Process, catalyst, and tert-amyl alcohol preparation process for the dehydrogenation of isopentane to isopentene

Abstract

The present application belongs to the technical field of alkane processing, and particularly relates to a process for preparing isoamylene by dehydrogenation of isoamylane, a catalyst and a process for preparing tertiary amyl alcohol. The process for preparing isoamylene by dehydrogenation of isoamylane comprises an isoamylane dehydrogenation device and an isoamylene separation device. In the isoamylane dehydrogenation device, the dehydrogenation reactor is a fixed bed adiabatic reactor, which is loaded with a Pt-Sn-K / alumina type alkane dehydrogenation catalyst. The catalyst has a hydrogen-assisted methanation performance under dehydrogenation reaction conditions containing CO2. The setting and operation of the dehydrogenation reactor group include modes A, B or C. The isoamylene separation device comprises a cooling separation system and a pressure swing adsorption separation system, and an optional rectification system. Tertiary amyl alcohol is further prepared from isoamylene under the condition of a fixed bed hydration reactor and solvents such as ethylene glycol monobutyl ether, and the tertiary amyl alcohol product is separated. The present application is superior to the prior art as a whole, and 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 isopentenene by dehydrogenation of isopentenane, a catalyst, and a process for preparing tert-amyl alcohol. Background Technology

[0002] The dehydrogenation of isopentane to produce isopentene, followed by further conversions such as dehydrogenation to produce isoprene, etherification with methanol to produce methyl tert-amyl ether (TAME), which can be used as a gasoline blending component, and reaction with water to produce tert-amyl alcohol, has certain commercial value. The isopentane dehydrogenation process and the catalyst used are key. Isopentane is a basic product of C5 byproducts from catalytic cracking and naphtha cracking to produce ethylene, and is readily available.

[0003] The dehydrogenation of propane and isobutane to produce propylene and isobutene is a related process for the dehydrogenation of isopentane to produce isopentene. However, due to the increased complexity of the isopentane dehydrogenation products, the process for producing isopentene from isopentane with limited improvements to the isobutane dehydrogenation process and catalyst generally yields limited results.

[0004] As disclosed in CN103232312A, the process for producing isobutene from isobutane by dehydrogenation uses hydrogen and water vapor as diluents and heat carriers. The dehydrogenation reaction of isobutane is carried out in a fixed-bed adiabatic reactor filled with a Pt-Sn-K / alumina type platinum-based dehydrogenation catalyst, and then isobutene is further separated and produced. This process has been industrialized and has achieved good technical results and economic benefits. In this process unit, the isobutane-isopentane dehydrogenation unit employs two dehydrogenation reactor series. Each series includes two dehydrogenation reactors connected in series. The dehydrogenation reactors adopt a radial or axial reaction structure, and each dehydrogenation reactor is equipped with a feed flow heater before it. The catalyst used in the dehydrogenation reactors is a spherical catalyst (outer diameter 2.5-3.0 mm) prepared based on the proportions and methods in CN103212411A. Pt is uniformly distributed and highly dispersed in the spherical particles. The inlet conditions for each hydrogenation reactor are: temperature 580-650℃, pressure 0.1-0.5 MPa, and space velocity 2000-6000 h⁻¹. -1 In each dehydrogenation reactor series, the feed gas volume ratio of the first dehydrogenation reactor is isobutane:hydrogen:water vapor = 1:(0.5-3):(0.5-3). The specific gas inlet conditions are adjusted according to the reaction effect, which can achieve the target reaction effect of isobutane single-pass conversion rate of more than 30% or even more than 35% and isobutene selectivity of more than 97%. After the dehydrogenation catalyst gradually accumulates coke and the reaction effect declines, it needs to be regenerated in situ in the reactor. The regeneration cycle can reach more than 400 hours or even more than 750 hours.

[0005] Two Pt-Sn-K / alumina type alkane dehydrogenation catalysts, which have been applied and achieved good technical results in the industrial apparatus of the process for producing isobutene from isobutane as described in CN103232312A, were used respectively. Specifically, catalyst A in Example 1 of CN103212411A (containing 0.35wt% Pt, 0.7wt% Sn, 1wt% K2O, with a support of 1.1wt% ZrO2 in a high-temperature zirconium-aluminum composite oxide) and catalyst B in Example 4 (containing 0.25wt% Pt, 0.7wt% Sn, 1wt% K2O, with a support of ZrO2 in a high-temperature zirconium-aluminum composite oxide). Based on the formulation and method of 2.2wt% high-temperature zirconium-aluminum composite oxide, a spherical catalyst with the same chemical composition (outer diameter 2.5-3.0 mm, Pt uniformly distributed in the spherical particles, i.e., uniform cross-sectional color and highly dispersed in the initial stage of each operating cycle) was prepared. In the same dehydrogenation pilot plant used in the development of the isobutane dehydrogenation to isobutene process described in CN103232312A, the process test of isopentane dehydrogenation to isopentene was carried out under the conditions of a fixed-bed adiabatic reactor. It was found that there was a significant difference in effect between isopentane dehydrogenation and isobutane dehydrogenation. It was difficult to simultaneously control and obtain isopentane single-pass conversion of more than 30%, isopentene selectivity of more than 90%, and regeneration cycle of more than 300 hours. The dehydrogenation pilot plant comprises two reactors, I and II, connected in series, with reactor I being the preceding reactor. The catalyst loading volume ratio in reactors I and II is 1:1. Each reactor has a heater installed before its gas flow inlet, with the specific configuration being the same as the connection method B in Catalyst Evaluation Example 1 below. The main operating conditions include: pressure 0.02-0.5 MPa, and liquid hourly space velocity (LHSV) of isopentane 0.6-6 h⁻¹. -1 The molar ratio of the feed gas stream to the first dehydrogenation reactor is isopentane:hydrogen:water vapor = 1:(0.5-3):(0-3). The inlet gas temperature of the two dehydrogenation reactors is adjusted in the range of 540-600℃ according to the reaction effect.

[0006] In the dehydrogenation pilot plant, using the same experimental control method, spherical catalysts (outer diameter 2.5-3.0 mm, catalyst E K content increased to 1 wt% (K2O), uniform Pt distribution in the spherical particles, i.e., uniform cross-sectional color and high dispersion in the initial stage of operation) prepared based on the CE ratio and method of catalyst CN103212411A were tested. All catalysts contained 0.35 wt% Pt, 0.7 wt% Sn, and 1 wt% K2O, with supports consisting of high-temperature magnesium-aluminum composite oxide containing 2.27 wt% MgO, high-temperature magnesium-aluminum composite oxide containing 1.13 wt% MgO, and ZnO, respectively. High-temperature zinc-aluminum composite oxide (8.14 wt%), commercially available low-carbon alkane (C3-C6 alkane) dehydrogenation catalysts, and low-carbon alkane dehydrogenation catalysts prepared by other superior existing technologies (such as CN107537485A and Liu Changcheng's article below) were used to conduct process tests on the dehydrogenation of isopentane to isopentene. It was found that all of these methods struggled to simultaneously achieve the basic technical requirements of isopentane single-pass conversion of over 30%, isopentene selectivity of over 90%, and regeneration cycle of over 300 hours. Specifically, the catalyst prepared by the method described in Liu Changcheng's article below could not maintain a regeneration cycle of over 280 hours (with chlorine regeneration) when the inlet temperature of the first dehydrogenation reactor was 540-550℃ and the inlet temperature of the second dehydrogenation reactor was 530-550℃. All tested catalysts were spherical with an outer diameter of 2.2-3.0 mm, a pore volume of over 0.40 mL / g, and a specific surface area of ​​over 80 m². 2 / g or more.

[0007] The catalyst disclosed in CN103212411A for the dehydrogenation of low-carbon alkanes to olefins is made of the following components: the support is a high-temperature composite oxide composed of at least two metal oxides, one of which is aluminum oxide, and the other is one or more of magnesium oxide, zirconium oxide, and zinc oxide; the active component is one or more of ruthenium, rhodium, platinum, and iridium; the first promoter is one or more of germanium, tin, or lead, all of Group IV elements; the second promoter is one or more of alkali metals, including lithium, sodium, potassium, rubidium, cesium, and francium; preferably, the active component is platinum, based on the support... The mass fraction is 0.2-0.5%; the first auxiliary agent is tin, and the molar ratio with the active component is (1-3):1; the second auxiliary agent is potassium, and the mass fraction based on the carrier as oxide is 0.1-3%; the high-temperature composite oxide carrier contains 77-95 wt% alumina, 1-10 wt% magnesium oxide, 1-5 wt% zirconium oxide, or 2-10 wt% zinc oxide; the preparation method of the high-temperature composite oxide carrier is: dissolving one or more soluble salts containing magnesium oxide, zirconium oxide, or zinc oxide in water to prepare a concentration of 0.05-1.0%. The catalyst is prepared by mixing the alumina precursor boehmite with a mol / L aqueous solution, then impregnating it at 0-100°C for 0.1-24 hours, drying it at 30-200°C under a pressure of 0.01-0.1 MPa for 1-24 hours, and finally calcining it in air at 700-1000°C for 2-10 hours. The catalyst is prepared by adding an active component such as chloroplatinic acid and a first auxiliary agent, stannous chloride, to a solution containing inorganic acids (hydrochloric acid or nitric acid) and / or organic acids (maleic acid, oxalic acid, acetic acid, citric acid, tartaric acid, etc.). A stable complex is formed by complexing with one or more lactic acids in an acidic solution. A dissolved second auxiliary agent, such as potassium nitrate, is then added to form a stable impregnation solution. A high-temperature composite oxide support is added to the impregnation solution at a volume ratio of (0.1-20):1. The impregnation time of the high-temperature composite oxide support in the impregnation solution is 0.1-48 hours, and the impregnation temperature is 0-100℃. After impregnation, it is dried in air at 50-200℃ for 3-24 hours and calcined in air at 400-600℃ for 1-24 hours to obtain the catalyst. The catalyst is reduced with hydrogen before contacting the low-carbon alkane reaction gas stream. The reduction is carried out at 400-600℃ for 1-24 hours.

[0008] CN107537485A discloses a dehydrogenation catalyst for the dehydrogenation of isopentane to isopentene, comprising, by weight, the following components: a) 0.001 to 1 part Pt; b) 0.1 to 10 parts Sn element or its oxide; 0.1 to 15 parts alkali metal element or its oxide; d) 74 to 99 parts spinel-type support; wherein preferably, Pt is 0.01 to 0.5 parts, Sn or its oxide is 0.2 to 2 parts, alkali metal or its oxide is 0.1 to 5 parts, preferably K, Na, or Li, and may also contain 0.1 to 10 parts tungsten or its oxide, and the spinel-type support is at least one selected from ZnAl2O4, MgAl2O4, NiAl2O4, CuAl2O4, CoAl2O4, or MnAl2O4. The preparation method of this isopentane dehydrogenation catalyst includes the following steps: adding a calculated amount of tin salt, such as SnCl2, to a calculated amount of H2PtCl6 solution, dissolving it, then adding a calculated amount of potassium nitrate solution, and finally adding a calculated amount of spinel support; grinding and drying under an infrared lamp, drying overnight at 110℃, and calcining in flowing air at 300-600℃ for 2-8 hours to obtain the isopentane dehydrogenation catalyst. The catalyst is used at a reaction temperature of 500-600℃, a reaction pressure of atmospheric pressure, and an isopentane weight hourly space velocity of 0.1-5.0 h⁻¹. -1 The molar ratio of H2 to isopentane is (1-10):1.

[0009] In their paper "Reaction Mechanism of Pt-Sn-K / Al2O3 Catalysis for the Conversion of C5-C7 Alkanes to Alkenes and Aromatics ([J / OL], Acta Petrolei Sinica (Petroleum Processing), 2025-05-07), Liu Changcheng et al. disclosed the following catalyst preparation method: A certain amount of aluminum hydroxide powder was weighed, and an appropriate amount of deionized water was added and stirred. An appropriate amount of dilute nitric acid was added dropwise to acidify the mixture into a sol. Under stirring, a hydrochloric acid solution containing SnCl4·5H2O was added to the sol for acidification. Then, kerosene and fatty alcohol polyoxyethylene ether were added dropwise under stirring. The mixture was then formed into droplets in an oil-ammonia column. The wet droplets were solidified in the ammonia-water phase, removed, rinsed with deionized water, dried, and then heated in a muffle furnace at 65°C. The Sn-Al2O3-containing support was prepared by calcining at 0℃ for 4 hours, followed by a second-stage calcination at 1000℃ for 4 hours. A chloroplatinic acid impregnation solution with a Pt mass fraction of 0.305% (based on dry Al2O3) was prepared using hydrochloric acid as a competitive adsorbent. The amount of hydrochloric acid added could be 0 wt% (no addition), 1 wt%, or 2 wt% of the Sn-Al2O3 support mass. The Sn-Al2O3-containing support was added to the chloroplatinic acid impregnation solution and impregnated for 4 hours, dried, and calcined. Then, it was impregnated with potassium chloride solution for 4 hours, dried, and calcined at 600℃ for 4 hours. Finally, it was reduced with H2 at 580℃ to obtain the Pt-Sn-K / Al2O3 catalyst.

[0010] Therefore, it is necessary to develop a process for preparing isopentene by dehydrogenation of isopentane, using a suitable catalyst, to achieve long-term operation with a single-pass conversion rate of isopentane of 30% or more, preferably 30% or more, isopentene selectivity of 90% or more, and catalyst regeneration cycle of 300 hours or more, preferably 400 hours or more, as well as a process for further application of isopentene. Summary of the Invention

[0011] To address the aforementioned technical problems, this invention provides a process for preparing isopentenene by dehydrogenation of isopentenane, a catalyst, and a process for preparing tert-amyl alcohol.

[0012] The present invention relates to a process for preparing isopentenene by dehydrogenation of isopentenane, comprising an isopentenane dehydrogenation unit and an isopentenene separation unit;

[0013] The isopentane dehydrogenation unit includes a dehydrogenation reactor group, as well as corresponding connecting pipelines, temperature measurement and control components, flow measurement and control components, and valves; the reactor of the dehydrogenation reactor group is a fixed-bed adiabatic reactor with an axial or radial reaction structure, and is filled with a Pt-Sn-K / alumina type alkane dehydrogenation catalyst; the catalyst has hydromethanation performance under CO2-containing dehydrogenation reaction conditions;

[0014] The setup and operation of dehydrogenation reactor units include any of the following modes:

[0015] Mode A: Set up 2-4 reactors, each with the same catalyst loading volume, and install a heater 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, the same below), 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 section temperature of the catalyst bed in each reactor is controlled above 530℃.

[0016] Mode B: Set up 2-3 parallel reactor series, each reactor series including two reactors I and II connected in series sequentially, 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 before the gas flow inlet of each reactor. During the stable operation of isopentane dehydrogenation, the inlet gas flow conditions of reactor I are as follows: 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 of reactor II include: temperature 510-540℃, and molar content of CO2 1-2%. The maximum temperature of the catalyst bed in each reactor is controlled below 550℃. The outlet gas flow temperature of reactor II or the outlet section temperature of the catalyst bed is controlled above 530℃.

[0017] Mode C: Set up 2-3 parallel reactor series, each reactor series including three reactors III, IV, and V connected in series sequentially, with reactor III being 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 installed before the gas flow inlet of each reactor. During the stable operation of isopentane dehydrogenation, the inlet gas flow conditions for reactor III are as follows: temperature 550-560℃, pressure ≤0.2MPa, and isopentane mass hourly space velocity 1.5-2.5h⁻¹. -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℃.

[0018] Specifically, when a reactor in Mode A or a series of reactors in Modes B and C experiences a drop in its reaction performance index 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.

[0019] In the process for preparing isopentene by isopentane dehydrogenation of this invention, the isopentane dehydrogenation unit is further equipped with a heat exchanger T1. A mixture of isopentane liquid flow and / or isopentane-rich gas flow, hydrogen flow and / or hydrogen-rich gas flow, and carbon dioxide flow flow M, along with the post-reaction gas flow P from the dehydrogenation reactor group, undergoes countercurrent heat exchange through heat exchanger T1 to obtain pre-reaction gas flow N and cooled gas flow Q. Pre-reaction gas flow N enters the dehydrogenation reactor group, and cooled gas flow Q goes to the isopentene separation unit. The impurities in the isopentane-rich gas flow and hydrogen-rich gas flow that are detrimental to the isopentane dehydrogenation process and catalyst performance should be controlled below a certain content.

[0020] This invention relates to a process for preparing isopentene from isopentane dehydrogenation. The alkane dehydrogenation catalyst can be spherical or strip-shaped, with Pt uniformly distributed within the catalyst particles (i.e., uniform cross-sectional color) and 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. The catalyst can be a zirconium-containing spherical or strip-shaped catalyst H, prepared by improving the proportions and methods in CN103212411A. The catalyst uses a high-temperature zirconium-aluminum composite oxide containing 0.8-1.5 wt% ZrO2 as a support, with the supported components (based on elemental composition) being Pt 0.2-0.4 wt%, Sn 0.5-1 wt%, and K 0.65-1 wt%, preferably containing 0.35 wt% Pt, 0.7 wt% Sn, and 0.83 wt% K. The support is ZrO2. 1.1 wt% of high-temperature zirconium-aluminum composite oxide spherical or strip-shaped catalyst HA; it was found that catalyst H, especially catalyst HA, under fixed-bed adiabatic reactor and hydrogen-containing conditions, such as the setup and operation conditions of the dehydrogenation reactor group A and B, exhibits good reaction performance for isopentane dehydrogenation to isopentene and matching CO2 hydrogenation to methane. Under the setup and operation conditions of the dehydrogenation reactor group C (without CO2), it exhibits good reaction performance for isopentane dehydrogenation to isopentene. Moreover, the catalytic activity is relatively stable and the deactivation is very slow. Long-term operation results of isopentane single-pass conversion of more than 30%, isopentene selectivity of more than 90%, and catalyst regeneration cycle of more than 300 hours can be obtained.

