Process for preparing isoamylene by dehydrogenation of isopentane, catalyst and tert-amyl alcohol preparation process

Abstract

The invention belongs to the technical field of alkane processing, and particularly relates to a process for preparing isopentene by dehydrogenation of isopentane, a catalyst and a tert-amyl alcohol preparation process. The process for preparing isopentene by dehydrogenation of isopentane comprises an isopentane dehydrogenation device and an isopentene separation device, in the isopentane dehydrogenation device, a dehydrogenation reactor is a fixed bed adiabatic reactor and is filled with a Pt-Sn-K / alumina type alkane dehydrogenation catalyst, the catalyst has hydromethanation performance under the CO2-containing dehydrogenation reaction condition, and the arrangement and operation of a dehydrogenation reactor group comprise a mode A, a mode B or a mode C; the isopentene separation device comprises a cooling separation system, a pressure swing adsorption separation system and an optional rectification system. The tert-amyl alcohol is further prepared from isoamylene under the conditions of a fixed bed hydration reactor, ethylene glycol monobutyl ether and other solvents, and the tert-amyl alcohol product is obtained through separation. The process is superior to the prior art on the whole, and has certain application prospects.

Description

TECHNICAL FIELD

[0001] The present application belongs to the technical field of alkane processing, and particularly relates to a process for preparing iso-pentene by dehydrogenation of iso-pentane, a catalyst and a process for preparing tertiary amyl alcohol. BACKGROUND

[0002] Preparation of iso-pentene by dehydrogenation of iso-pentane, and further conversion, such as continued dehydrogenation to prepare iso-pentadiene, etherification reaction with methanol to prepare methyl tertiary amyl ether (TAME) which can be used as a gasoline blending component, and reaction with water to prepare tertiary amyl alcohol, has certain commercial value; the iso-pentane dehydrogenation process and the catalyst used are the key.

[0003] Preparation of propylene and iso-butene by dehydrogenation of propane and iso-butane is a related process of preparation of iso-pentene by dehydrogenation of iso-pentane. Due to the increase in complexity of the iso-pentane dehydrogenation product, the process for preparing iso-pentene from iso-pentane based on limited improvement of the iso-butane dehydrogenation process and catalyst generally has poor effect.

[0004] In the process for preparing iso-butene by dehydrogenation of iso-butane disclosed in CN103232312A, hydrogen and steam are used as diluent and heat carrier, and the dehydrogenation reaction of iso-butane is carried out in a fixed bed adiabatic reactor loaded with a Pt-Sn-K / alumina type platinum dehydrogenation catalyst, and then iso-butene is prepared by further separation, which has achieved industrialization and reached good technical effect and economic benefit. In the process device, the iso-butane and iso-pentane dehydrogenation device adopts two dehydrogenation reactor series, each dehydrogenation reactor series includes two serially connected dehydrogenation reactors, the dehydrogenation reactors adopt radial or axial reaction structure, and each dehydrogenation reactor is provided with a stream heating furnace in front; the catalyst used in the dehydrogenation reactor is a spherical catalyst (outer diameter 2.5-3.0 mm) prepared based on the ratio and method in CN103212411A, Pt is uniformly distributed in the spherical particles and is in a highly dispersed state; the gas inlet conditions of each hydrogenation reactor are as follows: temperature 580-650℃, pressure 0.1-0.5MPa, and space velocity 2000-6000h -1 The volume ratio of the feed gas of the first dehydrogenation reactor in each dehydrogenation reactor series is iso-butane:hydrogen:steam=1:(0.5-3):(0.5-3), and the specific gas inlet conditions are adjusted according to the reaction effect, which can reach the target reaction effect of iso-butane single-pass conversion rate of more than 30% or even more than 35% and iso-butene selectivity of more than 97%; the dehydrogenation catalyst needs to be regenerated in situ in the reactor after gradual carbon deposition and reaction effect decline, and the regeneration period can reach more than 400h or even more than 750h.

[0005] Two kinds of Pt-Sn-K / alumina type alkane dehydrogenation catalysts, which have been applied in the industrial device of the process for preparing isobutene by dehydrogenation of isobutane described in CN103232312A and achieved good technical effects, i.e. catalyst A (containing 0.35wt% Pt, 0.7wt% Sn, 1wt% K2O, and the carrier being high-temperature zirconium-aluminum composite oxide containing ZrO21.1wt%) and catalyst B (containing 0.25wt% Pt, 0.7wt% Sn, 1wt% K2O, and the carrier being high-temperature zirconium-aluminum composite oxide containing ZrO22.2wt%) in Example 1 and Example 4 of CN103212411A, are used as the basis for the preparation and proportioning method of the spherical catalyst (with an outer diameter of 2.5-3.0mm, and Pt uniformly distributed in the spherical particles, i.e. the cross section color is uniform and in a highly dispersed state at the early stage of each running cycle) with the same chemical composition. The same set of dehydrogenation pilot plant used in the process development for preparing isopentene by dehydrogenation of isopentane described in CN103232312A is used to carry out the process test for preparing isopentene by dehydrogenation of isopentane under the condition of fixed bed adiabatic reactor, and it is found that there is a significant difference in effect between the dehydrogenation of isobutane and isopentane, and it is difficult to simultaneously control the single-pass conversion rate of isopentane to be more than 30%, the selectivity of isopentene to be more than 90%, and the regeneration cycle to be more than 300h. The dehydrogenation pilot plant includes two reactors I and II connected in series, wherein reactor I is the former reactor, the loading volume ratio of the catalysts in reactors I and II is 1:1, a heater is arranged in front of the gas inlet of each reactor, and the specific configuration is the same as that of the catalyst evaluation example 1 of the present application in the following text, connection mode B. The main operating conditions include: pressure 0.02-0.5MPa, liquid hourly space velocity of isopentane 0.6-6h -1 , the molar ratio of the feed gas stream of the first dehydrogenation reactor is isopentane:hydrogen:water vapor=1:(0.5-3):(0-3), and the inlet gas temperature of the two dehydrogenation reactors is adjusted according to the reaction effect within the range of 540-600℃.

[0006] In the dehydrogenation pilot plant, using the same experimental control method, the spherical catalyst prepared on the basis of the catalyst C-E ratio and method of CN103212411A (outer diameter 2.5-3.0 mm, K content of catalyst E increased to 1 wt% as K2O, Pt uniformly distributed in the spherical particles, i.e. uniform color in cross section and high dispersion state at the initial stage of operation; all containing 0.35 wt% Pt, 0.7 wt% Sn, 1 wt% K2O, the carrier in turn being high-temperature magnesium-aluminum composite oxide containing MgO 2.27 wt%, high-temperature magnesium-aluminum composite oxide containing MgO 1.13 wt%, high-temperature zinc-aluminum composite oxide containing ZnO 8.14 wt%) and commercially available low-carbon alkane (C3-C6 alkane) dehydrogenation catalysts, low-carbon alkane dehydrogenation catalysts prepared according to other relatively optimal existing technology methods (such as CN107537485A and the article by Liu Changzheng), were respectively subjected to process tests for preparing isoamylene from isoamylane dehydrogenation, and it was found that it was also very difficult or impossible to simultaneously achieve the basic technical requirements of isoamylane single-pass conversion rate of 30% or more, isoamylene selectivity of 90% or more and regeneration cycle of 300 h or more, wherein the catalyst prepared according to the method in the article by Liu Changzheng could not maintain a regeneration cycle (chlorine regeneration) of 280 h or more when the first dehydrogenation reactor inlet gas temperature was 540-550°C and the second dehydrogenation reactor inlet gas temperature was 530-550°C; the tested catalysts were all spherical with an outer diameter of 2.2-3.0 mm, a pore volume of 0.40 mL / g or more and a specific surface area of 80 m 2 / g or more.

