8. The process for making n-pentenes from n-pentane as claimed in claim 1, wherein, 0.01-0.03wt% of hydroquinone is added to the 2-pentene heavy-removal column bottom liquid as a polymerization inhibitor.

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