[0021] The zirconium-containing catalyst H can be prepared by the following steps:

[0022] (H-1) with a pore volume of 0.5-0.7 mL / g and a specific surface area of ​​100-130 m² 2 Using spherical or strip-shaped alumina with dimensions of Φ2.5-3.0mm as a carrier, a zirconium oxychloride aqueous solution of 80-100% of the carrier's saturated water absorption volume and the required concentration is uniformly applied. The carrier is then placed in a sealed container for homogenization treatment for 10-20 hours, dried, and then calcined at 880-930℃ for 2-4 hours and cooled to obtain a zirconium-modified alumina carrier.

[0023] (H-2) The zirconium-modified alumina support obtained in step (H-1) is uniformly coated with a Pt-Sn-acid-K aqueous solution at a concentration of 80-90% of its saturated water absorption volume and a temperature of 20-30℃. The mixture is then placed in a sealed container for homogenization treatment for 10-15 hours, dried, and then calcined at 500-600℃ for 2-5 hours. After cooling, an alkane dehydrogenation catalyst is obtained.

[0024] The Pt-Sn-acid-K aqueous solution described in step (H-2) is prepared by dissolving chloroplatinic acid, stannous chloride, hydrochloric acid, citric acid, and potassium nitrate in an aqueous solution of the required concentration. The preparation method is as follows: dissolve chloroplatinic acid and hydrochloric acid in an aqueous solution, add stannous chloride, potassium nitrate, and citric acid, and mix well to obtain the Pt-Sn-acid-K aqueous solution; wherein the concentration of HCl is 0.05-0.1 mol / L and the concentration of citric acid is 0.2-0.4 mol / L.

[0025] The alkane dehydrogenation catalyst can also be a magnesium-containing spherical or strip-shaped catalyst J: using magnesium-modified alumina containing 1-2 wt% MgO as a support, with the supported components (based on elemental composition) of Pt 0.2-0.4 wt%, Sn 0.5-1 wt%, and K 0.65-1 wt%, and prepared through the following steps:

[0026] (J-1) with a pore volume of 0.5-0.7 mL / g and a specific surface area of ​​100-130 m² 2 Using spherical or strip-shaped alumina with dimensions Φ2.5-3.0mm as a carrier, a magnesium bicarbonate aqueous solution of the required concentration, accounting for 40-80% of the carrier's saturated water absorption volume, is uniformly applied. The mixture is then placed in a sealed container for homogenization treatment for 5-20 hours. Under flowing air conditions, the temperature is raised to 70-75℃ and kept constant for 3-5 hours. After drying, the mixture is calcined at 880-930℃ for 2-4 hours and then cooled to obtain a magnesium-modified alumina carrier.

[0027] (J-2) The magnesium-modified alumina support obtained in step (J-1) is uniformly coated with a Pt-Sn-EDTA-K aqueous solution at a concentration of 80-90% of its saturated water absorption volume and a temperature of 40-70℃. The mixture is then sealed and kept at a warm temperature for 3-8 hours to homogenize. After drying, it is calcined at 500-600℃ for 2-5 hours and then cooled to obtain an alkane dehydrogenation catalyst.

[0028] The Pt-Sn-EDTA-K aqueous solution described in step (J-2) is prepared by mixing diammonium nitrite (Pt(NH3)2(NO2)2), potassium stannate (K2SnO3), EDTA (ethylenediaminetetraacetic acid), nitric acid, potassium hydroxide, and potassium nitrite to form an aqueous solution of the required concentration. The preparation method is as follows: (A) Potassium stannate, such as potassium stannate trihydrate, is added to a potassium hydroxide aqueous solution with a pH of 10-12 and dissolved. EDTA (powder) and nitric acid aqueous solution are added, and the reaction is carried out (heated or stirred for a long time) until clear. The pH is adjusted to 6.0-6.5 by adding potassium hydroxide aqueous solution. Potassium nitrite or its aqueous solution is added, dissolved, and / or mixed to obtain the Sn-EDTA-K aqueous solution; wherein the molar ratio of Sn, EDTA, and nitric acid is 1:(1-1.03):(2.1-2.3); (B) In step (A), the Sn-EDTA-K aqueous solution is heated to 50-70℃, and diammonium platinum nitrite aqueous solution at 60-70℃ is added and stirred to obtain Pt-Sn-EDTA-K aqueous solution.

[0029] Catalyst J exhibits good reactivity in a fixed-bed adiabatic reactor and under hydrogen-containing conditions, such as under the setup and operation conditions of dehydrogenation reactor group modes A and B, for the dehydrogenation of isopentane to isopentene and the matching hydrogenation of CO2 to methane. Under the setup and operation conditions of dehydrogenation reactor group mode C (without CO2), it also exhibits good reactivity in the dehydrogenation of isopentane to isopentene, and the overall catalytic activity of the reactor is relatively stable with slow deactivation. Under the setup and operation conditions of reactor group modes A and B, long-term operation effects of isopentane single-pass conversion of over 30% or even over 35%, isopentene selectivity of over 90%, and catalyst regeneration cycle of over 400 hours can be achieved. Under the setup and operation conditions of reactor group mode C, long-term operation effects of isopentane single-pass conversion of over 30%, isopentene selectivity of over 90%, and catalyst regeneration cycle of over 600 hours can also be achieved.

[0030] This invention relates to a process for preparing isopentene from isopentane dehydrogenation. Under the setup and operating conditions of modes A and B, the dehydrogenation reactor group of the isopentane dehydrogenation unit, when using either the zirconium-containing catalyst H or the magnesium-containing catalyst J, ultimately achieves a single-pass isopentane conversion rate of over 30% or even over 35%, isopentene selectivity of over 90%, CO2 methanation rate of over 80%, and a catalyst regeneration cycle of over 300 hours for long-term operation. This indicates that both catalysts possess CO2 hydromethanation performance that matches the isopentane dehydrogenation performance, effectively coupling the two reactions. By utilizing the exothermic nature of the CO2 hydromethanation reaction, the conversion rate of isopentane dehydrogenation is significantly improved. Since isopentane dehydrogenation is a significantly endothermic reaction, and CO2 hydromethanation is a strongly exothermic reaction, if the catalyst's CO2 hydromethanation performance is significantly dominant, the maximum temperature of the catalyst bed can easily exceed 560℃, leading to a decrease in isopentene selectivity and a shortened catalyst regeneration cycle. During the dehydrogenation operation, the temperature and / or CO2 content of the inlet gas flow to each reactor can be adjusted appropriately according to the overall performance of the reactor. For example, at the beginning of the operating cycle when the overall catalyst performance is high, a lower temperature and / or CO2 content of the inlet gas flow can be used. At the end of the operating cycle when the overall catalyst performance declines, the temperature and / or CO2 content of the inlet gas flow can be appropriately increased to obtain a relatively stable isopentene yield and production. Due to the use of a relatively low gas flow inlet temperature, the catalyst bed temperature is also lower than that of existing technologies, thus reducing side reactions and extending the regeneration cycle. The CO2 methanation rate can be estimated from the CO2 conversion rate and the distribution and content of the generated oxygen-containing compounds. In the operation of modes A and B, the CO2 conversion rate easily exceeds 90% or even 95%.

[0031] The magnesium-containing catalyst J exhibits better overall performance and greater operational flexibility than the zirconium-containing catalyst H during its dehydrogenation application before the first regeneration of the dehydrogenation reactor. This indicates that the dispersion and binding state of Pt and Sn reached a high level during the first regeneration cycle, and that the magnesium modification effectively eliminated the long-term and special effects of acidity on the alumina support surface. However, the operating conditions and effects in the second and subsequent cycles depend on the regeneration method. In the process of preparing isopentene from isopentane dehydrogenation of this invention, reduction is required after the reactor is filled with new catalyst and before the dehydrogenation operation. The final step of 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℃.

[0032] For reactors or series loaded with catalyst H or catalyst J, the in-situ regeneration process after the dehydrogenation operation is stopped, when the overall reaction performance indicators drop to a specified range, can be the same as the in-situ regeneration process of Pt-Sn-K / alumina catalysts for isobutane dehydrogenation to isobutene in existing related technologies, or essentially the same as the external regeneration process of semi-regenerated platinum reforming catalysts in refineries, including carbonization, oxychlorination, and reduction operations. Specifically, the carbonization operation uses an oxygen-containing air stream with an O2 molar content of 0.5-0.8%, obtained by mixing nitrogen (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℃, 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. -1 The pressure is 0.02-0.2 MPa, and the initial bed temperature is 480-520℃. First, water or steam, along with 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 heater or furnace. The inlet temperature of the gas stream is controlled at 500-520℃, with the amount of chloride added being 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 further water or chloride addition) is controlled at 520-540℃, and treatment is carried out for 4-6 hours, with the reactor outlet gas temperature above 515℃. After the oxychlorination operation is completed, the reactor is purged with a high-purity nitrogen gas stream (N2 molar content above 99.9%) at 480-500℃, 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 H or catalyst J is reduced under the aforementioned reduction operating conditions. After in-situ regeneration including the oxychlorination operation, the coke deposits of catalyst H or catalyst J are basically burned off, the Pt particles are redispersed, and the isopentane dehydrogenation reaction performance and CO2 hydromethanization performance are basically restored. Among them, catalyst H can basically recover to the effect level of the first operating cycle, while the overall performance of catalyst J is slightly lower than the effect level of the first operating cycle without chlorine, indicating that the dispersion and binding state of Pt and Sn has been restored to a higher level, but the overall performance of catalyst J is still higher than that of catalyst H.

[0033] Reactors or series loaded with catalyst H and reactors or series loaded with catalyst J, after undergoing one and multiple in-situ regeneration treatments under chlorine conditions, essentially recover their reaction performance. When performing isopentane dehydrogenation under the settings and operating conditions of the dehydrogenation reactor group modes A, B, and C, they can all achieve long-term operating results with a single-pass isopentane conversion rate of over 30% or even over 35%, isopentene selectivity of over 90%, methanation rate of over 80% in the presence of CO2, and a catalyst regeneration cycle of over 300 hours. Specifically, reactors or series loaded with catalyst J, after undergoing one and multiple in-situ regeneration treatments under chlorine conditions, can maintain a catalyst regeneration cycle of over 500 hours when performing isopentane dehydrogenation under the settings and operating conditions of the dehydrogenation reactor group modes A and B.

[0034] Reactors or series loaded with catalyst J can also undergo in-situ regeneration under chlorine-free conditions after the overall reaction performance indicators drop to the specified range and the dehydrogenation operation is stopped. This includes carbonization, oxidation, and reduction operations. Specifically, the carbonization operation uses an oxygen-containing airflow with an O2 molar content of 0.5-0.8%, obtained by mixing a nitrogen stream with a dry air stream (e.g., dew point below -20°C), at a space velocity of 500-2000 h⁻¹. -1 The inlet temperature is 450-490℃, the pressure is 0-0.1MPa, and the initial bed temperature is 450-500℃. During the charcoal burning process, the maximum bed temperature is controlled to not exceed 520℃. The process continues for 2-4 hours until the difference in O2 molar content between the outlet and inlet air streams decreases to below 0.1%. For oxidation operations, an oxygen-containing air stream with an O2 molar content of 10-20%, obtained by mixing nitrogen with dry air, or a dry air stream can be used directly, with a space velocity of 500-2000 h⁻¹. -1 The pressure is 0.02-0.2 MPa, the initial bed temperature is 480-520℃, the gas inlet temperature is 500-520℃, and the oxidation treatment lasts for 6-10 hours. After the oxidation operation is completed, the reactor is purged with nitrogen gas at 480-500℃ and high-purity nitrogen gas, with a space velocity of 500-2000 h⁻¹. -1Once the O2 molar content in the outlet gas stream is reduced to below 0.2%, catalyst J is reduced again under the aforementioned reduction operating conditions. Reactors or series loaded with catalyst J, after undergoing in-situ regeneration under the aforementioned chlorine-free conditions, can achieve long-term operating results such as isopentane single-pass conversion of over 30% or even over 35%, isopentene selectivity of over 90%, methanation rate of over 80% in the presence of CO2, and catalyst regeneration cycle of over 300 hours or even over 450 hours when performing isopentane dehydrogenation under the settings and operating conditions of the aforementioned dehydrogenation reactor group modes A, B, and C. In particular, the isomerization side reaction problem leading to n-pentane is less severe compared to catalyst J after in-situ regeneration including oxychlorination operations. After in-situ regeneration under chlorine-free conditions, catalyst J showed significant reduction in carbon deposits, and its isopentane dehydrogenation and CO2 hydromethanization performance was largely restored (slightly lower than the performance during the first operating cycle). This indicates that the dispersion and binding state of Pt and Sn can be restored to the required level, and the overall alkalinity and anti-coking effect of the catalyst can be basically maintained. The in-situ regeneration under chlorine-free conditions reduces the chlorine resistance requirements of reactors, pipelines, control components, and detection components, avoiding chloride corrosion and KCl formation. The catalyst's continued overall alkalinity and resulting good anti-coking ability compensate to some extent for the limited redispersion effect of Pt during regeneration under chlorine-free conditions.

[0035] In the process for preparing isopentene by dehydrogenation of isopentane of this invention, the isopentene separation device includes a cooling separation system and a pressure swing adsorption separation system. The cooling separation system first further cools and pressurizes the isopentane dehydrogenation gas stream Q 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 removal tower D0 (theoretical plate number 30-36), and a small amount of low-boiling gas stream G-1 containing C4 and below components and some 3-methyl-1-butene is separated from the top of the tower, and the hydrocarbon-containing liquid stream L is separated from the bottom of the tower. The pressure swing adsorption (PSA) unit includes a PSA-H2 unit for separating hydrogen from gas flow G0, and one or more PSA-CH units for separating hydrocarbons. The PSA-H2 unit employs conventional PSA technology for hydrogen extraction adsorbents, equipment configuration, and separation operation processes, while the PSA-CH unit employs known technology for hydrocarbon 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 isopentane dehydrogenation unit, and the remainder is supplied externally. A low-hydrogen gas flow G-2 containing CH4, ethane, ethylene, and CO2 is also obtained and used as fuel gas in the heater. The PSA-CH unit preferably includes two hydrocarbon separation units: an alkane-alkene separation unit PSA-CH-1 and an alkane separation unit PSA-CH-2. The PSA-CH-1 unit separates the vaporized hydrocarbon-containing liquid stream L into three streams: an alkane stream GL-1 of the desired purity, an isopentenene stream Z-2, and a heterohydrocarbon stream Z-3. The resulting isopentenene stream Z-2 is easily controlled to achieve a molar purity of over 98% or even 99%. Dienes are primarily absorbed into the heterohydrocarbon stream Z-3, and the amount of heterohydrocarbon stream Z-3 is very small and can be heated. The furnace is used as fuel or for liquefaction and collection, which is then sold externally. The alkane gas stream GL-1 is further separated in the PSA-CH-2 unit to obtain two streams: isopentane stream Y-1 and n-pentane stream Z-1. The isopentane molar purity of stream Y-1 and the n-pentane molar purity of stream Z-1 are easily controlled to above 96% or even above 98%. Isopentane stream Y-1 is recycled in the isopentane dehydrogenation unit, while n-pentane stream Z-1 is liquefied and collected for external supply or converted to isopentane via an isomerization unit before being used in the isopentane dehydrogenation unit. The resulting isopentene stream Z-2 can be further liquefied to obtain isopentene liquid stream Z-2L. A small amount of low-boiling gas stream G-1 containing C4 and below components and 3-methyl-1-butene is separated from the top of the light-light removal tower D0 and also used as fuel in the heating furnace.

[0036] The process of the cooling and separation system may include: cooling the gas stream Q to 20-60℃ (e.g., through a heat exchanger with circulating water), pressurizing it to 0.5-1MPa, cooling it to 20-60℃, then cooling it successively 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 stream G0 and hydrocarbon-containing liquid stream 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 the setting and operation modes A and B of the dehydrogenation reactor group, this is the water produced by the CO2 hydromethanation reaction), During the cooling process from -20 to -10℃, multiple cooling heat exchangers can be used in rotation to perform de-icing operations (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 stream. In mode C, since a large amount of water vapor is introduced into the inlet gas flow of reactor III, the gas flow Q is continuously separated into condensate by an 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 a gas-liquid separator into a gas flow G0 mainly containing H2, CH4 and a small amount of CO2, ethane, and ethylene, and a hydrocarbon-containing liquid flow L0.