[0007] The catalyst for preparing olefins by dehydrogenation of low carbon alkanes disclosed in CN103212411A is made of the following components: the carrier is a high-temperature composite oxide containing 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 metal ruthenium, rhodium, platinum and iridium; the first additive is one or more of the fourth main group elements germanium, tin or lead; the second additive is one or more of the alkali metal elements lithium, sodium, potassium, rubidium, cesium and francium; preferably, the active component is metal platinum, the mass fraction based on the carrier is 0.2-0.5%; the first additive is tin, the molar ratio to the active component is (1-3):1; the second additive is potassium, the mass fraction based on the carrier is 0.1-3% in terms of oxide; in the high-temperature composite oxide carrier, the content of aluminum oxide is 77-95wt%, the content of magnesium oxide is 1-10wt%, the content of zirconium oxide is 1-5wt% or the content of zinc oxide is 2-10wt%; the preparation method of the high-temperature composite oxide carrier is as follows: one or more soluble salts containing magnesium oxide, zirconium oxide or zinc oxide are dissolved in water to form an aqueous solution with a concentration of 0.05-1.0mol / L, a precursor of aluminum oxide, pseudo-boehmite, is mixed with the aqueous solution, then impregnation is carried out at a temperature of 0-100℃ for 0.1-24 hours, drying is carried out at a pressure of 0.01-0.1MPa and a temperature of 30-200℃ for 1-24 hours, and finally calcination is carried out in an air atmosphere at a temperature of 700-1000℃ for 2-10 hours; the preparation method of the catalyst is as follows: the active component such as chloroplatinic acid and the first additive stannous chloride are added to an acidic solution containing inorganic acid (hydrochloric acid or nitric acid) and / or organic acid (one or more of maleic acid, oxalic acid, acetic acid, citric acid, tartaric acid and lactic acid) to form a stable complex, and then the second additive such as potassium nitrate is added to form a stable impregnation solution; the high-temperature composite oxide carrier is added to the impregnation solution, the volume ratio of the impregnation solution to the carrier is (0.1-20):1, the impregnation time of the high-temperature composite oxide carrier in the impregnation solution is 0.1-48 hours, the impregnation temperature is 0-100℃, after impregnation, drying is carried out in air at a temperature of 50-200℃ for 3-24 hours, and calcination is carried out in air at a temperature of 400-600℃ for 1-24 hours to obtain the catalyst. The catalyst is reduced by hydrogen before being contacted with the reaction gas stream of low carbon alkanes, the reduction is carried out at a temperature of 400-600℃, and the reduction time is 1-24 hours.

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

[0009] Liu Changzheng et al. in "Reaction rule of Pt-Sn-K / Al2O3 catalyst for converting C5-C7 alkanes to olefins and aromatic hydrocarbons" [J / OL], Petroleum (Petroleum Processing), 2025-05-07, disclose a catalyst preparation method: a certain amount of aluminum hydroxide powder is taken, a proper amount of deionized water is added and stirred, and a proper amount of dilute nitric acid is added dropwise to obtain a sol; the sol is acidified by adding a hydrochloric acid solution containing SnCl4·5H2O under stirring, and then kerosene and fatty alcohol polyoxyethylene ether are added dropwise under stirring, followed by drop-sphere formation in an oil-ammonia column; the wet spheres are solidified in ammonia water, taken out, washed with deionized water, dried, then calcined at 650°C in a muffle furnace for 4h, and then two-stage calcination at 1000°C for 4h to prepare a spherical Sn-Al2O3-containing carrier; a chloroplatinic acid impregnation solution with a Pt mass fraction of 0.305% (based on dry Al2O3) is prepared using hydrochloric acid as a competitive adsorbent, and the amount of hydrochloric acid added can be 0wt% (no addition), 1wt%, 2wt% of the mass of the Sn-Al2O3 carrier; the Sn-Al2O3-containing carrier is immersed in the chloroplatinic acid impregnation solution for 4h, dried and calcined, then immersed in a potassium chloride solution for 4h, dried, calcined at 600°C for 4h, and finally reduced with H2 at a temperature of 580°C to obtain a Pt-Sn-K / Al2O3 catalyst.

[0010] Therefore, it is necessary to develop a process for preparing isoamylene from isoamylane by dehydrogenation, using a suitable catalyst, while achieving a single-pass conversion rate of isoamylane of 30% or more, preferably 30% or more, a selectivity of isoamylene of 90% or more, and a long-term operation effect of a catalyst regeneration cycle of 300 h or more, preferably 400 h or more, and further applying the process of isoamylene. SUMMARY

[0011] To solve the above technical problems, the present application provides a process for preparing isoamylene from isoamylane by dehydrogenation, a catalyst, and a process for preparing tertiary amyl alcohol.

[0012] The process for preparing isoamylene from isoamylane by dehydrogenation of the present application comprises an isoamylane dehydrogenation device and an isoamylene separation device.

[0013] The isoamylane dehydrogenation device comprises a dehydrogenation reactor group, and corresponding connecting pipelines, temperature measurement and control components, flow measurement and control components, and valves; the reactors of the dehydrogenation reactor group are fixed bed adiabatic reactors with axial or radial reaction structures, and are loaded with Pt-Sn-K / alumina type alkane dehydrogenation catalysts; the catalysts have hydrogen-assisted methanation performance under dehydrogenation reaction conditions containing CO2;

[0014] The dehydrogenation reactor group is set and operated in any of the following modes:

[0015] Mode A: 2-4 reactors are set, the loading volumes of the catalysts in each reactor are the same, and a heating furnace is arranged before the gas flow inlet of each reactor; the isoamylane dehydrogenation operation is carried out in parallel for each reactor; during the stable operation of isoamylane dehydrogenation, the inlet gas flow conditions of each reactor include: temperature 520-550℃, pressure ≤0.2 MPa (gauge pressure, the same below), isoamylane mass space velocity 1-2 h -1 -1, molar ratio of isoamylane to H2 1:(0.5-1.5), CO2 molar content 1.5-2.5%; the maximum temperature of the catalyst bed layer of each reactor is controlled to be below 550℃; the outlet gas flow temperature or the outlet section temperature of the catalyst bed layer of each reactor is controlled to be above 530℃;

[0016] Mode B: 2-3 reactor series are set in parallel, each reactor series comprises two reactors I and II connected in series, 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); a heating furnace is arranged before the gas flow inlet of each reactor; during the stable operation of isoamylane dehydrogenation, the inlet gas flow conditions of reactor I all include: temperature 520-550℃, pressure ≤0.2 MPa, isoamylane mass space velocity 1.5-3 h -1, the molar ratio of isopentane to H2 is 1:(0.5-1), and the molar content of CO2 is 1-2%; the inlet gas stream conditions of the reactor II include: temperature 510-540℃, and the molar content of CO2 is 1-2%; the highest temperature of the catalyst bed of each reactor is controlled to be below 550℃; and the outlet gas stream temperature or the temperature of the outlet section of the catalyst bed of the reactor II is controlled to be above 530℃;

[0017] Mode C: 2-3 reactor series are arranged in parallel, each of which includes three reactors III, IV and V connected in series, the reactor III is the first reactor, and the volume ratio of the catalysts in the reactors III, IV and V is 1:(0.8-1):(0.8-1); a heating furnace is arranged in front of the inlet of each reactor; during the stable operation of the isopentane dehydrogenation, the inlet gas stream conditions of the reactor III include: temperature 550-560℃, pressure ≤0.2MPa, the mass space velocity of isopentane 1.5-2.5h -1 , the molar ratio of isopentane to H2 is 1:(0.3-1):(1.5-3); the temperature of the inlet gas stream of the reactors IV and V is controlled to be 540-550℃; and the outlet gas stream temperature or the temperature of the outlet section of the catalyst bed of the reactor V is controlled to be above 530℃;

[0018] When the performance index of a reactor in mode A or a reactor series in mode B or C decreases to a specified range, the dehydrogenation operation is cut off and the catalyst is regenerated in situ, and after the regeneration is completed, the dehydrogenation operation is cut in.