[0037] The adsorbents used in each adsorption tower of the PSA-CH-1 unit, under their operating conditions (e.g., 40-100℃, absolute pressure 30-1000 kPa), should have an adsorption selectivity of 3 or higher, preferably 4 or higher, for isopentene / pentane in the vaporized stream of hydrocarbon liquid flow L, meaning they are more likely to adsorb isopentene. The pentane refers to the total amount of n-pentane and isopentene. Similarly, the adsorbents used in each adsorption tower of the PSA-CH-2 unit, under their operating conditions (e.g., 100-160℃, absolute pressure 40-1000 kPa), typically have an adsorption selectivity of 6 or higher for n-pentane / isopentane in the alkane stream GL-1. The separation principle is the shape selectivity of the adsorbent for n-pentane / isopentane. The slightly smaller n-pentane molecules can enter the inner pores (larger surface area) of the adsorbent microcrystals in large quantities and be adsorbed, while the slightly larger isopentene molecules cannot enter the inner pores of the adsorbent microcrystals in large quantities and are discharged.

[0038] In the process for preparing isopentene from isopentane dehydrogenation of this invention, the isopentene separation device, in addition to the aforementioned cooling separation system and pressure swing adsorption separation system, may also include a distillation system. The distillation system includes a distillation column D1 (theoretical tray number 50-60), which separates the hydrocarbon-containing liquid stream L from the bottom of the light hydrocarbon removal column D0 to a certain extent. With the distillation system in place, the PSA-CH unit only needs to include an alkane-ene separation unit PSA-CH-1 (the PSA-CH-2 unit is not required). From the top of distillation column D1, an isopentane liquid stream Y-2 is separated, and from the bottom, an isopentene-rich liquid stream Z containing n-pentane and dienes is obtained. The obtained isopentene-rich liquid stream Z is vaporized and then passed through the PSA-CH-1 unit to separate into three parts: n-pentane stream Z-11, isopentene stream Z-21, and diene-containing heterohydrocarbon stream Z-31. Among them, isopentene stream Z-21 and its liquefied isopentene liquid stream Z-21L are the target products for isopentene dehydrogenation, and the molar purity can be controlled to above 98% or even above 99%. The n-pentane stream Z-11 and diene-containing heterohydrocarbon stream Z-31 can be further utilized or liquefied and sold.

[0039] This invention also provides a process for preparing tert-amyl alcohol, using isopentenyl liquid stream Z-2L or Z-21L from the isopentenyl separation unit in the isopentenyl dehydrogenation process and water as raw materials, mixed with recycled residual isopentenyl liquid stream X-1, recycled water stream X-3 or aqueous solution stream X-3 containing tert-amyl alcohol, and circulating solvent stream X-4. The resulting liquid stream X is subjected to a liquid-phase reaction of isopentenyl hydration to produce tert-amyl alcohol in a fixed-bed hydration reactor packed with macroporous sulfonic acid type cationic resin catalyst (hydrogen form, such as our product JQE-02). The solvent contained in the circulating solvent stream X-4 can be isopropanol, tert-butanol, or ethylene glycol monoethyl ether, preferably ethylene glycol monobutyl ether with a higher boiling point. The mixture is then passed sequentially through a light removal column (theoretical plate number 18-25) and a distillation column (theoretical plate number 18-25). (40-50 plates) The liquid stream obtained from the hydration reaction is separated to obtain residual isopentenyl liquid stream X-1, tert-amyl alcohol and water mixture stream X-20, and circulating solvent liquid stream X-4; the tert-amyl alcohol and water mixture stream X-20 is then separated by known separation methods such as azeotropic distillation, extractive distillation, or extract-azeotropic distillation to obtain tert-amyl alcohol liquid stream X-2, and water stream X-3 or an aqueous solution containing tert-amyl alcohol stream X-3; the residual isopentenyl alcohol liquid stream X-1, water stream X-3 or the aqueous solution containing tert-amyl alcohol stream X-3, and circulating solvent liquid stream X-4 are all recycled for the hydration reaction; the operating conditions of the hydration reactor include: the molar ratio of water, isopentenyl alcohol, and ethylene glycol monobutyl ether in the liquid stream X is (1-3):1:(3-8), and the liquid hourly space velocity is 0.3-1.2 h⁻¹. -1The feed temperature is 40-65℃, the pressure is 0.5-1.0MPa, and the discharge temperature is 65-80℃. The hydration reactor is preferably equipped with an external circulation system, where 40-70% of the volumetric or mass flow rate of the effluent is reused as circulating material flow X-5 and mixed with flow X before re-entering the hydration reactor. Alternatively, heat transfer and temperature control components, such as coils internally circulated with 60-65℃ circulating water, can be installed in the catalyst bed. A portion of the remaining isopentenyl liquid flow X-1 (containing water) can be sent back to the isopentenyl fractionation reactor. The cooling and separation system of the device is combined with the hydrocarbon-containing liquid stream L0 for processing and utilization, so as to maintain the total content (molar content) of 2-methyl-2-butene and 2-methyl-1-butene in the remaining isopentenene liquid stream X-1 at above 90%, that is, to control the content of 3-methyl-1-butene, alkanes and other components that do not react with water, brought in by the isopentenene liquid streams Z-2L or Z-21L, to below 10%. The extractant used in the extractive distillation method or extractive-azeotropic distillation method can be cyclohexane or n-hexane. The hydration reactor can achieve and guarantee the reaction effect of isopentenene conversion rate ≥40% and tert-amyl alcohol selectivity ≥98% during long-term operation for more than one year. In the subsequent separation process of the hydration reaction, the separation yield of tert-amyl alcohol can be guaranteed to be above 98%. The tert-amyl alcohol content in the obtained tert-amyl alcohol liquid stream X-2 exceeds 99.5 wt% or even 99.8 wt%, and the water content is less than 0.1 wt%. After cooling, it can be packaged, stored and transported as tert-amyl alcohol product. In the aforementioned tert-amyl alcohol preparation process, 2-methyl-2-butene and 2-methyl-1-butene, two of the three isomers of isopentenene, react with water to form tert-amyl alcohol, while 3-methyl-1-butene does not react with water. Using ethylene glycol monobutyl ether as the solvent for the isopentenene hydration reaction has advantages beyond stability and good reaction performance. Furthermore, its high boiling point allows the heavy components (mainly isopentenene polymers and fusel alcohols) generated during the hydration reaction and separation process to more easily enter the circulating solvent stream X-4. Periodic fractionation and purification of the circulating solvent stream X-4 helps maintain and extend the reaction performance and lifespan of the macroporous sulfonic acid-type cationic resin catalyst in the hydration reactor, and is beneficial for obtaining tert-amyl alcohol with a purity of over 99.8 wt%. The atmospheric boiling point of tert-amyl alcohol is 102 °C, and that of ethylene glycol monobutyl ether is 171 °C. The aforementioned isopentenene hydration reaction process, the separation process of the hydrated product, and the process control and effects are all known in the art.The separation method for the mixed liquid stream X-20 of tert-amyl alcohol and water utilizes the fact that tert-amyl alcohol and water are azeotropic (at normal pressure, the azeotropic temperature is 87.35℃, and the azeotropic composition is 72.5wt% tert-amyl alcohol and 27.5wt% water), and that tert-amyl alcohol and water are immiscible at temperatures such as 50-75℃, easily separating into an upper tert-amyl alcohol phase containing water and a lower water phase containing tert-amyl alcohol. For details, please refer to the relevant records and studies in "Refining Tert-Amyl Alcohol by Azeotropic Distillation" (Gao Yuanming, [J] Sichuan Chemical Industry, 1995, 4, 29-31) and "Optimization Study of Process Flow of Isoprene and Tert-Amyl Alcohol" (Wang Shaoheng, Master's Thesis of Dalian University of Technology, 2024).

[0040] The beneficial effects of the present invention include:

[0041] 1. The process for preparing isopentene by dehydrogenation of isopentene, wherein the isopentene dehydrogenation unit, when operating under the settings and conditions of modes A, B, and C, can achieve long-term operation results with a single-pass conversion rate of isopentene of 30% or even 35%, isopentene selectivity of 90% or more, and catalyst regeneration cycle of 300 hours or even 600 hours or more; the isopentene separation unit can obtain isopentene with a molar purity of 98% or even 99% or more, and hydrogen.

[0042] 2. An effective process for preparing tert-amyl alcohol from isopentane is provided, offering a potentially better option for the application of isopentane; tert-amyl alcohol is a widely used organic raw material and solvent with a high market price.

[0043] 3. The process for preparing isopentene by dehydrogenation of isopentane and the process for preparing tert-amyl alcohol from isopentane in this invention are superior to the prior art and have certain application prospects.

[0044] Comprehensive analysis suggests that the main principles underlying the process for preparing isopentene from isopentane dehydrogenation in this invention, which achieves the aforementioned beneficial effects, include:

[0045] 4. Although the catalyst H is prepared based on the proportions and method of CN103212411A, it has a specific surface area of ​​100-130 m². 2Using spherical or strip-shaped alumina as a support, zirconium oxychloride aqueous solution is impregnated onto this alumina support to prepare a zirconium-containing alumina support, ultimately producing catalyst H with better isopentane dehydrogenation performance. In method CN103212411A, a zirconium-containing alumina support is prepared by impregnating aquabsite (a crystalline form of aluminum hydroxide without formed alumina grains and internal pores) as an alumina precursor with zirconium oxychloride aqueous solution. The resulting zirconium-containing alumina support and the catalyst prepared by further loading Pt, Sn, and K show significant differences in the dispersion state of Zr and its binding state with Al2O3 compared to catalyst H. This is likely the main reason for the performance difference between the two in the isopentane dehydrogenation reaction. The specific surface area is 100-130 m². 2 The spherical or strip-shaped alumina support, typically calcined at temperatures above 950℃, exhibits low surface activity. During the catalyst preparation step (H-1), involving impregnation with a strongly acidic zirconium oxychloride aqueous solution, drying, and calcination at 880-930℃ to prepare the zirconium-modified alumina support, the alumina support itself is eroded by the reaction with zirconium oxychloride (zirconia hydrolyzes to release HCl, which reacts with the alumina). However, this erosion mostly occurs on the inner pore surface of the alumina grains, resulting in a relatively uniform microscopic distribution of the final zirconium-alumina composite oxide coating. In the CN103212411A method, the zirconium oxychloride in the aqueous solution reacts fully with the surface layer of the boehmite grains. In the dried and calcined zirconium-containing alumina support, Zr is mainly distributed on the outer surface and between the alumina grains, with less distribution in the inner pores of the alumina grains. Consequently, the final zirconium-alumina composite oxide coating exhibits a non-uniform microscopic distribution. In the catalyst preparation method disclosed in Liu Changcheng's article, due to the same reason, most of the Sn in the Sn-Al2O3 support is distributed on the outer surface or between alumina grains (with less distribution on the inner pore surface of alumina grains), while most of the further loaded Pt is distributed on the inner pore surface of alumina grains. This results in an unsatisfactory dispersion and bonding state between Pt and Sn in the catalyst. Even after hydrogen reduction, the dispersion and bonding state between Pt and Sn does not reach the level of catalysts H and J of this invention, making it difficult to achieve a regeneration cycle of more than 300 hours. CN107537485A does not disclose the size and shape of the spinel-type support and catalyst it prepared, but its ratio inevitably leads to a contradiction between mechanical strength and inner pore volume, making it difficult to apply this catalyst in the isopentane dehydrogenation process in a fixed-bed reactor. The main reason is that it requires molding by means of powder pressing to obtain the mechanical strength required for a fixed-bed catalyst, thus limiting the inner pore volume and the ability to accommodate carbon deposits, ultimately making it difficult to achieve a regeneration cycle of more than 300 hours.

[0046] 5. In the preparation of the magnesium-containing catalyst J, during the process of impregnating the magnesium bicarbonate aqueous solution, drying, and calcining at 880-930℃ to prepare the magnesium-modified alumina support, the magnesium bicarbonate aqueous solution and the magnesium carbonate generated by heating to 70-75℃ under flowing air conditions and isothermal treatment to release carbon dioxide, drying, and calcining at 880-930℃ to prepare the magnesium (MgO)-modified alumina support, the resulting magnesium-aluminum composite oxide coating is microscopically uniformly distributed; during the process of impregnating the magnesium-modified alumina support with Pt-Sn-EDTA-K aqueous solution (pH value close to neutral), drying, and calcining at 500-600℃ to prepare the alkane dehydrogenation catalyst, the support skeleton is basically unaffected, the Pt distribution is relatively uniform, and the dispersion degree is high. In step (A) of preparing the Pt-Sn-EDTA-K aqueous solution described in step (J-2), potassium stannate, potassium hydroxide, and potassium nitrite together provide the K loading of the catalyst; the reaction of the potassium stannate aqueous solution with EDTA and nitric acid until the solution becomes clear indicates that EDTA has loaded Sn. 4+ After complete complexation, take one drop of solution and add 2-4 drops of 40wt% hexamethylenetetramine aqueous solution. Add one drop of indicator solution, such as 0.02wt% xylenol orange aqueous solution, to the mixture (pH around 5.5); it should turn yellow. When potassium stannate is in molar excess compared to EDTA, adding potassium hydroxide to adjust the pH to 6.0-6.5 in the aqueous solution will produce flocculent matter, and the solution will not be clear. The Sn-EDTA-K aqueous solution obtained after adding potassium nitrite (pH around 6.5) will also not be clear. Using potassium nitrite makes the near-neutral Pt-Sn-EDTA-K aqueous solution more stable during storage and carrier impregnation, mainly by preventing the decomposition of diammonium nitrite, which is not possible when the solution contains only potassium nitrate. In the preparation of the magnesium-containing spherical or strip-shaped catalyst J, diammonium nitrite, which has low solubility in aqueous solution but whose concentration meets the catalyst loading requirements, and Sn are used. 4+ The solution contains an EDTA complex, and the pH is near neutral. Therefore, the molecular structure of diammonium platinum nitrite can be considered intact (in reality, it partially ionizes to release [Pt(NH3)2]). 2+ Ion, [Pt(NH3)2] 2+ (The ionic structure of [Pt(NH3)2] remains intact) 2+ The adsorption on the inner and outer surfaces of magnesium-modified alumina supports is weak, and uniform distribution of Pt in the catalyst can be achieved without competitive adsorbents such as citric acid; if the Pt-Sn-EDTA-K aqueous solution contains a certain concentration of Cl... - (Chloride ions) then form PtCl4 2- The adsorption of Pt is strong on the inner and outer surfaces of the magnesium-modified alumina support, leading to uneven distribution of Pt in the absence of suitable competing adsorbents. Meanwhile, after drying in step (J-2), Sn... 4+The fine crystals and uniform dispersion of EDTA complex and diammonium nitrite in the pores of magnesium-modified alumina support result in a uniform radial distribution of the Pt-Sn component in the final catalyst, rather than an eggshell-shaped distribution. This is evident from the uniform color caused by the Pt component. Furthermore, the dispersion of Pt and Sn is higher than that of existing catalysts, including the zirconium-containing catalyst H, or the dispersion and bonding of Pt and Sn are better. This gives catalyst J not only better overall isopentane dehydrogenation performance but also CO2 hydromethanization performance that matches its isopentane dehydrogenation performance. The diammonium nitrite and potassium stannate used are readily available and of stable quality; potassium stannate is a tetravalent tin salt, eliminating the need for the oxidation of divalent tin to tetravalent tin.

[0047] 6. In the preparation process of the magnesium-containing catalyst J, the magnesium-modified alumina support is used, which basically eliminates the surface acidity of the alumina support. Under the alkaline conditions of the K loading, the catalyst has low isomerization activity from isopentane to n-pentane, and extremely low activities for cracking, carbonization, and deep dehydrogenation to generate dienes. Catalyst J exhibits good reactivity in a fixed-bed adiabatic reactor and under hydrogen-containing conditions, such as under the setup and operation conditions of dehydrogenation reactor group modes A and B, for the dehydrogenation of isopentane to isopentene and the matching hydrogenation of CO2 to methane. Under the setup and operation conditions of dehydrogenation reactor group mode C (without CO2), it also exhibits good reactivity in the dehydrogenation of isopentane to isopentene, and the overall catalytic activity of the reactor is relatively stable with slow deactivation. Under the setup and operation conditions of reactor group modes A and B, long-term operation effects of isopentane single-pass conversion of over 30% or even over 35%, isopentene selectivity of over 90%, and catalyst regeneration cycle of over 400 hours can be achieved. Under the setup and operation conditions of reactor group mode C, long-term operation effects of isopentane single-pass conversion of over 30%, isopentene selectivity of over 90%, and catalyst regeneration cycle of over 600 hours can also be achieved.

[0048] 7. Under the setup and operating conditions of modes A and B, the dehydrogenation reactor group of the isopentane dehydrogenation unit, when using either the zirconium-containing catalyst H or the magnesium-containing catalyst J, ultimately achieves a single-pass isopentane conversion rate of over 30% or even over 35%, isopentene selectivity of over 90%, CO2 methanation rate of over 80%, and a catalyst regeneration cycle of over 300 hours for long-term operation. This indicates that both catalysts possess CO2 hydromethanation performance that matches the isopentane dehydrogenation performance, effectively coupling the two reactions. This utilizes the exothermic nature of the CO2 hydromethanation reaction to significantly improve the isopentane dehydrogenation conversion rate. Since isopentane dehydrogenation is a significantly endothermic reaction, and CO2 hydromethanation is a strongly exothermic reaction, if the catalyst's CO2 hydromethanation performance is significantly dominant, the maximum temperature of the catalyst bed can easily exceed 560℃, leading to a decrease in isopentene selectivity and a shortened catalyst regeneration cycle. During the dehydrogenation operation, the temperature and / or CO2 content of the inlet gas flow in each reactor can be appropriately adjusted according to the overall performance of the reactor. For example, at the beginning of the operating cycle when the overall catalyst performance is high, a lower temperature and / or CO2 content of the inlet gas flow can be used. At the end of the operating cycle when the overall catalyst performance declines, the temperature and / or CO2 content of the inlet gas flow can be appropriately increased to obtain a relatively stable isopentene yield and production. Due to the use of a relatively low gas flow inlet temperature, the catalyst bed temperature is also lower than that of existing technologies, thus reducing side reactions and extending the regeneration cycle. The methanation rate of CO2 can be estimated from the CO2 conversion rate and the distribution and content of the generated oxygen-containing compounds. In the operation of modes A and B, the CO2 conversion rate easily exceeds 90% or even 95%, but the outlet gas of the dehydrogenation reactor does not contain methanol.

[0049] 8. During stable operation of each isopentane dehydrogenation reactor, the maximum temperature of the catalyst bed is controlled below 550℃ under the setting and operating conditions of modes A and B, and below 560℃ for reactor III and below 550℃ for reactors IV and V under the setting and operating conditions of mode C. This is a key factor in obtaining isopentene selectivity of over 90% and catalyst regeneration cycles of over 300 hours or even over 500 hours. In mode A reactors, as well as in modes B and C, the outlet gas flow temperature or catalyst bed outlet temperature of the downstream reactors is controlled above 530℃, which is a key factor in obtaining isopentane conversion of over 30%. The concentration of the main dehydrogenation product, 2-methyl-2-butene, is close to the equilibrium concentration between isopentane and the outlet product under specific reaction conditions such as outlet temperature and composition. Attached Figure Description

[0050] Appendix Figure 1 The following is a schematic diagram of the process flow for mode B in the process of preparing isopentene from isopentane by dehydrogenation in Example 1; wherein, each PSA unit only shows a summary of the material inflow and outflow relationship. Detailed Implementation

[0051] The technical solution of the present invention will be specifically described and explained below with reference to the embodiments.

[0052] Example 1

[0053] 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, and 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, the process for preparing isopentene from isopentane dehydrogenation in this embodiment was designed, including an isopentane dehydrogenation unit and an isopentene separation unit;

[0054] The isopentane dehydrogenation unit includes a dehydrogenation reactor group, as well as corresponding connecting pipelines, temperature measurement and control components, flow measurement and control components, and valves; the reactor of the dehydrogenation reactor group is a fixed-bed adiabatic reactor with an axial or radial reaction structure, and is filled with catalyst HA or catalyst JB.

[0055] The setup and operation of dehydrogenation reactor units include any of the following modes:

[0056] Mode A: Three reactors are set up, each with the same catalyst loading volume. 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 for each reactor include: temperature 520-550℃, pressure ≤0.2MPa, and isopentane mass hourly space velocity (MHV) 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 section temperature of the catalyst bed in each reactor is controlled above 530℃.

[0057] Mode B: Two parallel reactor series are set up. Each reactor series includes two reactors, I and II, connected in series sequentially, with reactor I being the first reactor. The catalyst loading volume ratio in reactors I and II is 1:1. A heater is installed before the gas flow inlet of each reactor. During the stable operation of isopentane dehydrogenation, the inlet gas flow conditions for reactor I are: 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 of reactor II include: temperature 510-540℃, and molar content of CO2 1-2%. The maximum temperature of the catalyst bed in each reactor is controlled below 550℃. The outlet gas flow temperature of reactor II or the outlet section temperature of the catalyst bed is controlled above 530℃.

[0058] Mode C: Two parallel reactor series are set up. Each reactor series includes three reactors (III, IV, and V) connected in series sequentially, with reactor III being the first reactor. The catalyst loading volume ratio in reactors III, IV, and V is 1:1:1. A heater is installed before the gas flow inlet of each reactor. During the stable operation of isopentane dehydrogenation, the inlet gas flow conditions for reactor III are: temperature 550-560℃, pressure ≤0.2MPa, and isopentane mass hourly space velocity (MHV) 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℃.

[0059] Specifically, when a reactor in mode A or a reactor series in modes B and C experiences a drop in its reaction performance index to a specified range, the dehydrogenation operation is switched off and the catalyst is regenerated in situ. After regeneration, the dehydrogenation operation is switched back on. The isopentane dehydrogenation unit is also equipped with a heat exchanger T1. The mixture of the raw material isopentane liquid flow, the recycled isopentane gas flow Y, the recycled hydrogen gas flow G-H2, and the carbon dioxide gas flow M, and the reaction gas flow P from the dehydrogenation reactor group, undergo countercurrent heat exchange through the heat exchanger T1 to obtain the pre-reaction gas flow N and the cooled gas flow Q. The pre-reaction gas flow N enters the dehydrogenation reactor group, and the cooled gas flow Q goes to the isopentene separation unit.

[0060] The isopentene separation unit includes a cooling separation system and a pressure swing adsorption (PSA) separation system. The cooling separation system first further cools and pressurizes the isopentene dehydrogenated gas stream Q, which has been cooled by heat exchanger T1, to (-20 to -10)℃ and pressurizes it to 0.5-1 MPa. This process separates water and yields a gas stream G0 mainly containing H2, CH4, and small amounts of CO2, ethane, and ethylene, as well as a hydrocarbon-containing liquid stream L0. The hydrocarbon-containing liquid stream L0 enters a light-light removal tower D0 (a packed tower with approximately 10 theoretical trays). A small amount of low-boiling gas stream G-1 containing C4 or lower is separated from the top of the tower, and the hydrocarbon-containing liquid stream L is separated from the bottom. The pressure swing adsorption unit includes a PSA-H2 unit for separating hydrogen from gas stream G0, and one or more PSA-CH units for separating hydrocarbons. The PSA-H2 unit uses conventional pressure swing adsorption adsorbents, equipment configuration, and separation operation processes, while the PSA-CH unit uses known hydrocarbon separation adsorbents, equipment configuration, and pressure swing adsorption separation operation processes. Gas flow G0 is separated in the PSA-H2 unit to obtain hydrogen gas G-H2 with a molar purity of 98% or even 99% or higher, and almost free of unsaturated hydrocarbons. Part of this hydrogen gas is recycled to the isopentane dehydrogenation unit as a feedstock for the dehydrogenation reaction, and the remainder is supplied externally. The low-hydrogen gas flow G-2, containing CH4, ethane, ethylene, and CO2, obtained from the PSA-H2 unit is used as fuel gas in the heater. The PSA-CH unit has two hydrocarbon separation units, including the alkane separation unit PSA-CH-1 and the alkane separation unit PSA-CH-2. The PSA-CH-1 unit separates the vaporized hydrocarbon-containing liquid flow L into three parts: alkane gas flow GL-1, isopentene gas flow Z-2, and mixed hydrocarbon gas flow Z-3, all of which have the required purity. The isopentene gas flow Z-2 can be controlled to have a molar purity of 98% or even 99% or higher. Dienes are basically sent to the mixed hydrocarbon gas flow Z-3, and the amount of mixed hydrocarbon gas flow Z-3 is very small, which can be used as fuel gas or liquid gas in the heater. After collection, the alkane stream GL-1 is sold externally. The alkane stream GL-1 is further separated in the PSA-CH-2 unit to obtain two streams: isopentane stream Y-1 and n-pentane stream Z-1. The molar purity of isopentane in isopentane stream Y-1 and the molar purity of n-pentane in n-pentane stream Z-1 can be easily controlled to above 96% or even above 98%. Isopentane stream Y-1 is recycled in the isopentane dehydrogenation unit as part of the feedstock for the dehydrogenation reaction. n-pentane stream Z-1 is liquefied, collected, and supplied externally, or converted to isopentane in the isopentane dehydrogenation unit and then used in the isopentane dehydrogenation unit. The resulting isopentene stream Z-2 is further liquefied to obtain isopentene liquid stream Z-2L. A small amount of low-boiling stream G-1 containing C4 and below components is separated from the top of the light-light content removal tower D0 and also used as fuel in the heater. The configuration and operating performance of the isopentene separation unit were simulated using Aspen Plus software.

[0061] The process of the cooling and separation system may include: cooling gas flow Q to 20-60℃ (via a heat exchanger for cooling via 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 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, and condensate is separated when cooling to 1-10℃ (in the setup and operation modes A and B of the dehydrogenation reactor group, this is the water produced by the CO2 hydromethanation reaction), During the cooling process from -20 to -10℃, multiple cooling heat exchangers are used and de-icing operations are performed in turn 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 stream. In mode C, since a large amount of water vapor is introduced into the gas flow at the inlet of reactor III, the gas flow Q is continuously separated into condensate by an 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 a gas-liquid separator into a gas flow G0 mainly containing H2, CH4 and a small amount of CO2, ethane, and ethylene, and a hydrocarbon-containing liquid flow L0.

[0062] The alkane adsorbents used in each adsorption tower of the PSA-CH-1 unit exhibit an adsorption selectivity of 3.3 for isopentene / pentane in the vaporized stream of hydrocarbon-containing liquid flow L, under operating conditions of 40-100℃ and 30-1000 kPa absolute pressure. The pentane refers to the total amount of n-pentane and isopentene. The alkane adsorbents used in each adsorption tower of the PSA-CH-2 unit exhibit an adsorption selectivity of 6.5 for n-pentane / isopentane in the alkane stream GL-1, under operating conditions of 100-160℃ and 40-1000 kPa absolute pressure. Both the PSA-CH-1 and PSA-CH-2 units employ a three-tower adsorption process. The operating sequence of each adsorption tower includes high-pressure adsorption, displacement, vacuum desorption, and pressurization processes, all cyclically performed. 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.

[0063] Each reactor or series requires reduction after being loaded with new catalyst and before being put into dehydrogenation operation. The final step of the regeneration process is also a reduction operation. The reduction operation conditions are: using a hydrogen gas flow with H2 molar content higher than 98% and olefin molar content lower than 0.1%, and a space velocity of 500-1000 h⁻¹. -1 The pressure is 0.02-0.2 MPa, the initial bed temperature is 480-500℃, and the maximum bed temperature does not exceed 520℃. The process continues for 2-4 hours until the water vapor molar content in the outlet gas flow is reduced to below 0.1%.

[0064] The in-situ regeneration process performed on each reactor or series after the isopentane dehydrogenation operation has been stopped when the overall reaction performance indicators have decreased to the specified range can be carried out under chlorine conditions. Chlorine regeneration methods include carbonization, oxychlorination, and reduction operations, specifically following the first in-situ regeneration method described in Catalyst Evaluation Example 1. Specifically, during the carbonization operation, an oxygen-containing air stream with an O2 molar content of 0.5-0.8% and a nitrogen flow (N2 molar content above 99%) is used, with a space velocity of 500-2000 h⁻¹. -1 The inlet temperature is 450-490℃, 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. -1 The pressure is 0.02-0.2 MPa, and the initial bed temperature is 480-520℃. First, water or steam, along with chlorides such as dichloroethane, trichloroethane, carbon tetrachloride, or tetrachloroethylene, are introduced into the oxygen-containing stream. The inlet temperature of the gas stream is controlled at 500-520℃, with the amount of chloride added being 0.4-0.6 wt% of the catalyst in the reactor (calculated as chlorine). The water-to-chlorine molar ratio is controlled at 5-8, and the introduction time is 3-6 hours. Then, the inlet temperature of the oxygen-containing stream (without further water or chloride introduction) is controlled at 520-540℃, and treatment is carried out for 4-6 hours, with the reactor outlet gas temperature above 515℃. After the oxychlorination operation is completed, the reactor is purged with a high-purity nitrogen gas stream (N2 molar content above 99.9%) at a temperature of 480-500℃ and 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 under the aforementioned reduction operating conditions. After one and multiple in-situ regenerations including the oxychlorination operation, the coke deposits of catalyst HA or catalyst JB are essentially burned off, and the performance of CO2 hydromethanization and isopentane dehydrogenation is essentially restored. Catalyst HA can essentially recover to the performance level of the first operating cycle, while the overall performance of catalyst JB after one and multiple chlorine-containing regenerations is lower than the performance level of the first operating cycle without chlorine but higher than that of catalyst HA.

[0065] Reactors or series loaded with catalyst JB can undergo in-situ regeneration under chlorine-free conditions, including carbonization, oxidation, and reduction operations. Specifically, the carbonization operation uses an oxygen-containing airflow with an O2 molar content of 0.5-0.8%, obtained by combining a nitrogen stream with a dry air stream (e.g., dew point below -20°C), at a space velocity of 500-2000 h⁻¹. -1The inlet temperature is 450-490℃, the pressure is 0-0.1MPa, and the initial bed temperature is 450-500℃. During the charcoal burning process, the maximum bed temperature is controlled to not exceed 520℃. The process continues for 2-4 hours until the difference in O2 molar content between the outlet and inlet air streams decreases to below 0.1%. For oxidation operations, an oxygen-containing air stream with an O2 molar content of 10-20%, obtained by mixing nitrogen with dry air, or a dry air stream can be used directly, with a space velocity of 500-2000 h⁻¹. -1 The pressure is 0.02-0.2 MPa, the initial bed temperature is 480-520℃, the gas inlet temperature is 500-520℃, and the oxidation treatment lasts for 6-10 hours. After the oxidation operation is completed, the reactor is purged with nitrogen gas at 480-500℃ and high-purity nitrogen gas, 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%, the catalyst is reduced again under the aforementioned reduction operating conditions. Reactors or series loaded with catalyst JB, after undergoing in-situ regeneration under the aforementioned chlorine-free conditions, can achieve long-term operating results of over 30% or even 35% single-pass isopentane conversion, over 90% isopentene selectivity, over 80% methanation rate in the presence of CO2, and over 300 hours or even 450 hours of catalyst regeneration cycle when performing isopentane dehydrogenation under the settings and operating conditions of the aforementioned dehydrogenation reactor group modes A, B, and C. In particular, the isomerization side reaction problem of n-pentane formation is less severe compared to catalyst JB after in-situ regeneration including oxychlorination operations. After in-situ regeneration under the aforementioned chlorine-free conditions, the coke deposits of catalyst JB are basically burned off, and the isopentane dehydrogenation reaction performance and CO2 hydromethanization performance are largely restored (slightly lower than the effect level in the first operating cycle).

[0066] Example 2

[0067] This embodiment designs another process for preparing isopentene by dehydrogenation of isopentane, which is basically the same as the process in Example 1. The difference is that the isopentene separation device includes not only the cooling separation system and the pressure swing adsorption separation system, but also a distillation system. The PSA-CH device in the pressure swing adsorption unit only has an alkene separation unit PSA-CH-1 (without a PSA-CH-2 unit). The distillation unit includes a distillation column D1 (theoretical plate number 55), which separates the hydrocarbon-containing liquid stream L to a certain extent. The specific operation process includes: the hydrocarbon-containing liquid stream L enters the light removal column D1, the isopentane liquid stream Y-2 is separated from the top of the column, and the isopentene-rich liquid stream Z containing n-pentane and diene is obtained from the bottom of the column. The isopentene-rich liquid stream Z is vaporized and then passed through PSA-CH-1 to separate into three parts: a n-pentane stream Z-11, an isopentene stream Z-21, and a diene-containing hydrocarbon stream Z-31. The isopentene stream Z-21 and its liquefied isopentene liquid stream Z-21L are the target products for isopentene dehydrogenation in this embodiment, with a molar purity of over 98% or even 99%. The n-pentane stream Z-11 and the diene-containing hydrocarbon stream Z-31 can be further utilized or liquefied for sale. This isopentene separation device achieves essentially the same overall effect as in Example 1.

[0068] In Examples 1 and 2, the processes for preparing isopentene by dehydrogenation of isopentene, when the isopentene dehydrogenation unit is operated under the settings and conditions of modes A, B, and C, can achieve long-term operation results with a single-pass conversion rate of isopentene of 30% or even 35%, isopentene selectivity of 90% or more, and catalyst regeneration cycle of 300 hours or more. The isopentene separation unit can obtain isopentene with a molar purity of 98% or even 99% or more, and hydrogen. The service life of catalyst HA and dehydrogenation catalyst JB should be more than 3 years.

[0069] Example 3

[0070] This embodiment prepared a Pt-Sn-K / alumina type alkane dehydrogenation catalyst HA (2.1L), which is spherical with an average outer diameter of Φ2.8mm. It is a zirconium-containing spherical catalyst prepared by improving the proportions and methods based on CN103212411A: using a high-temperature zirconium-aluminum composite oxide containing 1.1wt% ZrO2 as a support, and the supported components are Pt 0.35wt%, Sn 0.7wt%, and K 0.83wt% (calculated values ​​in the catalyst). It was prepared through the following steps:

[0071] (H-1) with a pore volume of 0.62 mL / g and a specific surface area of ​​118 m² 2Using spherical alumina (1200g, Φ2.8mm, pre-baked at 400℃ for 3h) as a carrier (room temperature, 1200g, tapped volume 2310mL, pore volume and specific surface area measured by low-temperature N2 adsorption method), a small, sealable drum (12L capacity) was uniformly sprayed with 74.5mL of a 1.453mol / L zirconium oxychloride aqueous solution. The spray volume was 90% of the carrier's saturated water absorption volume (69.1mL / 100g), and the spraying time was 2 hours. After 5 minutes, the drum continued to rotate for 20 minutes until the surface of the small balls was completely dry. The material was then discharged, bagged, and placed in a sealed container for homogenization treatment for 15 hours. It was then placed on a tray and dried in an oven at 130°C for 8 hours. After that, it was calcined in a muffle furnace at 900°C for 3 hours. When the temperature was lowered to 350°C, the material was discharged and placed into a 5L stainless steel container that had been pre-purged with nitrogen. The container was sealed and nitrogen was purged to maintain positive pressure inside (to prevent moisture absorption). The container was then placed and cooled to room temperature to obtain a zirconium-modified alumina carrier with a calculated ZrO2 content of 1.1 wt%.