[0019] In the process for preparing isoamylene by dehydrogenation of isopentane, 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 gas 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 goes to the isoamylene separation device. In the isopentane-rich gas stream and the hydrogen-rich gas stream, the content of the impurity components that are not conducive to the dehydrogenation process of isopentane and the reaction performance of the catalyst should be controlled to be below a certain content.

[0020] The present application is a process for preparing iso-pentene from iso-pentane dehydrogenation, the alkane dehydrogenation catalyst can be spherical or strip-shaped, Pt is uniformly distributed in the catalyst particles (i.e. the cross-section color is uniform) and is in a highly dispersed state at the beginning of each operating cycle, wherein the size of the spherical catalyst can be Φ2.5-3.0mm, and the size of the strip-shaped catalyst can be Φ(1.2-2)×(2-10)mm; it can be a spherical or strip-shaped catalyst H containing zirconium prepared on the basis of the improvement of the proportioning and method of CN103212411A: taking high-temperature zirconium-aluminum composite oxide containing 0.8-1.5wt% ZrO2 as the carrier, and loading components in terms of elements: Pt 0.2-0.4wt%, Sn 0.5-1wt%, K 0.65-1wt%, preferably taking spherical or strip-shaped catalyst HA with 0.35wt% Pt, 0.7wt% Sn, 0.83wt% K, and the carrier being high-temperature zirconium-aluminum composite oxide containing 1.1wt% ZrO2; it is found that the catalyst H, especially the catalyst HA, has good reaction performance for iso-pentane dehydrogenation to generate iso-pentene and matching CO2 hydrogenation to generate methane under the conditions of fixed bed adiabatic reactor and hydrogenation, such as under the setting and operating conditions of the dehydrogenation reactor group mode A, B, and has good reaction performance for iso-pentane dehydrogenation to generate iso-pentene under the setting and operating conditions of the dehydrogenation reactor group mode C (without CO2), and the catalytic activity is relatively stable, deactivation is very slow, and long-term operation effects of more than 30% iso-pentane single-pass conversion, more than 90% iso-pentene selectivity and more than 300h catalyst regeneration cycle can be obtained.

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

[0022] (H-1) taking spherical alumina with a pore volume of 0.5-0.7mL / g, a specific surface area of 100-130m 2 / g, and a size of Φ2.5-3.0mm or strip-shaped alumina with a size of Φ(1.2-2)×(2-10)mm as the carrier, uniformly applying zirconium oxychloride aqueous solution accounting for 80-100% of the saturated water absorption volume of the carrier and the required concentration, and placing in a closed container for 10-20h for homogenization treatment, drying, and then calcining at 880-930℃ for 2-4h, and cooling to obtain zirconium-modified alumina carrier;

[0023] (H-2) uniformly applying Pt-Sn-acid-K aqueous solution accounting for 80-90% of the saturated water absorption volume of the zirconium-modified alumina carrier obtained in step (H-1) and the required concentration and temperature of 20-30℃, placing in a closed container for 10-15h for homogenization treatment, drying, and then calcining at 500-600℃ for 2-5h, and cooling to obtain the alkane dehydrogenation catalyst;

[0024] The Pt-Sn-acid-K aqueous solution in step (H-2) is an aqueous solution with desired concentrations prepared from chloroplatinic acid, stannous chloride, hydrochloric acid, citric acid and potassium nitrate. The preparation method is as follows: chloroplatinic acid and hydrochloric acid are prepared into an aqueous solution, and stannous chloride, potassium nitrate and citric acid are added and dissolved and uniformly mixed to obtain the Pt-Sn-acid-K aqueous solution; 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 spherical or strip-shaped catalyst J containing magnesium: the carrier is magnesium-modified alumina containing 1-2 wt% of MgO, and the loading components are, in terms of elements, 0.2-0.4 wt% of Pt, 0.5-1 wt% of Sn and 0.65-1 wt% of K, and the catalyst J is prepared by the following steps:

[0026] (J-1) taking spherical alumina with a pore volume of 0.5-0.7 mL / g, a specific surface area of 100-130 m 2 / g and a size of Φ 2.5-3.0 mm or strip-shaped alumina with a size of Φ (1.2-2) × (2-10) mm as the carrier, uniformly applying an aqueous solution of magnesium bicarbonate with a saturated water absorption volume of 40-80% of the carrier and a desired concentration, placing for homogenization treatment for 5-20 h in a closed state, heating to 70-75 °C under the condition of flowing air and treating at a constant temperature for 3-5 h, then drying, and then calcining at 880-930 °C for 2-4 h and cooling to obtain the magnesium-modified alumina carrier;

[0027] (J-2) uniformly applying the Pt-Sn-EDTA-K aqueous solution with a saturated water absorption volume of 80-90% of the carrier and a desired concentration and a temperature of 40-70 °C to the magnesium-modified alumina carrier obtained in step (J-1), placing for homogenization treatment for 3-8 h in a closed and heat-insulated state, drying, then calcining at 500-600 °C for 2-5 h and cooling to obtain the alkane dehydrogenation catalyst;

[0028] The Pt-Sn-EDTA-K aqueous solution in step (J-2) is prepared by mixing diamminenitrito platinum (Pt(NH3)2(NO2)2), potassium stannate (K2SnO3), EDTA (ethylenediaminetetraacetic acid), nitric acid, potassium hydroxide, and potassium nitrite into an aqueous solution with a desired 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 value of 10-12, dissolved, and then EDTA (powder) and a nitric acid aqueous solution are added. The mixture is reacted (heated or stirred for a long time) until it is clear. A potassium hydroxide aqueous solution is added to adjust the pH value to 6.0-6.5. Potassium nitrite or an aqueous solution thereof is added, dissolved, and / or mixed uniformly to obtain a Sn-EDTA-K aqueous solution. 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°C, and a diamminenitrito platinum aqueous solution at 60-70°C is added and stirred uniformly to obtain a Pt-Sn-EDTA-K aqueous solution.

[0029] The catalyst J has good reaction performance for the dehydrogenation of isopentane to isopentene and the hydrogenation of matched CO2 to methane under the conditions of a fixed bed adiabatic reactor and hydrogenation, such as the settings and operating conditions of the dehydrogenation reactor group modes A and B. It has good reaction performance for the dehydrogenation of isopentane to isopentene under the settings and operating conditions of the dehydrogenation reactor group mode C (without the addition of CO2), and the overall catalytic activity of the reactor is relatively stable and deactivation is very slow. Under the settings and operating conditions of the reactor group modes A and B, the single-pass conversion rate of isopentane is more than 30% or even more than 35%, the selectivity of isopentene is more than 90%, and the long-term operation effect of the catalyst regeneration cycle is more than 400 h. Under the settings and operating conditions of the reactor group mode C, the single-pass conversion rate of isopentane is more than 30%, the selectivity of isopentene is more than 90%, and the long-term operation effect of the catalyst regeneration cycle is more than 600 h.

[0030] The present application is a process for preparing iso-pentene by dehydrogenation of iso-pentane. Under the setting and operating conditions of Mode A and B, the dehydrogenation reactor group of the iso-pentane dehydrogenation device can ultimately obtain a long-term operating effect of a single-pass conversion rate of iso-pentane of more than 30% or even more than 35%, a selectivity of iso-pentene of more than 90%, a methanation rate of CO2 of more than 80%, and a catalyst regeneration cycle of more than 300 hours, which indicates that the two catalysts have CO2 methanation performance matched with the dehydrogenation performance of iso-pentane, and effectively couple the two reactions, thereby significantly improving the dehydrogenation reaction conversion rate of iso-pentane by using the heat release of the CO2 methanation reaction. Since the dehydrogenation of iso-pentane is a significantly endothermic reaction, and the CO2 methanation is a strong exothermic reaction, if the CO2 methanation performance of the catalyst is significantly dominant, the maximum temperature of the catalyst bed will easily exceed 560°C, thereby resulting in a decrease in the selectivity of iso-pentene and a shortening of the catalyst regeneration cycle. During the dehydrogenation operation, the temperature and / or CO2 content of the inlet gas stream can be appropriately adjusted according to the overall performance of the reactor, such as using a lower temperature and / or CO2 content of the inlet gas stream at the beginning of the operating cycle when the overall performance of the catalyst is higher, and appropriately increasing the temperature and / or CO2 content of the inlet gas stream at the end of the operating cycle when the overall performance of the catalyst decreases, so as to obtain a relatively stable yield and output of iso-pentene. Since a relatively low inlet temperature of the gas stream is used, the overall temperature of the catalyst bed is lower than that of the prior art, thereby reducing the side reactions and prolonging the regeneration cycle. The methanation rate of CO2 can be calculated by the CO2 conversion rate and the distribution and content of the generated oxygen-containing compounds. During the operation of Mode A and B, the CO2 conversion rate is easily more than 90% or even 95%.

[0031] The magnesium-containing catalyst J has better overall performance and greater operating flexibility than the zirconium-containing catalyst H during the dehydrogenation application before the first regeneration of the dehydrogenation reactor group, which indicates that the dispersion and combination state of Pt and Sn reach a high level during the operation of the first regeneration cycle, and the magnesium modification can basically eliminate the long-term and special effects of the surface acidity of the alumina carrier. However, the operating conditions and effects of the second operating cycle and thereafter depend on the regeneration method. In the process for preparing iso-pentene by dehydrogenation of iso-pentane, the reactor needs to be reduced after being filled with new catalyst and before being put into dehydrogenation operation, and the last step of the regeneration process is also a reduction operation. The reduction operation conditions are as follows: using a hydrogen stream with a H2 molar content of more than 98% and an olefin molar content of less than 0.1%, an airspeed of 500-1000 h -1 , a pressure of 0-0.2 MPa, a reduction time of 3-6 hours at a temperature of more than 480°C, and a maximum bed temperature of not more than 520°C.

[0032] The in-situ regeneration process of the reactor or series filled with catalyst H or catalyst J, when the overall reaction performance index falls to a specified range, or after the dehydrogenation operation is cut off, can be the same as the in-situ regeneration process of the Pt-Sn-K / alumina catalyst for preparing isobutene from isobutane in the related prior art, or substantially the same as the ex-situ regeneration process of the semi-regenerated platinum reforming catalyst in a refinery, including carbon burning, oxychlorination, and reduction operations. In the carbon burning operation, an oxygen-containing gas stream with an O2molar content of 0.5-0.8% is used, which is obtained by mixing air into a nitrogen stream (N2molar content of more than 99%), the space velocity is 500-2000h-1, the inlet temperature is 450-490℃, the pressure is 0-0.1 MPa, the initial bed temperature is 450-500℃, the maximum bed temperature is not more than 520℃, and the difference between the O2molar content of the outlet gas stream and that of the inlet gas stream is reduced to less than 0.1%, and then the process is extended for 2-4h. In the oxychlorination operation, an oxygen-containing gas stream with an O2molar content of 4-15% is used, which is obtained by mixing air into a nitrogen stream, the space velocity is 500-2000h-1, the pressure is 0.02-0.2 MPa, and the initial bed temperature is 480-520℃. Water or steam is first mixed into the oxygen-containing gas stream, and chlorides such as dichloroethane, trichloroethane, carbon tetrachloride, or tetrachloroethylene are also mixed into the oxygen-containing gas stream, or a chloroacetic acid aqueous solution (containing water and organic chlorine) with a mass concentration of 40-50% is directly mixed into the oxygen-containing gas stream before a heating furnace or a heater, the inlet temperature of the gas stream is controlled to be 500-520℃, the amount of chlorides mixed into the reactor 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 process is performed for 4-6h, and the outlet gas stream temperature of the reactor is higher than 515℃. After the oxychlorination operation is completed, the reactor is purged with a high-purity nitrogen stream (N2molar content of more than 99.9%) at a temperature of 480-500℃, the space velocity is 500-2000h-1, and the O2molar content of the outlet gas stream is reduced to less than 0.2%. Then, the reduction of catalyst H or catalyst J is performed according to the above reduction operation conditions. After the in-situ regeneration including the oxychlorination operation, the carbon deposition of catalyst H or catalyst J is substantially burned off, the Pt particles are redispersed, and the isopentane dehydrogenation reaction performance and the CO2hydrogenation to methane performance are substantially restored. The overall performance of catalyst H can be substantially restored to the effect level of the first running cycle, and the overall performance of catalyst J is slightly lower than that of the first running cycle without chlorine, indicating that the dispersion and combination state of Pt and Sn are restored to a high level, but the overall performance of catalyst J is still higher than that of catalyst H. -1 -1 -1

[0033] ​​​The reactor or series filled with catalyst H or the reactor or series filled with catalyst J, after being treated by one or more in-situ regeneration processes under the presence of chlorine, can basically restore the reaction performance, and when isopentane dehydrogenation is carried out under the setting and operating conditions of the dehydrogenation reactor group mode A, B or C, the long-term operation effect of isopentane single-pass conversion rate of 30% or even 35% or more, isopentene selectivity of 90% or more, methanation rate of 80% or more in the presence of CO2, and catalyst regeneration cycle of 300h or more can be obtained; wherein the reactor or series filled with catalyst J, after being treated by one or more in-situ regeneration processes under the presence of chlorine, when isopentane dehydrogenation is carried out under the setting and operating conditions of the dehydrogenation reactor group mode A or B, the catalyst regeneration cycle can be maintained for 500h or more.

[0034] The reactor or series filled with catalyst J can also be subjected to in-situ regeneration process under the absence of chlorine when the overall reaction performance index decreases to a specified range or after the dehydrogenation operation is cut off, including carbon burning, oxidation and reduction operations. When the carbon burning operation is carried out, an oxygen-containing gas stream with O2 molar content of 0.5-0.8% obtained by mixing dry air stream (such as the dew point being lower than -20℃) into nitrogen stream is used, the space velocity is 500-2000h -1 -1, the inlet temperature is 450-490℃, the pressure is 0-0.1MPa, and the bed initial temperature is 450-500℃; during the carbon burning process, the highest temperature of the bed is controlled to be not more than 520℃, and then the difference between the O2 molar content in the outlet gas stream and the inlet gas stream is reduced to be less than 0.1%, and then extended for 2-4h. When the oxidation operation is carried out, an oxygen-containing gas stream with O2 molar content of 10-20% obtained by mixing dry air stream into nitrogen stream or directly using dry air stream is used, the space velocity is 500-2000h -1 -1, the pressure is 0.02-0.2MPa, the bed initial temperature is 480-520℃, the gas inlet temperature is 500-520℃, and the oxidation treatment is carried out for 6-10h. After the oxidation operation is completed, the reactor is replaced by nitrogen stream and high-purity nitrogen stream with temperature of 480-500℃ for purging, the space velocity is 500-2000h -1, the molar content of O2 in the outlet gas stream is reduced to below 0.2%, and then the catalyst J is reduced again under the above reduction operation conditions. After the reactor or series loaded with the catalyst J is treated by the in-situ regeneration under the chlorine-free conditions, when the dehydrogenation of isopentane is carried out under the setting and operation conditions of the dehydrogenation reactor group mode A, B or C, the long-term operation effect of the isopentane single-pass conversion rate of 30% or even above 35%, the isopentene selectivity of 90% or above, the methanation rate of 80% or above in the presence of CO2, and the catalyst regeneration cycle of 300h or even above 450h can be obtained; in particular, the problem of isomerization side reaction to produce n-pentane is less than that of the catalyst J after the in-situ regeneration under the oxygen-chlorination operation. After the catalyst J is treated by the in-situ regeneration under the chlorine-free conditions, the carbon deposition is basically burned off, and the isopentane dehydrogenation reaction performance and the CO2 hydrogen-assisted methanation performance are greatly restored (slightly lower than the effect level in the first operation cycle), which indicates that the dispersion and combination state of Pt and Sn can be restored to the required degree, and the overall basicity of the catalyst and the anti-carbon deposition effect can also be basically maintained. The in-situ regeneration treatment under the chlorine-free conditions reduces the chlorine resistance requirement of the equipment such as the reactor, the pipeline, the control components, and the detection components, can avoid the corrosion problem of chlorides, and can also avoid the generation of KCl. The catalyst still exhibits the overall basicity and the good anti-carbon deposition ability, which to some extent makes up for the defect of the limited Pt re-dispersion effect during the regeneration treatment under the chlorine-free conditions.