[0072] (H-2) Take 1100g of the zirconium-modified alumina support obtained in step (H-1) and place it in a rotating drum. Spray 626mL of Pt-Sn-acid-K aqueous solution (temperature 26℃) evenly. The amount of sprayed solution is 85% of the saturated water absorption volume of the zirconium-modified alumina support (67.1L / 100g). The spraying time is 25min. The drum continues to rotate for 28min until the surface of the small ball is completely dry. The material is discharged, bagged, sealed and placed for homogenization treatment for 12h. It is then placed on a tray and dried in an oven at 130℃ for 8h. After that, it is calcined in a muffle furnace at 550℃ for 3h. After cooling, the material is discharged to obtain the alkane dehydrogenation catalyst HA with a tapped volume of 2110mL.

[0073] The Pt-Sn-acid-K aqueous solution described in step (H-2) is prepared by dissolving chloroplatinic acid, stannous chloride, hydrochloric acid, citric acid, and potassium nitrate into an aqueous solution of the required concentration. The preparation method is as follows: dissolve the required amount of chloroplatinic acid and hydrochloric acid in an appropriate amount of water, add the required amount of stannous chloride, potassium nitrate, and citric acid and stir until dissolved, add water to the required total volume and stir well to obtain 630 mL of Pt-Sn-acid-K aqueous solution; wherein the concentration of chloroplatinic acid is 0.0327 mol / L, the concentration of stannous chloride is 0.107 mol / L, the concentration of potassium nitrate is 0.387 mol / L, the concentration of HCl introduced by hydrochloric acid is 0.075 mol / L, the concentration of citric acid is 0.3 mol / L (the concentrations of each element or substance are calculated values), and the pH value is approximately 1.

[0074] The above preparation process was repeated to obtain 4.2 L of alkane dehydrogenation catalyst HA. The obtained alkane dehydrogenation catalyst HA has a uniform appearance and cross-sectional color, indicating that Pt is evenly distributed in the catalyst particles.

[0075] Example 4

[0076] This embodiment prepared a Pt-Sn-K / alumina type alkane dehydrogenation catalyst JB (2.1L), which is a sphere with an average outer diameter of Φ2.8mm. It used magnesium-modified alumina containing 1.5wt% MgO as the support, and the supported components were Pt 0.3wt%, Sn 0.7wt%, and K 0.83wt% (calculated content in the catalyst), and were prepared through the following steps:

[0077] (J-1) The pore volume used in step (H-1) of preparing the zirconium-containing spherical catalyst HA in Example 3 was 0.62 mL / g, and the specific surface area was 118 m². 2 Using spherical alumina (1200g, 2310mL tapped volume) with a diameter of Φ2.8mm and pre-dried at 400℃ for 3h as a carrier (room temperature), the alumina was placed in a rotating drum (the same as used in step H-1 of Example 2). 497mL of a 1.453mol / L magnesium bicarbonate aqueous solution was uniformly sprayed onto the alumina, the spray volume being 60% of the carrier's saturated water absorption volume (69.1mL / 100g). The spraying time was 25min, and the drum continued to rotate for 6min until the surface of the spheres was completely dry. The alumina was then bagged, sealed, and homogenized for 10h. It was then placed on a tray and heated to 73℃ in an oven with flowing air, and kept at that temperature for 4h. It was then heated to 130℃ and dried for 8h. Afterward, it was calcined in a muffle furnace at 900℃ for 3h, and then cooled to 350℃ before being discharged into a stainless steel container (the same as in Example 2). The same process as in step H-1 (pre-purged with nitrogen) is followed by sealing and maintaining positive pressure inside the container with nitrogen (to prevent moisture absorption). The container is then placed and cooled to room temperature to obtain a magnesium-modified alumina support with a calculated MgO content of 1.5 wt%.

[0078] (J-2) Take 1100g of the magnesium-modified alumina support obtained in step (J-1) and place it in a rotating drum. Spray 626mL of Pt-Sn-EDTA-K aqueous solution (temperature 57℃) evenly. The spray volume is 85% of the saturated water absorption volume of the magnesium-modified alumina support (67.0mL / 100g). The spraying time is 25min. The drum continues to rotate for 22min until the surface of the small ball is completely dry. The material is discharged, bagged, sealed and placed for homogenization treatment for 12h. It is then placed on a tray and dried in an oven at 130℃ for 8h. After that, it is calcined in a muffle furnace at 550℃ for 3h. After cooling, the material is discharged to obtain alkane dehydrogenation catalyst JB with a tapped volume of 2120mL.

[0079] The Pt-Sn-EDTA-K aqueous solution described in step (J-2) is prepared by mixing diammonium nitrite (Pt(NH3)2(NO2)2), potassium stannate (K2SnO3), EDTA (ethylenediaminetetraacetic acid), nitric acid, potassium hydroxide, and potassium nitrite to form an aqueous solution of the required concentration. The preparation method is as follows: (A) Add 20.2g of potassium stannate trihydrate to 300mL of potassium hydroxide aqueous solution with a pH of 11, stir until dissolved, add 20.3g of EDTA powder and 46.1mL of nitric acid aqueous solution (concentration 3.22mol / L), heat to about 60℃ and stir for 30min until clear, add 15.8mL of potassium hydroxide aqueous solution (concentration 4.6mol / L) to adjust the pH to 6.3, add 3.1g of potassium nitrite and stir until dissolved to make up the total amount of K, stir well to obtain the Sn-EDTA-K aqueous solution; wherein the molar ratio of K2SnO3, EDTA, and HNO3 is 1:1.03:2.2; (B) In step (A), the Sn-EDTA-K aqueous solution was kept at 60℃. 234.6 mL of diammonium nitrite aqueous solution (Pt concentration of 0.0752 mol / L) at 60℃ was added, and water was added to a total volume of 630 mL and stirred well to obtain a Pt-Sn-EDTA-K aqueous solution. The Pt concentration was 0.028 mol / L, the Sn concentration was 0.107 mol / L, the K concentration was 0.387 mol / L, and the EDTA concentration was 0.11 mol / L (prepared as needed; the concentrations of each element or substance are calculated values). The pH value was approximately 6.5.

[0080] The above preparation process was repeated to obtain 4.2 L of alkane dehydrogenation catalyst JB. The obtained alkane dehydrogenation catalyst JB has a uniform appearance and cross-sectional color, indicating that Pt is evenly distributed in the catalyst particles.

[0081] Catalyst Evaluation Example 1

[0082] In a specially designed dehydrogenation pilot plant, the alkane dehydrogenation catalyst HA prepared in Example 2 was subjected to performance evaluation tests for the dehydrogenation of isopentane to isopentene under fixed-bed and quasi-adiabatic reactor conditions. This dehydrogenation pilot plant was located at the production site of an isobutane dehydrogenation unit, and its connection method B had been used for a long time during the development of the isobutane dehydrogenation to isobutene process described in CN103232312A.

[0083] The main equipment of the dehydrogenation pilot plant includes three vertical fixed-bed adiabatic reactors I, II and III with axial reaction structures, as well as heaters I, II and III, pipe mixers I and II, two high-temperature ovens A and B with internal circulation fans and a design temperature of 600°C, and two condensers A and B. Pipe mixer I, heater I and reactor I are installed in high-temperature oven A, pipe mixer II, heater II and reactor II are installed in high-temperature oven B, and when connection method C is adopted, heater III and reactor III are also installed in high-temperature oven B.

[0084] Reactors I, II, and III all have a circular plate-shaped lower head, a cylindrical filling section (65mm inner diameter, 2mm wall thickness), an upper flange, and a thermocouple sheath (the upper flange and lower head are respectively welded with an 8mm inner diameter and 1mm wall thickness air inlet and outlet; the thermocouple sheath (6mm outer diameter) is welded to the lower head with its central axis coinciding with the cylindrical filling section and extending upwards to the top of the filling section; the airflow direction within the reactor / catalyst bed is vertically downwards). They are respectively densely filled from bottom to top with 200mL of alumina ceramic balls (3mm outer diameter) and catalyst HA. 1000 mL (catalyst bed height approximately 300 mm, height-to-diameter ratio 4.6), 200 mL alumina ceramic balls (outer diameter 3 mm), and four thermocouples inserted into the thermocouple sheath to detect the gas inlet temperature and the temperatures of the upper, middle, and lower sections of the catalyst bed (three thermocouples for detecting the catalyst bed temperature, with their ends at heights of 50 mm, 150 mm, and 250 mm respectively from bottom to top within the catalyst bed). The unfilled space at the top of the reactor is approximately 220 mL. The entire outer surface of the reactor is tightly covered with 25-30 mm thick alumina fiber felt as insulation material, and the alumina fiber felt is wrapped with quartz gauze for fixation. Reactors I, II, and III all basically meet the conditions for adiabatic dehydrogenation reaction.

[0085] Heaters I, II, and III are all shell-and-tube heat exchangers without a shell side, with heat exchange areas of 1.3 m² and 1.3 m² respectively. 2 0.6m 2 0.6m 2 During the dehydrogenation operation, the temperature of the mixed raw material gas stream or intermediate reaction gas stream is adjusted to raise the temperature to the required reactor inlet temperature. During the stable operation, the temperature difference between ovens A and B and the gas inlet temperature of reactors I and II is within 2℃. Pipe mixers I and II have the same specifications, with an internal volume of 320mL. They are vertically installed, and the material flow enters from the bottom interface. Pipe mixer I also has the function of vaporizing isopentane.

[0086] Condensers A and B are shell-and-tube heat exchangers of the same specifications, with a heat exchange area of ​​1m². 2The tube side flows through the cooling medium, while the shell side volume and maximum liquid capacity are 12L and 0.6L, respectively. Condenser A is a pre-cooling heat exchanger using circulating water as the cooling medium (circulating temperature 20-30℃), and condenser B uses antifreeze (ethylene glycol-water solution) for vehicle engines with a freezing point of -40℃ as the cooling medium (circulating temperature can be adjusted within the range of -30℃ to -25℃). Condensers A and B are used in series to cool the reactor outlet gas flow to -22℃ to -20℃. The connecting pipe between condensers A and B... An adjustable back pressure valve is installed in the pipeline to adjust the system pressure; condenser B discharges uncondensed gas flow and condensate flow. The uncondensed gas flow is continuously monitored for volumetric flow rate and temperature, and samples are taken regularly to test its chemical composition (every 1-3 hours) before being sent to the flare network. The condensate flow is sampled regularly to test its chemical composition (every 2-5 hours) before flowing through a U-shaped liquid seal into a storage tank (50L, placed in a freezer with temperature control at -30℃, and weighed, emptied, or replaced regularly) for temporary storage; condenser B and the condensate flow pipeline are respectively insulated with foamed plastic sheets.

[0087] Both sides of the high-temperature ovens A and B have low-position openings for connecting pipes and thermocouples (K-type armored, 1mm outer diameter), which are blocked with alumina cotton. The gas outlet of reactor I and the gas outlet of reactor II when using connection method C are also connected to high-temperature gas flow sampling tubes. The high-temperature gas flow sampling tubes extend outside the oven through the side wall openings of the high-temperature oven and are equipped with needle valves that can withstand short-term high temperatures at the end for measuring the composition of the gas flow at the reactor outlet.

[0088] During the evaluation test of catalyst HA, based on the overall flow sequence of the reactant stream, the following three equipment configurations and connection methods (B, C, and A) were used successively:

[0089] Connection method B: The combined flow of isopentane liquid, hydrogen gas and carbon dioxide gas I → pipeline mixer I → heater I → reactor I → supplementary carbon dioxide gas II → pipeline mixer II → heater II → reactor II → condenser A → back pressure valve → condenser B; This connection method B is used to test the performance of catalyst HA in the dehydrogenation reactor group setup and operation mode B described in Example 1.

[0090] Connection method C: The combined flow of isopentane liquid, water and hydrogen gas → pipeline mixer I → heater I → reactor I → heater II → reactor II → heater III → reactor III → condenser A → back pressure valve → condenser B; This connection method C is used to test the performance of catalyst HA in the dehydrogenation reactor group setup and operation mode C described in Example 1;

[0091] Connection method A: The combined flow of isopentane liquid, hydrogen gas and carbon dioxide gas I → pipeline mixer I → heater I → reactor I → condenser A → back pressure valve → condenser B; This connection method A is used to test the performance of catalyst HA in the dehydrogenation reactor group setup and operation mode A described in Example 1;

[0092] The feed configuration for reactor I during the dehydrogenation process is as follows: a hydrogen gas stream at a certain pressure (from the hydrogen pipeline network of the isobutane dehydrogenation production unit, with a molar purity of ≥98% and a hydrocarbon content of ≤350 mg / Nm³). 3 Unsaturated hydrocarbon content is less than 100 mg / Nm 3 (Dew point below -18℃), carbon dioxide gas flow I (gasified from steel cylinder, molar purity above 99.98%, hydrocarbon content below 3mg / Nm³). 3 The flow rates of the isopentane liquid and isopentane are controlled separately by mass flow meters. The isopentane liquid is pressurized and its flow rate is controlled by a plunger flow pump. The combined liquids then enter the pipeline mixer I installed in the high-temperature oven A. The resulting gas stream of reaction feedstock is heated by heater I before entering reactor I. The main composition of the isopentane used is: isopentane 99.2 wt%, components below C4 0.31 wt%, components above C6 0.16 wt%, n-pentane 0.11 wt%, total olefins 0.04 wt%, and total sulfur 0.8 mg / Nm³. 3 ;

[0093] All reactors, heaters, coolers, pipe mixers, sampling tubes, connecting pipes, ferrules, and other fittings are made of 310S heat-resistant and corrosion-resistant stainless steel.

[0094] Based on the flow rates and chemical composition of the uncondensed gas stream and condensate stream discharged from condenser B, the effects of isopentane conversion, carbon dioxide conversion, isopentene selectivity, and methane formation in the dehydrogenation reaction were calculated. The chemical composition of the high-temperature gas stream at reactor I outlet (sampled every 2-5 hours, sampler temperature around 100℃), the uncondensed gas stream from condenser B (sampler temperature at room temperature), and the condensate stream (sampler temperature -30℃) was detected by two gas chromatographs.

[0095] After reactors I and II are loaded with new catalyst, assembled according to connection method B, and pass a nitrogen tightness test at room temperature (0.5 MPa), catalyst reduction is performed first. The reduction operation conditions are: using a hydrogen flow from the hydrogen pipeline network of the isobutane dehydrogenation production unit (molar purity above 98%, hydrocarbon content below 350 mg / Nm³). 3 Unsaturated hydrocarbon content is less than 100 mg / Nm 3 (Dew point below -18℃) 750 NL / h (Gas hourly space velocity of the catalyst bed in both reactors is 750 h⁻¹) -1The back pressure valve in the connecting pipeline between condensers A and B is set at 0.05 MPa as the heating and reducing gas flow for the reactor. The heating rate of both high-temperature ovens A and B is set at 5℃ / min and the constant temperature is set at 490℃. Hydrogen gas flow is introduced when heating begins. Condenser A is activated (circulating water is supplied) while condenser B is not activated (antifreeze is at room temperature and not circulated). About 1.2 hours after the upper section temperature of the catalyst bed in reactors I and II reaches 480℃, the lower section temperature also reaches above 480℃. Then, reduction continues for 3.5 hours. During the reduction process, the temperature of the upper, middle and lower sections of the two catalyst beds does not exceed 516℃.

[0096] After the reduction operation is completed, immediately activate condenser B (cooling medium circulation temperature -30℃), switch in isopentane liquid flow and carbon dioxide gas flow, simultaneously adjust the temperatures of high-temperature ovens A and B, and adjust the hydrogen flow rate to begin evaluating the dehydrogenation reaction in mode B.

[0097] Process tests were conducted on catalyst HA under the operating conditions listed in Table 1, in Mode B. In Table 1, the isopentane mass hourly space velocity (MHSV) is 2 h⁻¹. -1 That is, an isopentane liquid flow rate of 2000 g / h is equivalent to an isopentane flow rate of 620.9 NL / h after vaporization. The CO2 molar content is calculated. The CO2 molar content in the inlet gas stream of reactor I is adjusted by adjusting the flow rate of carbon dioxide stream I. The CO2 molar content in the inlet gas stream of reactor II is adjusted by adjusting the flow rate of carbon dioxide stream II, based on the estimated flow rate and CO2 content of the outlet gas stream of reactor I. Furthermore, the control of the CO2 molar content in both inlet gas streams is based on ensuring that the temperatures of the catalyst beds in reactors I, II, and II do not exceed 550℃, while also considering the isopentene yield.

[0098] Table 1. Evaluation of dehydrogenation reaction conditions during the first operating cycle of catalyst HA.