[0035] The present application is prepared by isopentane dehydrogenation process, the isopentene separation device includes cooling separation system and pressure swing adsorption separation system;The isopentane dehydrogenation gas stream Q after cooling in heat exchanger T1 is further cooled and pressurized, such as cooling to (-20 to-10) ℃ and pressurizing to 0.5-1 MPa, separating water and obtaining gas stream G0 mainly containing H2, CH4 and a small amount of CO2, ethane, ethylene, and hydrocarbon liquid stream L0;The hydrocarbon containing liquid stream L0 enters the light removal column D0 (theoretical tray number 30-36), and the overhead is separated into a small amount of low boiling gas stream G-1 containing C4 and below components and part of 3-methyl-1-butene, and the bottom is separated into a hydrocarbon containing liquid stream L. The pressure swing adsorption unit includes PSA-H2 unit for separating hydrogen from gas stream G0, and one or more sets of PSA-CH unit for separating hydrocarbons, wherein the PSA-H2 unit adopts conventional technology of pressure swing adsorption hydrogenation adsorbent, equipment configuration and separation operation process, and the PSA-CH unit adopts known technology of hydrocarbon separation adsorbent, equipment configuration and separation operation process. Gas stream G0 is separated in PSA-H2 unit to obtain hydrogen stream G-H2 with a molar purity of more than 98% or even more than 99%, almost no unsaturated hydrocarbons, part of which is recycled to isopentane dehydrogenation device, and the rest is externally supplied;Also obtain low hydrogen stream G-2 containing CH4, ethane, ethylene, CO2 for heating furnace as fuel gas. The PSA-CH unit is preferably provided with two sets of hydrocarbon separation units, including alkylene separation unit PSA-CH-1 and alkane separation unit PSA-CH-2;The gasified hydrocarbon containing liquid stream L is separated by PSA-CH-1 unit to obtain alkane gas stream GL-1, isopentene gas stream Z-2 and hetero-hydrocarbon gas stream Z-3 three parts, the obtained isopentene gas stream Z-2 is easy to control to isopentene molar purity of more than 98% or even more than 99%, diene hydrocarbon basically enters the hetero-hydrocarbon gas stream Z-3 and the amount of the hetero-hydrocarbon gas stream Z-3 is very small, which can be used as fuel gas or liquefied and collected for external sale;The alkane gas stream GL-1 is further separated by PSA-CH-2 unit to obtain isopentane gas stream Y-1 and n-pentane gas stream Z-1 two parts, the isopentane molar purity of the obtained isopentane gas stream Y-1 and the n-pentane molar purity of the n-pentane gas stream Z-1 are all easy to control to more than 96% or even more than 98%, the isopentane gas stream Y-1 is recycled to the isopentane dehydrogenation device, and the n-pentane gas stream Z-1 is liquefied and collected for external supply or converted into isopentane by isomerization device and then used in isopentane dehydrogenation device. The obtained isopentene gas stream Z-2 can be further liquefied to obtain isopentene liquid stream Z-2L. The light removal column D0 overhead is separated into a small amount of low boiling gas stream G-1 containing C4 and below components and 3-methyl-1-butene, which is also used as fuel gas in the heating furnace.

[0036] The process operation of the cooling separation system can include: cooling the gas stream Q to 20-60°C (such as through a heat exchanger cooled by circulating water), increasing the pressure to 0.5-1 MPa, cooling to 20-60°C, and then successively cooling to 1-10°C and (-20 to -10) °C and carrying out water separation and ice removal operations, and cooling to (-20 to -10) °C to obtain a gas stream G0 and a hydrocarbon-containing liquid stream L0; the two-step cooling to 1-10°C and (-20 to -10) °C uses a cooling heat exchanger with water separation function, and when cooled to 1-10°C, condensed water is separated out (water produced by the CO2 hydrogen-assisted methanation reaction when the dehydrogenation reactor group is set and operated in mode A, B), and during the (-20 to -10) °C cooling process, the ice on the heat exchange surface can be removed by using multiple cooling heat exchangers and carrying out ice removal operations (short-term shutdown of the cooling medium circulation) in rotation to maintain long-term heat exchange capacity and deeply remove water in the stream; when mode C, a large amount of water vapor is added to the gas stream at the inlet of reactor III, and after the gas stream Q is cooled to 20-60°C before and after being increased to 0.5-1 MPa, condensed water is continuously separated out through an oil-water separator; the gas-liquid mixed stream cooled to (-20 to -10) °C is then separated into a gas stream G0 mainly containing H2, CH4 and a small amount of CO2, ethane and ethylene, and a hydrocarbon-containing liquid stream L0.

[0037] The adsorbent used in each adsorption tower of the PSA-CH-1 unit has an adsorption selectivity for isopentene / pentane in the gasified stream L containing hydrocarbons of more than 3, preferably more than 4, under its operating conditions (such as 40-100°C, absolute pressure 30-1000 kPa), i.e. more easily adsorbing isopentene, and the pentane is the total amount of n-pentane and isopentane; similarly, the adsorbent used in each adsorption tower of the PSA-CH-2 unit has an adsorption selectivity for n-pentane / isopentane in the alkane gas stream GL-1 of more than 6 under its operating conditions (such as 100-160°C, absolute pressure 40-1000 kPa), and the separation principle is the shape selection of the adsorbent for n-pentane / isopentane, and the n-pentane molecule size is slightly smaller, which can enter the inner pores of the adsorbent microparticles (with a larger surface area) and be adsorbed, while the isopentane cannot enter the inner pores of the adsorbent microparticles in large amounts due to its slightly larger molecular size and is discharged.

[0038] The present application is a process for preparing isoamylene from isoamylane dehydrogenation, wherein the isoamylene separation device comprises a cooling separation system, a pressure swing adsorption separation system, and a rectification system. The rectification system comprises a rectification column D1 (theoretical plate number 50-60) to separate the hydrocarbon-containing liquid stream L from the bottom of the light-removing column D0. With the rectification system, the PSA-CH unit can only have the alkylene separation unit PSA-CH-1 (without the PSA-CH-2 unit). The rectification column D1 separates the isoamylane liquid stream Y-2 from the top and the isoamylene-rich liquid stream Z containing n-pentane and dienes from the bottom. The isoamylene-rich liquid stream Z is gasified and then separated into the n-pentane gas stream Z-11, the isoamylene gas stream Z-21, and the hetero-hydrocarbon gas stream Z-31 containing dienes by the PSA-CH-1 unit. The isoamylene gas stream Z-21 and the isoamylene liquid stream Z-21L obtained by liquefaction are the target products of isoamylane dehydrogenation, and the molar purity can be controlled to more than 98% or even more than 99%. The n-pentane gas stream Z-11 and the hetero-hydrocarbon gas stream Z-31 containing dienes can be further utilized or liquefied for sale.