[0099]

[0100] Under the operating conditions listed in Table 1, the temperature difference between the inlet gas flow and the oven temperature for both reactors is less than 3°C. Specifically, the isopentane dehydrogenation reaction performance includes: during the evaluation test at the specified time, the maximum temperature of the catalyst bed in both reactors can be controlled below 550°C; the temperature at the outlet of the catalyst bed in reactor I can be controlled above 500°C, and the temperature at the outlet of the catalyst bed in reactor II can be controlled above 530°C; a stable total isopentane conversion rate of over 30% can be achieved (operation with a total isopentane conversion rate of over 35%). The overall reaction effect was evaluated over a time period of 206 hours, with isopentene selectivity exceeding 90% and total CO2 methanation exceeding 80%. Reactor I consistently achieved isopentene conversion exceeding 20%, isopentene selectivity exceeding 90%, and CO2 methanation exceeding 83%. The total isopentene conversion was primarily related to the overall temperature of the catalyst bed in reactor II, and less related to the total amount of carbon dioxide gas introduced. The isopentene conversion in reactor I was closely related to the temperature of the lower section of the catalyst bed, i.e., the outlet temperature. By the 478th hour, the total isopentene conversion had decreased to below 30%, and the system pressure drop (i.e., the bed pressure drop of both reactors) had increased to 0.009 MPa, indicating significant coking had occurred.

[0101] At the 480-hour mark of evaluation, the isopentane and carbon dioxide feeds were shut off, the cooling medium circulation in condenser B was stopped (the circulating water in condenser A remained continuous), the back pressure valve was fully opened, and the first operating cycle was ended. The doors and inlet / outlet ports of high-temperature ovens A and B were opened for about 20 minutes to lower the temperature of ovens A and B to about 490℃. The temperature of ovens A and B was set to 480℃. The hydrogen flow was switched to a 99.9% nitrogen flow at a flow rate of 2000 NL / h (introduced via a hydrogen flow meter and inlet pipeline). After 1 hour of replacement, all six temperatures of the catalyst beds in both reactors were lowered to below 490℃. The first in-situ regeneration of the catalysts in both reactors was then initiated, including carbonization, oxychlorination, and reduction operations. During the carbonization process, 80 NL / h of dry air (dew point -32℃) is introduced through the flow meter and inlet pipe of carbon dioxide gas stream I during the dehydrogenation operation, with an O2 molar content of approximately 0.76% added to the inlet gas stream of reactor I. The O2 content in the outlet gas stream of reactor I is measured every 30 minutes. Based on the measured O2 content in the outlet gas stream of reactor I, an appropriate flow rate of dry air is introduced through the flow meter and inlet pipe of carbon dioxide gas stream II during the dehydrogenation operation to ensure that the O2 molar content in the inlet gas stream of reactor II is approximately 0.7%. When the measured O2 molar content in the outlet gas stream of reactor I rises to 0.7%, the O2 molar content in the outlet gas stream of reactor II is measured (sampled from the exhaust pipe of condenser B). After the O2 molar content in the outlet gas stream of reactor II reaches 0.7%, carbonization continues for another 3 hours. The highest temperature of the catalyst bed during the carbonization process is 515℃.After the carbonization operation is completed, the temperatures of ovens A and B are set to 520℃. The flow rate of dry air supplied through the flow meter and inlet pipe of carbon dioxide gas stream I during the dehydrogenation operation is increased to 500 NL / h, while the flow rate of 99.9% nitrogen gas supplied through the flow meter and inlet pipe is reduced to 500 NL / h (O2 molar content in the inlet gas stream of reactor I is 10%). The back pressure valve is controlled at 0.05 MPa. A chloroacetic acid aqueous solution with a mass concentration of 44.5% and a flow rate of 6.6 g / h (water to chlorine molar ratio 6.5) is supplied through the isopentane flow pump during the dehydrogenation operation. The chloroacetic acid aqueous solution is stopped after 5 hours (the amount of chloroacetic acid supplied, calculated as chlorine, is 0.5 wt% of the total catalyst in both reactors). After about 1 hour, the temperature of the lower section of the catalyst bed in both reactors reached above 515℃. Then, the temperature of ovens A and B was set to 530℃, and the oxychlorination operation was completed after 5 hours. After that, the dry air was stopped, and the doors and inlets / outlets of the high-temperature ovens A and B were opened for about 20 minutes to reduce the temperature of ovens A and B to about 490℃. The temperature of ovens A and B was then set to 480℃, and the flow rate of 99.9% nitrogen gas was increased to 2000 NL / h. After 1 hour, the molar content of O2 in the gas flow at the outlet of reactor II decreased to below 0.1%. The 99.9% nitrogen gas flow was then switched to the hydrogen gas flow from the pipeline at 750 NL / h (the gas hourly space velocity of the catalyst bed in both reactors was 750 h⁻¹). -1 The catalyst reduction ended 5 hours later.

[0102] By appropriately increasing the set temperatures of ovens A and B, and conducting the second operating cycle evaluation test of the isopentane dehydrogenation reaction of catalyst HA under the operating conditions described in Table-1 above, the results showed that the effect of the first operating cycle was basically repeated. The maximum temperature of the catalyst bed in both reactors could be controlled below 550℃. The temperature of the catalyst bed outlet section in reactor I could be controlled above 500℃, and the temperature of the catalyst bed outlet section in reactor II could be controlled above 530℃. The overall reaction effect of achieving a total isopentane conversion rate of over 30% (the operating time for achieving a total isopentane conversion rate of over 35% is 202h), isopentene selectivity of over 90%, and total CO2 methanation rate of over 80% could be stably obtained. Among them, reactor I could stably obtain a reaction effect of an isopentane conversion rate of over 20%, isopentene selectivity of over 90%, and CO2 methanation rate of over 85%. By the 491-hour mark, the total conversion rate of isopentane had dropped below 30%, and the system pressure drop, i.e. the bed pressure drop of the two reactors, had increased to 0.011 MPa, indicating that a significant amount of coking had occurred.

[0103] 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 isopentane 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 isopentane mass hourly space velocity of 2.5 h⁻¹. -1 The molar ratio of isopentane to H2 was 1:0.5; the inlet gas temperature of reactor II was 530℃; 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 550℃. Results included: the outlet temperatures of the catalyst beds in reactors I and II could be maintained above 510℃ and 530℃, respectively; a stable overall reaction effect of over 30% total isopentane conversion (operating time above 35% was 281h), over 90% isopentene selectivity, and over 80% total CO2 methanation rate could be achieved, with reactor I consistently achieving over 20% isopentane conversion, over 90% isopentene selectivity, and over 82% CO2 methanation rate. By the 455th hour of evaluation, the total isopentane conversion had dropped below 30%, and the system pressure drop, i.e., the bed pressure drop of both reactors, had increased to 0.016 MPa (indicating significant coking).

[0104] Following the first in-situ regeneration method described above, the catalysts in both reactors underwent a third in-situ regeneration. Then, using the operating conditions and control methods of the third operating cycle, an evaluation test of the isopentane dehydrogenation reaction in the fourth operating cycle of catalyst HA, in mode B, was conducted until the total isopentane conversion rate dropped below 30%. Afterward, the catalysts in both reactors underwent a fourth in-situ regeneration using the same method as the first regeneration. Finally, using the operating conditions and control methods of the fourth operating cycle, an evaluation test of the isopentane dehydrogenation reaction in the fifth operating cycle of catalyst HA, in mode B, was conducted until the total isopentane conversion rate dropped below 30%. The results essentially replicated the effects of the third operating cycle. The operating times with a total isopentane conversion rate above 30% were 443 h and 457 h, respectively, and the operating times with a total isopentane conversion rate above 35% were 275 h and 284 h, respectively.

[0105] During the evaluation tests of the isopentane dehydrogenation reaction (Mode B) in the third, fourth, and fifth operating cycles, the CO2 molar content in the inlet gas stream of reactor I gradually increased from the initial 1.0 wt% to 2.0 wt%, and the CO2 molar content in the inlet gas stream of reactor II gradually increased from the initial 1.3 wt% to 2.0 wt%.

[0106] Following the first in-situ regeneration method described above, the catalysts in both reactors underwent a fifth in-situ regeneration. Before reduction, heater III and reactor III were connected (other connections were appropriately modified, and reactor III contained new catalyst HA for reduction). The dehydrogenation reaction of isopentane was evaluated in mode C and during the sixth operating cycle. Carbon dioxide streams I and II were shut off, and water was injected through the isopentane pipeline (pressurized and flow controlled by a plunger flow pump). Specific conditions included: inlet gas temperature of reactor I 560℃, pressure 0.03 MPa, and isopentane mass hourly space velocity (MSV) 2 h⁻¹. -1 The molar ratio of isopentane to H2 and water was 1:1:2 (CO2 was not added); the inlet gas flow temperature of reactors II and III was 550℃. Results included: the outlet temperatures of the catalyst beds in reactors I, II, and III could be maintained above 498℃, 510℃, and 530℃, respectively; a stable overall reaction effect of over 30% total isopentane conversion (the operating time for over 35% isopentane conversion was 295 hours) and over 90% isopentene selectivity could be achieved, with reactor I consistently achieving over 15% isopentane conversion and over 90% isopentene selectivity. At 572 hours of evaluation, the total isopentane conversion dropped below 30%, and the system pressure drop, i.e., the bed pressure drop of the two reactors, increased to 0.019 MPa (indicating significant coking).

[0107] Following the first in-situ regeneration method described above, the reactor system underwent a sixth in-situ regeneration (the main difference being that the amount of chloroacetic acid added, calculated as chlorine, was 0.5 wt% of the total catalyst in the three reactors, and the addition time of the 44.5% chloroacetic acid aqueous solution (water to chlorine molar ratio 6.5) at a flow rate of 6.6 g / h was 7.5 h; the catalysts in reactors I and II underwent the sixth in-situ regeneration, while the catalyst in reactor III underwent the first in-situ regeneration). Following this, the isopentane dehydrogenation reaction was evaluated under Mode C and during the seventh operating cycle. The isopentane dehydrogenation reaction conditions included: an inlet gas flow temperature of 560℃ and a pressure of 0.05 MPa for reactor I, and an isopentane mass hourly space velocity of 3 h⁻¹. -1The molar ratio of isopentane to H2 and water was 1:0.5:2.5 (without CO2 added); the inlet gas flow temperature of reactors II and III was 550℃. Results showed that the outlet temperatures of the catalyst beds in reactors I, II, and III could be maintained above 498℃, 510℃, and 530℃, respectively; a stable overall reaction effect of over 30% total isopentane conversion (operating time above 35% was 386 hours) and over 90% isopentene selectivity could be achieved, with reactor I consistently achieving over 15% isopentane conversion and over 90% isopentene selectivity. At 605 hours of evaluation, the total isopentane conversion dropped below 30%, and the system pressure drop, i.e., the bed pressure drop of the two reactors, increased to 0.015 MPa (indicating significant coking).

[0108] Following the sixth in-situ regeneration method described above, the reactor system underwent a seventh in-situ regeneration. Then, under the operating conditions and control methods of the seventh operating cycle (Mode C), an evaluation test of the isopentane dehydrogenation reaction of catalyst HA was conducted for the eighth operating cycle until the total isopentane conversion rate dropped below 30%. The results essentially replicated the effects of the seventh operating cycle. The operating time with a total isopentane conversion rate above 30% was 600 hours, and the operating time with a total isopentane conversion rate above 35% was 373 hours.

[0109] Following the sixth in-situ regeneration method described above, the reactor system underwent its eighth in-situ regeneration. Afterward, the outlet port of reactor I in oven A was connected to condenser A (the high-temperature gas sampling tube continued to be used; the carbon dioxide gas inlet system II, oven B, and the internal piping mixer II, heaters II / III, and reactors II / III were no longer used). The water pump and piping were disconnected, and the isopentane dehydrogenation reaction of mode A and reactor I in its ninth operating cycle was evaluated. Specific conditions included: inlet gas temperature of reactor I 540℃, pressure 0.05MPa, and isopentane mass hourly space velocity 2h⁻¹. -1The molar ratio of isopentane to H2 was 1:1. By adjusting the flow rate of carbon dioxide gas stream I, the overall temperature of the catalyst bed in reactor I was controlled to be as high as possible, but the maximum temperature did not exceed 550℃, and the outlet temperature was maintained above 530℃. Results included: the outlet temperature of the catalyst bed in reactor I could be maintained above 530℃ for an extended period; the evaluation time for isopentane conversion of ≥30% was 452 hours (the operating time for total isopentane conversion of ≥35% was 233 hours); and the reaction effect of isopentene selectivity of ≥90% and CO2 methanation rate of ≥85% was achieved simultaneously. The control method mainly involved gradually increasing the CO2 molar content of the inlet gas stream of reactor I from the initial 1.7% to 2.2%. After 454 hours of evaluation, the reaction effect of isopentane conversion of ≥30% and isopentene selectivity of ≥90% could no longer be achieved simultaneously; the system pressure drop, i.e., the pressure drop of the reactor I bed, rose to 0.08 MPa (indicating significant coking).

[0110] During the nine operating cycles mentioned above, the main dehydrogenation products and side reactions of isopentane included: among the main dehydrogenation product isopentene, 2-methyl-2-butene accounted for 70-82% of the molar percentage, and 3-methyl-1-butene accounted for 1-3% of the molar percentage; the main side reaction was the isomerization reaction to produce n-pentane, with n-pentane accounting for about 80% (by carbon) of the isopentane side reaction products, and the molar ratio of n-pentane to its dehydrogenation products (various straight-chain pentenes) was higher than 20:1; the minor side reaction was the cracking reaction to produce low-carbon hydrocarbons with C4 or less, accounting for less than 15% (by carbon) of the isopentane side reaction products; C5 dienes accounted for less than 6% (by carbon) of the isopentane side reaction products, mainly 3-methyl-1,2-butadiene, with less 2-methyl-1,3-butadiene, i.e., isopentene; and hydrocarbons with C9 or more accounted for less than 2.7% (by carbon) of the isopentane side reaction products.

[0111] From the overall reaction effect of Example 1, it can be seen that under the operating conditions, catalyst HA has both moderate CO2 hydromethanization performance and high isopentane dehydrogenation performance (isopentane dehydrogenation to isopentene). The exothermic CO2 methanation reaction significantly improves the conversion rate of isopentane in the dehydrogenation reaction, while basically not increasing side reactions. In the above evaluation process, an isopentane conversion rate of over 30% was stably obtained at the lower reaction temperature, the isopentene selectivity was stably maintained at over 90%, and a regeneration cycle of over 350 hours was achieved. Controlling the catalyst bed temperature to the aforementioned lower level is the key condition for obtaining a longer regeneration cycle. In-situ regeneration basically did not change or reduce the reaction performance of catalyst HA. The isopentane conversion rate in the dehydrogenation reaction is limited to a certain extent by the thermodynamic reaction equilibrium. The above three equipment configurations and connection methods A, B, and C correspond to the configurations and operating modes A, B, and C in Example 1, respectively. Among them, the reactor conditions in modes B and C are easier to achieve and have a certain degree of controllability, while the reactor conditions in mode A have a narrower controllability but are achievable and relatively simple.

[0112] Catalyst Evaluation Example 2

[0113] In another dehydrogenation pilot plant identical to that used in Catalyst Evaluation Example 1, the Pt-Sn-K / alumina type alkane dehydrogenation catalyst JB prepared in Example 3 was used to conduct a process test of isopentene preparation from isopentane in Evaluation Example 2 under fixed bed and quasi-adiabatic reactor conditions.

[0114] The loading of reactors I-III (the loading amount of catalyst JB is 1000mL), external insulation, connection method, installation and operation of heaters I-III, pipeline mixers I and II, high-temperature ovens A and B, condensers A and B, high-temperature gas flow sampling pipe, carbon dioxide gas flow II inlet pipe, and overall flow direction sequence of the reaction gas flow are completely the same as in catalyst evaluation example 1; the source and purity of isopentane, carbon dioxide and hydrogen used are completely the same as in catalyst evaluation example 1.

[0115] After the first operating cycle of the system (connection method B, using reactors I and II) was assembled and passed the nitrogen tightness test at room temperature and 0.5 MPa, the catalyst was reduced. The reduction operating conditions were exactly the same as those in catalyst evaluation example 1. During the reduction process, the temperature of the upper, middle and lower sections of the two catalyst beds did not exceed 519°C.

[0116] After catalyst JB was reduced, condenser B (cold medium circulation temperature -30℃) was immediately activated, and isopentane liquid and carbon dioxide gas flows were introduced. The temperature conditions and hydrogen flow rates of high-temperature ovens A and B were adjusted to begin the dehydrogenation reaction evaluation. The isopentane dehydrogenation reaction evaluation was conducted for ten operating cycles using essentially the same method as in Catalyst Evaluation Example 1, but with the adoption of in-situ regeneration methods under both chlorine-free and chlorine-containing conditions. First, process tests were conducted on catalyst JB under the operating conditions listed in Table-2 below, in Mode B. The CO2 molar content in the inlet gas flow of reactors I and II was controlled to ensure that the temperatures of the upper, middle, and lower sections of the catalyst bed in reactors I and II did not exceed 550℃, while also considering the isopentene yield.

[0117] Table 2. Evaluation of dehydrogenation reaction conditions during the first operating cycle of catalyst JB.