[0039] The present application also provides a process for preparing tertiary amyl alcohol. The isoamylene liquid stream Z-2L or Z-21L from the isoamylene separation device in the isoamylane dehydrogenation process and water are used as raw materials, mixed with the recycled residual isoamylene liquid stream X-1, the recycled water stream X-3 or the isoamyl alcohol-containing aqueous solution stream X-3, and the circulating solvent liquid stream X-4. The obtained liquid stream X is subjected to liquid-phase reaction for isoamylene hydration to produce tertiary amyl alcohol in a fixed-bed hydration reactor filled with macroporous sulfonic acid-type cation resin catalyst (hydrogen type, such as the product JQE-02 of the company). The solvent contained in the circulating solvent liquid stream X-4 can be isopropyl alcohol, tertiary butyl alcohol, or ethylene glycol monoethyl ether, and is preferably ethylene glycol monobutyl ether with a higher boiling point. The liquid stream obtained by hydration is separated into the residual isoamylene liquid stream X-1, the isoamyl alcohol and water miscible liquid stream X-20, and the circulating solvent liquid stream X-4 by passing through a light-removing column (theoretical plate number 18-25) and a rectification column (theoretical plate number 40-50) in sequence. The isoamyl alcohol and water miscible liquid stream X-20 is separated into the isoamyl alcohol liquid stream X-2 and the water stream X-3 or the isoamyl alcohol-containing aqueous solution stream X-3 by known separation methods such as azeotropic rectification, extractive rectification, or extractive-azeotropic rectification. The residual isoamylene liquid stream X-1, the water stream X-3 or the isoamyl alcohol-containing aqueous solution stream X-3, and the circulating solvent liquid stream X-4 are all recycled for hydration reaction. The operating conditions of the hydration reactor include that the molar ratio of water, isoamylene, 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, the feed temperature is 40-65°C, the pressure is 0.5-1.0 MPa, and the discharge temperature is 65-80°C; the hydration reactor is preferably provided with an external circulation system, 40-70% of the volume or mass flow of the discharge liquid stream is recycled as a circulation material liquid stream X-5 and mixed into the liquid stream X before being fed into the hydration reactor, and / or a heat-removal temperature-controlling component such as a coil pipe through which 60-65°C circulating water is circulated is arranged in the catalyst bed; the total content (molar content) of 2-methyl-2-butene and 2-methyl-1-butene in the residual isopentene liquid stream X-1 (containing water) can be maintained at 90% or more by sending part of the flow of the residual isopentene liquid stream X-1 back to the cooling separation system of the isopentene separation device, and combining and processing the residual isopentene liquid stream X-1 with the hydrocarbon-containing liquid stream L0, so that the content of 3-methyl-1-butene, alkanes and other components that do not react with water brought in by the isopentene liquid stream Z-2L or Z-21L is controlled to be less than 10%; the extractant used in the extractive rectification method or the extractive-azeotropic rectification method can be cyclohexane or n-hexane. The hydration reactor can achieve and ensure a reaction effect of isopentene conversion ≥40% and tert-amyl alcohol selectivity ≥98% during a long-term operation of more than one year; the separation yield of tert-amyl alcohol in the subsequent separation process of the hydration reaction can be ensured to be more than 98%; the content of tert-amyl alcohol in the obtained tert-amyl alcohol liquid stream X-2 is more than 99.5wt% or even 99.8wt%, and the water content is less than 0.1wt%, so that the obtained tert-amyl alcohol liquid stream X-2 can be directly packed, stored and transported as a tert-amyl alcohol product after being cooled. In the preparation process of the tert-amyl alcohol, 2-methyl-2-butene and 2-methyl-1-butene in the three isomers of isopentene react with water to generate tert-amyl alcohol, while 3-methyl-1-butene does not react with water. The use of ethylene glycol monobutyl ether as the isopentene hydration reaction solvent has the advantages of stable performance, good reaction effect, and the fact that the heavy components (mainly isopentene polymers and fusel alcohols) generated in the hydration reaction and the separation process are more likely to enter the circulation solvent liquid stream X-4, so that the fractionation and purification treatment of the circulation solvent liquid stream X-4 on a regular basis helps to maintain and prolong the reaction performance and service life of the macroporous sulfonic acid type cation resin catalyst in the hydration reactor, and is conducive to obtaining a tert-amyl alcohol product with a purity of more than 99.8wt%. The normal boiling point of tert-amyl alcohol is 102°C, and the normal boiling point of ethylene glycol monobutyl ether is 171°C. The isopentene hydration reaction process, the separation process of the hydration reaction discharge liquid, and the process control and effect are known in the art.The mixture of tertiary amyl alcohol and water stream X-20 is separated by using the azeotrope of tertiary amyl alcohol and water (azeotrope temperature 87.35℃ under normal pressure, azeotrope composition: 72.5wt% of tertiary amyl alcohol and 27.5wt% of water) and the immiscibility of tertiary amyl alcohol and water at a temperature of 50-75℃, which can be easily separated into the upper tertiary amyl alcohol phase containing water and the lower water phase containing tertiary amyl alcohol. For details, please refer to the relevant records and research in "Purification of Tertiary Amyl Alcohol by Azeotropic Rectification" (Gaoyuanming, [J] Sichuan Chemical Industry, 1995, 4, 29-31) and "Process Optimization Research of Isopentene and Tertiary Amyl Alcohol" (Wang Shaoheng, Master Thesis of Dalian University of Technology, 2024).

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

[0041] 1. The process for preparing isopentene from isopentane by dehydrogenation, when the isopentane dehydrogenation device is operated under the setting and operating conditions of modes A, B and C, can obtain a long-term operation effect of isopentane single-pass conversion rate of 30% or even more than 35%, isopentene selectivity of 90% or more, and catalyst regeneration cycle of 300h or more, even 600h or more; the isopentene separation device can obtain isopentene with a molar purity of 98% or even more than 99%, and hydrogen.

[0042] 2. An effective process for preparing tertiary amyl alcohol from isopentane is provided, which provides a potential and better choice for the application of isopentane; tertiary amyl alcohol is a widely used organic raw material and solvent with a relatively high market price.

[0043] 3. The process for preparing isopentene from isopentane by dehydrogenation and the process for preparing tertiary amyl alcohol from isopentane are superior to the prior art as a whole and have certain application prospects.

[0044] It is believed that the main principle of the process for preparing isopentene from isopentane by dehydrogenation, which can achieve the above beneficial effects, includes:

[0045] 4. The catalyst H, although prepared on the basis of the proportioning and method of CN103212411A, has a specific surface area of 100-130m 2 / g, and the zirconium-containing alumina carrier is prepared by impregnating the alumina carrier with an aqueous solution of zirconyl chloride, and finally the catalyst H with better isopentane dehydrogenation performance is prepared. In the method of CN103212411A, the zirconium-containing alumina carrier is prepared by impregnating the boehmite precursor of alumina (a crystal form of aluminum hydroxide, which has not yet formed alumina crystal grains and pores in the crystal grains) with an aqueous solution of zirconyl chloride, and the catalyst prepared by further loading Pt, Sn and K in the obtained zirconium-containing alumina carrier, the dispersion state of Zr and the combination state with Al2O3 will be significantly different compared with those in the catalyst H, which should be the main reason for the performance difference between the two in the isopentane dehydrogenation reaction. The specific surface area of the spherical or strip-shaped alumina carrier is 100-130 m 2 / g, and the zirconium-containing alumina carrier is prepared by impregnating the alumina carrier with an aqueous solution of zirconyl chloride, and finally the catalyst H with better isopentane dehydrogenation performance is prepared. In the method of CN103212411A, the zirconium-containing alumina carrier is prepared by impregnating the boehmite precursor of alumina (a crystal form of aluminum hydroxide, which has not yet formed alumina crystal grains and pores in the crystal grains) with an aqueous solution of zirconyl chloride, and the catalyst prepared by further loading Pt, Sn and K in the obtained zirconium-containing alumina carrier, the dispersion state of Zr and the combination state with Al2O3 will be significantly different compared with those in the catalyst H, which should be the main reason for the performance difference between the two in the isopentane dehydrogenation reaction. The specific surface area of the spherical or strip-shaped alumina carrier is 100-130 m