[0118]

[0119] Under the operating conditions listed in Table 2, the temperature difference between the inlet gas flow and the oven temperature for both reactors is less than 3°C. Specifically, the isopentane dehydrogenation reaction performance includes: during the evaluation test at the specified time, the maximum temperature of the catalyst bed in both reactors can be controlled below 550°C; the temperature at the outlet of the catalyst bed in reactor I can be controlled above 500°C, and the temperature at the outlet of the catalyst bed in reactor II can be controlled above 530°C; a stable total isopentane conversion rate of over 30% can be achieved (operation with a total isopentane conversion rate of over 35%). The overall reaction effect was evaluated over a time period of 389 hours, with isopentene selectivity exceeding 90% and total CO2 methanation exceeding 80%. Reactor I consistently achieved isopentene conversion exceeding 20%, isopentene selectivity exceeding 90%, and CO2 methanation exceeding 85%. The total isopentene conversion was primarily related to the overall temperature of the catalyst bed in reactor II, and less related to the total amount of carbon dioxide gas introduced. The isopentene conversion in reactor I was closely related to the temperature of the lower section of the catalyst bed, i.e., the outlet temperature. By the 633-hour mark, the total isopentene conversion had decreased to below 30%, and the system pressure drop (i.e., the bed pressure drop of both reactors) had increased to 0.012 MPa (indicating significant coking).

[0120] At the 635th hour of evaluation, the isopentane and carbon dioxide feeds were shut off, the cooling medium circulation of condenser B was stopped (the circulating water in condenser A was kept running), the back pressure valve was fully opened, and the first operating cycle was ended. The doors and inlet / outlet of high-temperature ovens A and B were opened for about 20 minutes to reduce the temperature of ovens A and B to about 490℃. The temperature of ovens A and B was set to 480℃. The hydrogen flow was switched to a 99.9% nitrogen flow at a flow rate of 2000 NL / h (introduced through a hydrogen flow meter and inlet pipeline). After 1 hour of replacement, the six temperatures of the catalyst beds in both reactors were reduced to below 490℃. The first in-situ regeneration of the catalysts in both reactors was then started, including carbonization, oxidation, and reduction operations. During the carbonization process, 80 NL / h of dry air (dew point -32℃) is introduced through the flow meter and inlet pipe of carbon dioxide gas stream I during the dehydrogenation operation, resulting in an O2 molar content of approximately 0.7% in the inlet gas stream of reactor I. The O2 content in the outlet gas stream of reactor I is measured every 30 minutes. Based on the measured O2 content in the outlet gas stream of reactor I, an appropriate flow rate of dry air is introduced through the flow meter and inlet pipe of carbon dioxide gas stream II during the dehydrogenation operation to ensure that the O2 molar content in the inlet gas stream of reactor II is approximately 0.7%. When the measured O2 molar content in the outlet gas stream of reactor I rises to 0.65%, the O2 molar content in the outlet gas stream of reactor II is measured (sampled from the exhaust pipe of condenser B). After the O2 molar content in the outlet gas stream of reactor II reaches 0.65%, carbonization continues for another 3 hours. The highest temperature of the catalyst bed during carbonization is 513℃. The carbonization operation is then completed. Afterwards, the temperatures of ovens A and B were set to 510℃. The flow rate of dry air supplied through the flow meter and inlet pipe of carbon dioxide gas stream I during the dehydrogenation operation was increased to 500 NL / h, while the flow rate of 99.9% nitrogen gas supplied through the flow meter and inlet pipe was reduced to 500 NL / h (O2 molar content in the inlet gas stream of reactor I was 10%). The back pressure valve pressure was controlled at 0.05 MPa. The oxidation operation was completed after 7 hours of treatment. Then, the dry air supply was stopped, and the doors and inlet / outlet ports of high-temperature ovens A and B were opened for approximately 20 minutes to reduce the temperature of both ovens to approximately 490℃. The temperature of ovens A and B was then set to 480℃, and the flow rate of 99.9% nitrogen gas was increased to 2000 NL / h. After 1 hour, the O2 molar content in the outlet gas stream of reactor II decreased to below 0.1%. The 99.9% nitrogen gas stream was then switched to the aforementioned pipeline hydrogen gas stream at 750 NL / h (the gas hourly space velocity of the catalyst bed in both reactors was 750 h⁻¹). -1 The catalyst reduction ended 5 hours later.

[0121] Then, appropriately increase the set temperatures of ovens A and B, and then conduct the second operating cycle of catalyst JB, mode B, for the isopentane dehydrogenation reaction evaluation test; specific conditions include: inlet gas flow temperature of reactor I 540℃, pressure 0.03MPa, and isopentane mass hourly space velocity 2h⁻¹. -1 The molar ratio of isopentane to H2 was 1:0.5; the inlet gas temperature of reactor II was 540℃; 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 550℃. Results included: the outlet temperatures of the catalyst beds in reactors I and II could be maintained above 510℃ and 530℃, respectively; a stable overall reaction effect of over 30% total isopentane conversion (operating time above 35% was 377h), over 90% isopentene selectivity, and over 80% total CO2 methanation rate could be achieved, with reactor I consistently achieving over 20% isopentane conversion, over 90% isopentene selectivity, and over 82% CO2 methanation rate. By the 506th hour of evaluation, the total isopentane conversion had dropped below 30%, and the system pressure drop, i.e., the bed pressure drop of both reactors, had increased to 0.015MPa (indicating significant coking).

[0122] Following the first in-situ regeneration method (chlorine-free regeneration) in Example 2 of this evaluation, the catalysts in both reactors underwent a second in-situ regeneration. Then, under the operating conditions and control methods of the second operating cycle, an evaluation test of the isopentane dehydrogenation reaction in the third operating cycle of catalyst JB, in Mode B, was conducted until the total isopentane conversion rate dropped below 30%. Subsequently, the catalysts in both reactors underwent a third in-situ regeneration using the first in-situ regeneration method (chlorine-free regeneration) in Example 2 of this evaluation. Finally, under the operating conditions and control methods of the second operating cycle, an evaluation test of the isopentane dehydrogenation reaction in the fourth operating cycle of catalyst JB was conducted until the total isopentane conversion rate dropped below 30%. The results essentially replicated the effects of the second operating cycle. The operating times with a total isopentane conversion rate above 30% were 491 h and 505 h, respectively, and the operating times with a total isopentane conversion rate above 35% were 364 h and 372 h, respectively.

[0123] During the evaluation tests of the isopentane dehydrogenation reaction (Mode B) in the second, third, and fourth operating cycles, the CO2 molar content in the inlet gas stream of reactor I gradually increased from the initial 1.0 wt% to 2.0 wt%, and the CO2 molar content in the inlet gas stream of reactor II gradually increased from the initial 1.3 wt% to 2.0 wt%.

[0124] Following the first in-situ regeneration method (chlorine-free regeneration) in Example 2 of this evaluation, the catalysts in both reactors underwent a fourth in-situ regeneration. Heater III and reactor III were connected before reduction (other connections were appropriately modified, and reactor III was filled with new catalyst JB for reduction). Subsequently, the dehydrogenation reaction of isopentane was evaluated in Mode C and during the fifth operating cycle. Carbon dioxide streams I and II were shut off, and water was injected through the isopentane pipeline (pressurized and flow controlled by a plunger flow pump). Specific conditions included: inlet gas temperature of reactor I 560℃, pressure 0.03 MPa, and isopentane mass hourly space velocity (MSV) 2.5 h⁻¹. -1 The molar ratio of isopentane to H2 and water was 1:0.5:3.0 (without CO2 added); the inlet gas flow temperature of reactors II and III was 550℃. Results included: the outlet temperatures of the catalyst beds in reactors I, II, and III could be maintained above 495℃, 510℃, and 530℃, respectively; a stable overall reaction effect of over 30% total isopentane conversion (operating time above 35% was 419 h) and over 90% isopentene selectivity could be achieved, with reactor I consistently achieving over 15% isopentane conversion and over 90% isopentene selectivity. At 641 h, the total isopentane conversion dropped below 30%, and the system pressure drop, i.e., the bed pressure drop of the two reactors, increased to 0.017 MPa (indicating significant coking).

[0125] Following the first in-situ regeneration method (chlorine-free regeneration) in Example 2 of this evaluation, the catalysts in both reactors underwent a fifth in-situ regeneration. Subsequently, under the operating conditions and control methods of the fifth operating cycle, an evaluation test of the isopentane dehydrogenation reaction in the sixth operating cycle of catalyst JB and mode C was conducted until the total isopentane conversion rate dropped below 30%. Then, the reactor system was regenerated again using the first in-situ regeneration method of Example 2 (chlorine-free regeneration, with the catalysts in reactors I and II undergoing a sixth in-situ regeneration, and the catalyst in reactor III undergoing a first in-situ regeneration). Following the operating conditions and control methods of the fifth operating cycle, an evaluation test of the isopentane dehydrogenation reaction in the seventh operating cycle of catalyst JB was conducted until the total isopentane conversion rate dropped below 30%. The results essentially replicated the effects of the fifth operating cycle. The operating times with a total isopentane conversion rate above 30% were 653 h and 640 h, respectively, and the operating times with a total isopentane conversion rate above 30% were 415 h and 433 h, respectively.

[0126] Following the first in-situ regeneration method in Example 2 of this evaluation, the catalyst of the reactor system underwent its seventh in-situ regeneration (chlorine-free regeneration). Afterwards, the outlet port of reactor I in oven A was connected to condenser A (the high-temperature gas sampling tube continued to be used; the carbon dioxide gas inlet system II, oven B, and pipeline mixer II, heater II, and reactor II were no longer used). The water pump and pipelines were disconnected, and the dehydrogenation reaction of isopentane in mode A and reactor I was evaluated during the eighth operating cycle. Specific conditions included: inlet gas temperature of reactor I 540℃, pressure 0.05MPa, and isopentane mass hourly space velocity 1.8h⁻¹. -1 The molar ratio of isopentane to H2 was 1:1. By adjusting the flow rate of carbon dioxide gas stream I, the overall temperature of the catalyst bed in reactor I was controlled to the highest possible level, but the maximum temperature did not exceed 550℃, and the outlet temperature was maintained above 530℃ as much as possible. Results included: the outlet temperature of the catalyst bed in reactor I could be maintained above 530℃ for an extended period; the evaluation time for isopentane conversion of 30% or higher was 439 hours (the operating time for total isopentane conversion of 35% or higher was 192 hours); and reaction effects of isopentene selectivity of over 90% and CO2 methanation rate of over 85% were achieved. The control method mainly involved gradually increasing the molar CO2 content of the inlet gas stream in reactor I from the initial 1.8% to 2.5%.

[0127] Subsequently, the carbon dioxide gas flow II intake system, pipeline mixer II, heater II, and reactor II were reconnected. Following the first in-situ regeneration method in Example 2 of this evaluation, the catalysts in reactors I and II were regenerated in situ (chlorine-free regeneration; reactor I underwent its eighth regeneration, and reactor II underwent repeated regeneration for a total of eight times). Then, following the second operating cycle, the operating conditions and control methods of Mode B, the isopentane dehydrogenation reaction evaluation test of catalyst JB continued until the total isopentane conversion rate decreased to below 30%. The results essentially replicated the effects of the second operating cycle, with an isopentane conversion rate above 30% for 486 hours and an isopentane conversion rate above 35% for 369 hours.

[0128] Following the first in-situ regeneration method described in Evaluation Example 1, the catalysts in reactors I and II underwent a ninth in-situ regeneration (with chlorine regeneration). Then, under the operating conditions and control methods of the third operating cycle of Evaluation Example 1, Mode B, an evaluation test of the isopentane dehydrogenation reaction of catalyst JB was conducted until the total isopentane conversion rate decreased to below 30%. Afterwards, the catalysts in both reactors underwent a tenth in-situ regeneration (with chlorine regeneration) following the first in-situ regeneration method described in Evaluation Example 1. Finally, under the operating conditions and control methods of the third operating cycle of Evaluation Example 1, an evaluation test of the isopentane dehydrogenation reaction of catalyst JB was conducted until the total isopentane conversion rate decreased to below 30%. The operating conditions included: inlet gas flow temperature of reactor I 540℃, pressure 0.05MPa, and isopentane mass hourly space velocity 2.5h⁻¹. -1 The molar ratio of isopentane to H2 is 1:0.5; the inlet gas flow temperature of reactor II is 530℃; by adjusting the flow rates of carbon dioxide gas flow I and II, the overall temperature of the catalyst beds in reactors I and II is controlled to a reasonably high level, but the maximum temperature never exceeds 550℃. The results include: the temperatures at the catalyst bed outlet sections of reactors I and II can be maintained above 510℃ and 530℃, respectively; a stable overall reaction effect of over 30% total isopentane conversion (operating time of over 35% isopentane conversion is 281h), over 90% isopentene selectivity, and over 80% total CO2 methanation rate can be achieved, with reactor I achieving a stable reaction effect of over 18% isopentane conversion, over 90% isopentene selectivity, and over 85% CO2 methanation rate; the control method mainly involves gradually increasing the CO2 molar content of the inlet gas flow of reactor I from the initial 1.1% to 1.8%, and gradually increasing the CO2 molar content of the inlet gas flow of reactor II from the initial 1.3% to 2.0%; the operating times of over 30% isopentane conversion are 756h and 742h, respectively, and the operating times of over 35% isopentane conversion are 501h and 515h, respectively.

[0129] In the eleven operating cycles described in Example 2 of this evaluation, the main dehydrogenation products and side reactions of isopentane included: In the main dehydrogenation product isopentene, 2-methyl-2-butene accounted for 78-86% of the molar percentage, and 3-methyl-1-butene accounted for 1-2.5% of the molar percentage; the main side reaction was the isomerization reaction to produce n-pentane, which accounted for approximately 70% (by carbon) of the isopentane side reaction products, with a molar ratio of n-pentane to its dehydrogenation products (various straight-chain pentenes) exceeding 20:1; the minor side reaction was the cracking reaction to produce low-carbon hydrocarbons with C4 or less, accounting for less than 20% (by carbon) of the isopentane side reaction products; C5 dienes accounted for less than 5% (by carbon) of the isopentane side reaction products, primarily 3-methyl-1,2-butadiene, with 2-methyl-1,3-butadiene (i.e., isopentene) frequently undetectable; and hydrocarbons with C9 or more accounted for less than 2.1% (by carbon) of the isopentane side reaction products.

[0130] The overall reaction effect of Example 2 shows that under the operating conditions, catalyst JB exhibits both moderate CO2 hydromethanization performance and high isopentane dehydrogenation performance (isopentane dehydrogenation to isopentene). The exothermic CO2 methanation reaction significantly improves the isopentane conversion rate in the dehydrogenation reaction without significantly increasing side reactions. During the evaluation, an isopentane conversion rate of over 30% was stably achieved at the lower reaction temperature, the isopentene selectivity was stably maintained above 90%, and a regeneration cycle of over 450 hours, even exceeding 700 hours, was obtained. Controlling the catalyst bed temperature to the aforementioned lower level is the key condition for obtaining a longer regeneration cycle. Both the chlorine-free and chlorine-containing in-situ regeneration methods showed good regeneration effects on catalyst JB. The isopentane conversion rate in the dehydrogenation reaction was also significantly limited by the thermodynamic reaction equilibrium. The three equipment configurations and connection methods A, B, and C described in this embodiment 2 correspond to the configuration and operation modes A, B, and C in embodiment 1, respectively. In mode B and C, the reactor conditions are easier to achieve and have a certain degree of control flexibility. In mode A, the control flexibility of the reactor conditions is narrower but achievable and relatively simple.

[0131] Example 5

[0132] Example 5 presents a process for preparing tert-amyl alcohol. Using isopentenyl alcohol stream Z-2L or Z-21L (molar purity ≥ 99%) obtained from Example 1 or 2 and water as raw materials, it is mixed with recycled residual isopentenyl alcohol stream X-1, recycled aqueous solution stream containing tert-amyl alcohol X-31, and circulating solvent stream X-4. The resulting stream X is then subjected to a liquid-phase reaction of isopentenyl alcohol hydration to tert-amyl alcohol in a fixed-bed hydration reactor packed with a macroporous sulfonic acid-type cationic resin catalyst (hydrogen form, our product JQE-02). The circulating solvent stream X-4 contains ethylene glycol monobutyl ether as solvent. The mixture is then passed sequentially through a light-removal column D3 (theoretical plate number approximately 20) and a distillation column D4 (theoretical plate number 40-50). The liquid stream obtained from the hydration reaction is split to obtain the remaining isopentenyl liquid stream X-1, the tert-amyl alcohol and water mixture stream X-20, and the circulating solvent liquid stream X-4. The tert-amyl alcohol and water mixture stream X-20 is then passed through a distillation column D5 (theoretical plate number 40-50) with a top phase separator and azeotropic distillation to separate the tert-amyl alcohol liquid stream X-2 and the aqueous solution containing tert-amyl alcohol stream X-3. The remaining isopentenyl alcohol liquid stream X-1, the aqueous solution containing tert-amyl alcohol stream X-3, and the circulating solvent liquid stream X-4 are all recycled for the hydration reaction. The operating conditions of the hydration reactor include: the molar ratio of water, isopentenyl alcohol, and ethylene glycol monobutyl ether in the feed stream X is (1-2):1:(4-6), and the liquid hourly space velocity is 0.6-1 h⁻¹. -1 The feed temperature is 40-65℃, the pressure is 0.5-1.0MPa, and the discharge temperature is 65-80℃. The hydration reactor is also equipped with an external circulation system, which reuses 40-50% of the volumetric flow rate of the discharge liquid as circulating material liquid X-5 and mixes it with liquid X before it re-enters the hydration reactor. A coil with 60-65℃ circulating water is installed in the resin catalyst bed as a heat transfer and temperature control component. A portion of the remaining isopentenene liquid X-1 is also sent back to the isopentenene separation unit. The cooling separation system is combined with the hydrocarbon-containing liquid stream L0 for processing (mainly to remove alkanes and 3-methyl-1-butene) to maintain the total content (molar content) of 2-methyl-2-butene and 2-methyl-1-butene in the remaining isopentenene liquid stream X-1 at more than 90%. The circulating solvent liquid stream X-4 is periodically fractionated and purified to separate and remove components other than ethylene glycol monobutyl ether that are not conducive to improving the purity of the tert-amyl alcohol liquid stream X-2, especially isopentenene polymers and fusel alcohols.