[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 crystallization and uniform dispersion of the EDTA complex and the diamminonitrosoplatinum in the pores of the magnesium-modified alumina carrier, the uniform radial distribution of the Pt-Sn components in the spherical or strip-shaped catalyst J, the uniform color caused by the Pt component, and the higher dispersion of the Pt and Sn than in the related prior art catalysts, including the catalyst H containing zirconium, or the better combined dispersion of the Pt and Sn, make the catalyst J not only have better overall isopentane dehydrogenation performance, but also have CO2 hydrogenation to methane performance that matches the isopentane dehydrogenation performance. The diamminonitrosoplatinum and potassium stannate used are both reagents or products that are easy to purchase and have stable quality; the potassium stannate is a tetravalent tin salt, and the process of oxidizing divalent tin to tetravalent tin is omitted.

[0047] 6. In the preparation of the catalyst J containing magnesium, the use of the magnesium-modified alumina carrier substantially eliminates the surface acidity of the alumina carrier, and under the basic conditions of the K loading amount, the catalyst has lower isomerization activity of isopentane to n-pentane, and extremely low cracking, carbonization, and deep dehydrogenation to diene activity. The catalyst J has good reaction performance of isopentane dehydrogenation to isopentene and matching CO2 hydrogenation to methane under fixed bed adiabatic reactor and hydrogenation conditions, such as under the setting and operating conditions of the dehydrogenation reactor group modes A and B, and good reaction performance of isopentane dehydrogenation to isopentene under the setting and operating conditions of the dehydrogenation reactor group mode C (without CO2), and the overall catalytic activity of the reactor is relatively stable and deactivates very slowly; under the setting and operating conditions of the reactor group modes A and B, long-term operation effects of isopentane single-pass conversion rate of 30% or even 35% or more, isopentene selectivity of 90% or more, and catalyst regeneration cycle of 400h or more can be obtained; under the setting and operating conditions of the reactor group mode C, long-term operation effects of isopentane single-pass conversion rate of 30% or more, isopentene selectivity of 90% or more, and catalyst regeneration cycle of 600h or more can be obtained.

[0048] 7. Under the setting and operating conditions of Mode A, B, the isopentane dehydrogenation reactor group can obtain the long-term operation effects of isopentane single-pass conversion rate of 30% or even 35% above, isopentene selectivity of 90% above, CO2 methanation rate of 80% above and catalyst regeneration cycle of 300h above, when the catalyst H containing zirconium or catalyst J containing magnesium is used, which shows that both of the catalysts have the CO2 methanation performance matched with the isopentane dehydrogenation performance, and both of the catalysts effectively couple the two reactions, so as to use the heat release of the CO2 methanation reaction to significantly improve the isopentane dehydrogenation reaction conversion rate. Since the isopentane dehydrogenation is a significantly endothermic reaction and the CO2 methanation is a strong exothermic reaction, if the CO2 methanation performance of the catalyst occupies a significant advantage, the highest temperature of the catalyst bed will easily exceed 560℃, thereby resulting in the reduction of isopentene selectivity and the shortening of catalyst regeneration cycle. During the dehydrogenation operation, the temperature and / or CO2 content of the inlet gas stream can be appropriately adjusted according to the overall performance of the reactor, such as using a lower temperature and / or CO2 content of the inlet gas stream at the initial stage of the operation cycle and when the overall performance of the catalyst is higher, and appropriately increasing the temperature and / or CO2 content of the inlet gas stream at the end of the operation cycle and when the overall performance of the catalyst is reduced, so as to obtain a relatively stable isopentene yield and output; since a relatively low inlet gas stream temperature is used, the overall catalyst bed temperature is also lower than that of the prior art, so that the side reactions are reduced and the regeneration cycle is prolonged. The CO2 methanation rate can be calculated by the CO2 conversion rate and the distribution and content of the generated oxygen-containing compounds, and the CO2 conversion rate is easily more than 90% or even 95% during the operation of Mode A, B, but the dehydrogenation reactor outlet gas does not contain methanol.

[0049] 8. During the stable operation of each isopentane dehydrogenation reactor, the highest temperature of the catalyst bed is controlled to be 550℃ or lower under the setting and operating conditions of Mode A, B, and is controlled to be 560℃ or lower for reactor III and 550℃ or lower for reactors IV and V under the setting and operating conditions of Mode C, which is a key factor to obtain the isopentene selectivity of 90% or above and the catalyst regeneration cycle of 300h or even 500h or above; the temperature of the outlet gas stream or the temperature of the outlet section of the catalyst bed of each reactor of Mode A, and the outlet gas stream temperature of the rear reactor of Mode B, C is controlled to be 530℃ or above, which is a key factor to obtain the isopentane conversion rate of 30% or above; the concentration of the main dehydrogenation product 2-methyl-2-butene is close to the equilibrium concentration between isopentane and 2-methyl-2-butene under the specific reaction conditions of the outlet section temperature and composition. BRIEF DESCRIPTION OF DRAWINGS

[0050] ATTACHMENT Figure 1 It is a process flow diagram of Mode B in the process of preparing isopentene from isopentane dehydrogenation in Example 1 below; wherein each PSA unit only shows the outline of the material in and out relationship. DETAILED DESCRIPTION

[0051] The technical solutions of the present application are described and explained in detail below in combination with examples.

[0052] Example 1

[0053] According to the performance evaluation of the dehydrogenation catalyst HA (containing 0.35wt% Pt, 0.7wt% Sn, 0.83wt% K, and a carrier of spherical high-temperature zirconium-aluminum composite oxide containing ZrO21.1wt%) in Catalyst Evaluation Example 1 below and the performance of the dehydrogenation catalyst JB (containing Pt 0.3wt%, Sn 0.7wt%, K 0.83wt%, and a carrier of magnesium-modified alumina containing MgO 1.5wt%) in Catalyst Evaluation Example 2 below, the process of preparing isoamylene from isoamylane by dehydrogenation in this example is designed, including an isoamylane dehydrogenation device and an isoamylene separation device.