[0133] The tert-amyl alcohol and water mixture stream X-20, together with the tert-amyl alcohol-water azeotrope separated from the top of distillation column D4, is temperature-controlled and separated into layers in the phase separator at the top of distillation column D4. The upper layer is a water-containing tert-amyl alcohol solution stream X-21, and the lower layer is an aqueous solution stream X-31 containing tert-amyl alcohol. The water-containing tert-amyl alcohol solution stream X-21 is used as the middle feed of distillation column D4. Finally, a tert-amyl alcohol liquid stream X-2 with a purity of 99 wt% or even 99.5 wt% or higher and a water content of less than 0.1 wt% is obtained. After cooling, it can be packaged, stored, and transported as a tert-amyl alcohol product.

[0134] In the tert-amyl alcohol preparation process of this embodiment 5, the hydration reactor can achieve and guarantee the reaction effect of isopentenene conversion rate ≥40% and tert-amyl alcohol selectivity ≥98% during long-term operation for more than one year; in the subsequent separation process of the hydration reaction, the separation yield of tert-amyl alcohol can be guaranteed to be above 98%; the purity of the tert-amyl alcohol product can be controlled to be above 99.5wt% or even above 99.8wt%.

Claims

1. A process for preparing iso-pentene by dehydrogenation of iso-pentane, comprising an iso-pentane dehydrogenation device and an iso-pentene separation device; the iso-pentane dehydrogenation device comprises a dehydrogenation reactor group, and corresponding connecting pipelines, temperature measuring and control components, flow measuring and control components and valves; the reactors of the dehydrogenation reactor group are fixed bed adiabatic reactors with axial or radial reaction structure, and are loaded with Pt-Sn-K / alumina type alkane dehydrogenation catalyst H or catalyst J; the catalysts have hydro-methanation performance under dehydrogenation reaction conditions containing CO2; The setting and operation of the dehydrogenation reactor group comprises any of the following modes: During the stable operation of iso-pentane dehydrogenation, the inlet gas flow conditions of each reactor comprise: Mode A: 2-4 reactors are set, the catalyst loading volume in each reactor is the same, and a heating furnace is set before the gas inlet of each reactor; the dehydrogenation of isopentane is carried out in parallel in each reactor; Mode B: 2-3 reactor series are arranged in parallel, each reactor series comprises two reactors I and II connected in series, wherein reactor I is the former reactor, and the loading volume ratio of the catalysts in reactors I and II is 1:(0.8-1.2); Temperature 520-550°C, pressure ≤0.2 MPa, isopentane mass space velocity 1-2 h -1 , isopentane to H2 molar ratio 1:(0.5-1.5), CO2 molar content 1.5-2.5%; the maximum temperature of each reactor catalyst bed is controlled to be below 550°C; the outlet gas stream temperature or the catalyst bed outlet section temperature of each reactor is controlled to be above 530°C; A heating furnace is arranged in front of the gas flow inlet of each reactor; When the performance index of a reactor in mode A or a reactor series in modes B or C decreases to a specified range, the dehydrogenation operation is cut off and in-situ catalyst regeneration is carried out, and after the regeneration is completed, the dehydrogenation operation is cut in; the pressure is in bar; During the stable operation of the dehydrogenation of isopentane, the inlet gas stream conditions of reactor I include: temperature 520-550℃, pressure ≤0.2MPa, isopentane mass space velocity 1.5-3h -1 -1, molar ratio of isopentane to H2 1:(0.5-1), CO2 molar content 1-2%; the inlet gas stream conditions of reactor II include: temperature 510-540℃, CO2 molar content 1-2%; the maximum temperature of the catalyst bed of each reactor is controlled to be below 550℃; the outlet gas stream temperature of reactor II or the temperature of the outlet section of the catalyst bed is controlled to be above 530℃; Mode C: 2-3 reactor series in parallel are set, each reactor series includes three reactors III, IV, V in series, wherein reactor III is the first reactor, the volume ratio of catalyst loading in reactors III, IV, V is 1:(0.8-1):(0.8-1); a heating furnace is arranged in front of the gas inlet of each reactor; during the stable operation of isopentane dehydrogenation, the inlet gas flow conditions of reactor III include: temperature 550-560℃, pressure ≤0.2MPa, isopentane mass space velocity 1.5-2.5h -1 -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, V is controlled at 540-550℃; the outlet gas flow temperature or the outlet section temperature of catalyst bed of reactor V is controlled at above 530℃; The catalyst H uses high-temperature zirconium-aluminum composite oxide containing 0.8-1.5wt% ZrO2 as the carrier, and is loaded with Pt 0.2-0.4wt%, Sn 0.5-1wt% and K 0.65-1wt% in terms of elements, and is prepared by the following steps: (H-2) The zirconium-modified alumina carrier obtained in step (H-1) is uniformly applied with a Pt-Sn-acid-K aqueous solution with a saturation water absorption volume of 80-90% of the carrier and a required concentration and temperature of 20-30℃, is placed in a closed container for homogenization treatment for 10-15h, is dried, and then is calcined at 500-600℃ for 2-5h, and is cooled to obtain an alkane dehydrogenation catalyst; (H-1) with a pore volume of 0.5-0.7 mL / g and a specific surface area of ​​100-130 m² 2 Using spherical or strip-shaped alumina with dimensions of Φ2.5-3.0mm as a carrier, a zirconium oxychloride aqueous solution of 80-100% of the carrier's saturated water absorption volume and the required concentration is uniformly applied. The carrier is then placed in a sealed container for homogenization treatment for 10-20 hours, dried, and then calcined at 880-930℃ for 2-4 hours and cooled to obtain a zirconium-modified alumina carrier. The Pt-Sn-acid-K aqueous solution in step (H-2) is an aqueous solution prepared from chloroplatinic acid, stannous chloride, hydrochloric acid and citric acid, and potassium nitrate, and the preparation method comprises: chloroplatinic acid and hydrochloric acid are prepared into an aqueous solution, stannous chloride, potassium nitrate and citric acid are added and dissolved, and are uniformly mixed to obtain the Pt-Sn-acid-K aqueous solution; wherein the concentration of HCl is 0.05-0.1mol / L, and the concentration of citric acid is 0.2-0.4mol / L; The catalyst J uses magnesium-modified alumina containing 1-2wt% MgO as the carrier, and is loaded with Pt 0.2-0.4wt%, Sn 0.5-1wt% and K 0.65-1wt% in terms of elements, and is prepared by the following steps: (J-2) The magnesium-modified alumina carrier obtained in step (J-1) is uniformly applied with a Pt-Sn-EDTA-K aqueous solution with a saturation water absorption volume of 80-90% of the carrier and a required concentration and temperature of 40-70℃, is placed in a closed container for homogenization treatment for 3-8h, is dried, and then is calcined at 500-600℃ for 2-5h, and is cooled to obtain an alkane dehydrogenation catalyst; (J-1) with a pore volume of 0.5-0.7 mL / g and a specific surface area of ​​100-130 m² 2 Using spherical or strip-shaped alumina with dimensions Φ2.5-3.0mm as a carrier, a magnesium bicarbonate aqueous solution of the required concentration, accounting for 40-80% of the carrier's saturated water absorption volume, is uniformly applied. The mixture is then placed in a sealed container for homogenization treatment for 5-20 hours. Under flowing air conditions, the temperature is raised to 70-75℃ and kept constant for 3-5 hours. After drying, the mixture is calcined at 880-930℃ for 2-4 hours and then cooled to obtain a magnesium-modified alumina carrier. ​ The Pt-Sn-EDTA-K aqueous solution in step (J-2) is an aqueous solution prepared from diamminosnitrato platinum, potassium stannate, EDTA, nitric acid, potassium hydroxide and potassium nitrite, and the preparation method is as follows: (A) potassium stannate is added into an aqueous solution of potassium hydroxide with a pH value of 10-12, and then dissolved, and then EDTA powder and an aqueous solution of nitric acid are added, and the reaction is carried out until the solution is clear, and then an aqueous solution of potassium hydroxide is added to adjust the pH value to 6.0-6.5, and then potassium nitrite or an aqueous solution thereof is added, and then dissolved and / or mixed, to obtain the Sn-EDTA-K aqueous solution; wherein the molar ratio of Sn, EDTA and nitric acid is 1:(1-1.03):(2.1-2.3); (B) the Sn-EDTA-K aqueous solution in step (A) is controlled at a temperature of 50-70℃, and then an aqueous solution of diamminosnitrato platinum at 60-70℃ is added, and then stirred, to obtain the Pt-Sn-EDTA-K aqueous solution.

2. The process for the preparation of iso-pentene by dehydrogenation of iso-pentane as claimed in claim 1 wherein, The isopentane dehydrogenation device is further provided with a heat exchanger T1; the mixed stream M of the isopentane liquid stream and / or the isopentane-rich gas stream, the hydrogen stream and / or the hydrogen-rich stream and the carbon dioxide gas stream is countercurrently exchanged with the post-reaction gas stream P from the dehydrogenation reactor group through the heat exchanger T1, to obtain the pre-reaction gas stream N and the post-cooling gas stream Q, the pre-reaction gas stream N enters the dehydrogenation reactor group, and the post-cooling gas stream Q enters the isopentene separation device.

3. The process for the production of iso-pentene by dehydrogenation of iso-pentane as claimed in claim 1 wherein, The reduction operation before the dehydrogenation operation after the reactor is charged with the new catalyst H or catalyst J, and the last step of the reduction operation in the in-situ regeneration process of the catalyst, all under the conditions of using a hydrogen stream with a H2 molar content higher than 98% and an olefin molar content lower than 0.1%, an airspeed of 500-1000h -1 , a pressure of 0-0.2MPa, a reduction time of 3-6h at a temperature above 480℃, and a maximum bed temperature not exceeding 520℃.

4. The process for the production of iso-pentene by dehydrogenation of iso-pentane as claimed in claim 1 wherein, The reactor or series filled with the catalyst H or the catalyst J is subjected to an in-situ regeneration process after the dehydrogenation operation is cut off, which includes carbon burning, oxychlorination and reduction operations; ​ The oxygen-containing gas stream with 0.5-0.8% O2 molar content is obtained by mixing air stream with nitrogen stream, the space velocity is 500-2000h -1 -1, the inlet temperature is 450-490℃, the pressure is 0-0.1MPa, the initial temperature of the bed is 450-500℃, the maximum temperature of the bed is not more than 520℃, and the difference between the O2 molar content of the outlet gas stream and the inlet gas stream is reduced to below 0.1%, and then the operation is prolonged for 2-4h. The oxygen-containing gas stream with 4-15% of O2 molar content is used in the oxygen chlorination operation, the air stream is mixed with nitrogen stream, the space velocity is 500-2000h -1 -1, the pressure is 0.02-0.2MPa, the initial temperature of the bed is 480-520℃; water or steam and dichloroethane, trichloroethane, carbon tetrachloride or tetrachloroethylene are mixed in the oxygen-containing gas stream, or the aqueous solution of chloroacetic acid with 40-50% of mass concentration is directly mixed before the heating furnace or heater, the inlet temperature of the gas stream is controlled to be 500-520℃, the mixed amount of chlorides is 0.4-0.6wt% of the catalyst in the reactor in terms of chlorine, the molar ratio of water to chlorine is controlled to be 5-8, and the mixing time is controlled to be 3-6h; then the inlet temperature of the oxygen-containing gas stream without water and chlorides is controlled to be 520-540℃, the treatment time is 4-6h, and the outlet temperature of the reactor is above 515℃.

5. The process for the production of iso-pentene by dehydrogenation of iso-pentane as claimed in claim 1 wherein, The reactor or series filled with the catalyst J is subjected to an in-situ regeneration process under a chlorine-free condition after the dehydrogenation operation is cut off, which includes carbon burning, oxidation and reduction operations; ​ The oxygen-containing stream with 0.5-0.8% O2 mole content is obtained by mixing dry air stream with nitrogen stream, the space velocity is 500-2000h -1 -1, the inlet temperature is 450-490℃, the pressure is 0-0.1MPa, the initial temperature of the bed is 450-500℃; during the burning process, the highest temperature of the bed is controlled to be not more than 520℃, until the difference of O2 mole content between the outlet stream and the inlet stream is reduced to below 0.1%, and then extended for 2-4h. The oxidation operation is carried out using an oxygen-containing stream with a molar content of 10-20% of O2, obtained by feeding dry air into a nitrogen stream, or a dry air stream, at a space velocity of 500-2000 h -1 -1, a pressure of 0.02-0.2 MPa, a bed initial temperature of 480-520°C, an inlet temperature of the stream of 500-520°C, and for a time of 6-10 h.

6. The process for the production of iso-pentene by dehydrogenation of iso-pentane as claimed in claim 1 wherein, The isopentene separation device comprises a post-cooling separation system and a pressure swing adsorption separation system; the isopentane dehydrogenation post-reaction gas stream Q is further cooled and pressurized in the post-cooling separation system, to be cooled to-20 to-10℃ and pressurized to 0.5-1 MPa, and then water is separated, to obtain a gas stream G0 containing H2, CH4 and CO2, ethane and ethylene, and a hydrocarbon-containing liquid stream L0; the hydrocarbon-containing liquid stream L0 enters the de-light column D0, the low-boiling gas stream G-1 containing carbon four and 3-methyl-1-butene is separated from the top of the column, and the hydrocarbon-containing liquid stream L is separated from the bottom of the column; The pressure swing adsorption unit comprises a PSA-H2 unit for separating hydrogen from the gas stream G0, and an alkene separation unit PSA-CH-1 and an alkane separation unit PSA-CH-2; The gas stream G0 is separated in the PSA-H2 unit to obtain a hydrogen stream G-H2 with a molar purity of more than 98%, part of which is recycled to the isopentane dehydrogenation device, and the rest is externally supplied; A low-hydrogen gas stream G-2 containing CH4, ethane, ethylene and CO2 is also obtained, which is used as fuel gas in a heating furnace; The gasified hydrocarbon-containing liquid stream L is separated in the PSA-CH-1 unit into three parts of an alkane gas stream GL-1, an isopentene gas stream Z-2 and a hetero-hydrocarbon gas stream Z-3; the alkane gas stream GL-1 is further separated in the PSA-CH-2 unit into two parts of an isopentane gas stream Y-1 and a n-pentane gas stream Z-1, and the isopentane gas stream Y-1 is recycled to the isopentane dehydrogenation device.

7. A process for the preparation of tertiary amyl alcohol, characterized in that, The process of any one of claims 1-6 is used to prepare isopentene and water, and the isopentene and water are used as raw materials, mixed with the recycled residual isopentene stream X-1, the recycled water stream X-3 or the isopentanol-containing aqueous solution stream X-3, and the circulating solvent stream X-4 to obtain stream X, and the liquid phase reaction of isopentene hydration to produce tertiary amyl alcohol is carried out in a fixed bed hydration reactor packed with macroporous sulfonic acid cation resin catalyst, and the solvent contained in the circulating solvent stream X-4 is ethylene glycol monobutyl ether; the stream obtained by the hydration reaction is separated by a light component removal column D3 and a rectification column D4 in sequence to obtain the residual isopentene stream X-1, the isopentanol and water miscible stream X-20, and the circulating solvent stream X-4; the isopentanol and water miscible stream X-20 is separated by azeotropic rectification, extractive rectification or extractive-azeotropic rectification to obtain the isopentanol stream X-2, and the water stream X-3 or the isopentanol-containing aqueous solution stream X-3; the residual isopentene stream X-1, the water stream X-3 or the isopentanol-containing aqueous solution stream X-3, and the circulating solvent stream X-4 are all recycled for use in the hydration reaction; the operating conditions of the hydration reactor include that the molar ratio of water, isopentene and ethylene glycol monobutyl ether in the stream X is (1-3):1:(3-8), the liquid hourly space velocity is 0.3-1.2 h-1, the feed temperature is 40-65℃, the pressure is 0.5-1.0 MPa, and the discharge temperature is 65-80℃; and the isopentanol stream X-2 obtained is cooled to obtain the isopentanol product. -1 ​ 8. The t-amyl alcohol production process of claim 7, wherein, The hydration reactor is provided with an external circulation system, 40-70% of the volume or mass flow of the effluent stream is recycled as a recycle material stream X-5 and mixed into the stream X to be fed into the hydration reactor; a coil pipe is arranged in the catalyst bed of the hydration reactor to circulate 60-65℃ water; the total molar content of 2-methyl-2-butene and 2-methyl-1-butene in the remaining isopentene stream X-1 is maintained above 90% by sending the required flow of the remaining isopentene stream X-1 back to the cooling separation system of the isopentene separation device and combining and processing with the hydrocarbon-containing stream L0; the extractant used in the extractive rectification method or extractive-azeotropic rectification method is cyclohexane or n-hexane; the recycle solvent stream X-4 is periodically subjected to fractionation purification treatment.

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