[0054] The isoamylane dehydrogenation device includes a dehydrogenation reactor group, and corresponding connecting pipelines, temperature measurement and control components, flow measurement 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 catalyst HA or catalyst JB;

[0055] The dehydrogenation reactor group is set and operated in any of the following modes:

[0056] Mode A: three reactors are set, the loading volume of catalyst in each reactor is the same, and a heating furnace is arranged before the gas flow inlet of each reactor; the isoamylane dehydrogenation operation is carried out in parallel for each reactor; during the stable operation of isoamylane dehydrogenation, the inlet gas flow conditions of each reactor include: temperature 520-550℃, pressure ≤0.2MPa, isoamylane mass space velocity 1-2h -1 , molar ratio of isoamylane to H2 1:(0.5-1.5), CO2 molar content 1.5-2.5%; the maximum temperature of the catalyst bed of each reactor is controlled to be below 550℃; the outlet gas flow temperature or the outlet section temperature of the catalyst bed of each reactor is controlled to be above 530℃;

[0057] Mode B: two parallel reactor series are set, each reactor series includes two reactors I and II connected in series, reactor I is the former reactor, and the loading volume ratio of catalyst in reactors I and II is 1:1; a heating furnace is arranged before the gas flow inlet of each reactor; during the stable operation of isoamylane dehydrogenation, the inlet gas flow conditions of reactor I include: temperature 520-550℃, pressure ≤0.2MPa, isoamylane mass space velocity 1.5-3h -1, the molar ratio of isopentane to H2 is 1:(0.5-1), and the molar content of CO2 is 1-2%; the inlet gas stream of reactor II has a temperature of 510-540 ℃ and a molar content of CO2 of 1-2%; the maximum temperature of the catalyst bed of each reactor is controlled to be below 550 ℃; and the outlet gas stream of reactor II or the temperature of the outlet section of the catalyst bed of reactor II is controlled to be above 530 ℃;

[0058] Mode C: two reactor series in parallel are provided, each reactor series including three reactors III, IV and V connected in series, wherein reactor III is the first reactor, and the volume ratio of the catalysts in reactors III, IV and V is 1:1:1; a heating furnace is provided before the inlet of each reactor; during stable operation of isopentane dehydrogenation, the inlet gas stream of reactor III has a temperature of 550-560 ℃, a pressure of ≤0.2 MPa, a mass space velocity of isopentane of 1.5-2.5 h -1 , the molar ratio of isopentane to H2 is 1:(0.3-1):(1.5-3); the inlet gas stream of reactors IV and V has a temperature of 540-550 ℃; and the outlet gas stream of reactor V or the temperature of the outlet section of the catalyst bed of reactor V is controlled to be above 530 ℃;

[0059] When the performance of a reactor in mode A or a reactor series in mode B or C drops to a specified range, the dehydrogenation operation is cut off, and in-situ regeneration of the catalyst is performed, and after regeneration, the dehydrogenation operation is cut in; the isopentane dehydrogenation device is further provided with a heat exchanger T1; a mixed stream M of the raw isopentane liquid stream, the reused isopentane gas stream Y, the reused hydrogen gas stream G-H2 and the carbon dioxide gas stream, and a post-reaction gas stream P from the dehydrogenation reactor group, are countercurrently exchanged in the heat exchanger T1 to obtain a pre-reaction gas stream N and a cooled gas stream Q, the pre-reaction gas stream N enters the dehydrogenation reactor group, and the cooled gas stream Q goes to the isopentene separation device;

[0060] The isopentene separation device comprises a cooling separation system and a pressure swing adsorption separation system; the isopentane dehydrogenation gas stream Q cooled by the heat exchanger T1 is further cooled and pressurized in the cooling separation system, cooled to (-20 to -10) ℃ and pressurized to 0.5-1 MPa, separated from water and obtained 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 component removal column D0 (a packed column with about 10 theoretical plates), a small amount of low boiling gas stream G-1 containing carbon four and below components is separated at the top, and a hydrocarbon-containing liquid stream L is separated at the bottom. The pressure swing adsorption unit comprises a PSA-H2 unit for separating hydrogen from the gas stream G0, and one or more PSA-CH units for separating hydrocarbons, wherein the PSA-H2 unit adopts a conventional technology of pressure swing adsorption hydrogen separation adsorbent, equipment configuration and separation operation process, and the PSA-CH unit adopts a known technology of hydrocarbon separation adsorbent, equipment configuration and pressure swing adsorption separation operation process. The hydrogen G-H2 with a molar purity of more than 98% or even more than 99% and almost no unsaturated hydrocarbons is separated from the gas stream G0 in the PSA-H2 unit, part of which is recycled to the isopentane dehydrogenation device as a dehydrogenation reaction ingredient, and the rest is externally supplied; the low hydrogen stream G-2 containing CH4, ethane, ethylene and CO2 separated by the PSA-H2 unit is used as fuel gas in the heating furnace. The PSA-CH unit is provided with two sets of hydrocarbon separation units, including an alkylene separation unit PSA-CH-1 and an alkane separation unit PSA-CH-2; the gasified hydrocarbon-containing liquid stream L is separated by the PSA-CH-1 unit to obtain an alkane gas stream GL-1 and an isopentene gas stream Z-2 and a hetero-hydrocarbon gas stream Z-3 in three parts, the obtained isopentene gas stream Z-2 can be controlled to have a molar purity of isopentene of more than 98% or even more than 99%, dienes basically enter the hetero-hydrocarbon gas stream Z-3 and the amount of the hetero-hydrocarbon gas stream Z-3 is very small, which can be used as fuel gas in the heating furnace or liquefied and collected for external sale; the alkane gas stream GL-1 is further separated by the PSA-CH-2 unit to obtain an isopentane gas stream Y-1 and a n-pentane gas stream Z-1 in two parts, the isopentane molar purity of the obtained isopentane gas stream Y-1 and the n-pentane molar purity of the n-pentane gas stream Z-1 are both easily controlled to be more than 96% or even more than 98%, the isopentane gas stream Y-1 is recycled to the isopentane dehydrogenation device as part of the dehydrogenation reaction ingredient, and the n-pentane gas stream Z-1 is liquefied and collected for external supply or converted into isopentane by an isomerization device and then used in the isopentane dehydrogenation device. The obtained isopentene gas stream Z-2 is further liquefied to obtain an isopentene liquid stream Z-2L. The small amount of low boiling gas stream G-1 containing carbon four and below components separated at the top of the light component removal column D0 is also used as fuel gas in the heating furnace. Part of the configuration and operation effect of the isopentene separation device is simulated by Aspen Plus software.

[0061] The process operation of the cooling separation system can include: cooling the gas stream Q to 20-60℃ (by a heat exchanger cooled by circulating water), increasing the pressure to 0.5-1 MPa, cooling to 20-60℃, and then successively cooling to 1-10℃ and (-20 to -10)℃ and carrying out water separation and ice removal operation, and cooling to (-20 to -10)℃ to obtain the gas stream G0 and the hydrocarbon-containing liquid stream L0 after gas-liquid separation; the two-step cooling to 1-10℃ and (-20 to -10)℃ uses a cooling heat exchanger with water separation function, and the condensed water is separated when cooled to 1-10℃ (the water produced by the CO2 hydrogen-assisted methanation reaction in the setting and operation mode A and B of the dehydrogenation reactor group), and the ice on the heat exchange surface is melted and removed by using multiple cooling heat exchangers and carrying out ice removal operation (short-term shutdown of the circulation of the cooling medium) in turn and at regular time intervals during the cooling process to (-20 to -10)℃, so as to maintain long-term heat exchange capacity and deeply remove the water in the stream; in mode C, a large amount of water vapor is added to the gas stream at the inlet of the reactor III, and the condensed water is continuously separated by an oil-water separator after the gas stream Q is cooled to 20-60℃ before and after being increased to 0.5-1 MPa; the gas-liquid mixed stream cooled to (-20 to -10)℃ is separated by a gas-liquid separator to obtain the gas stream G0 mainly containing H2, CH4 and a small amount of CO2, ethane and ethylene, and the hydrocarbon-containing liquid stream L0.

[0062] The adsorbent of the alkene used in each adsorption tower of the PSA-CH-1 unit has an adsorption selectivity of 3.3 for isoamylene / pentane in the gasification stream L containing hydrocarbons under the operating conditions of 40-100℃ and an absolute pressure of 30-1000kPa, and the pentane is the total amount of n-pentane and isoamylene; the adsorbent of the alkane used in each adsorption tower of the PSA-CH-2 unit has an adsorption selectivity of 6.5 for n-pentane / isoamylene in the alkane gas stream GL-1 under the operating conditions of 100-160℃ and an absolute pressure of 40-1000kPa. The PSA-CH-1 unit and the PSA-CH-2 unit both use a three-tower adsorption process; the operation time sequence of each adsorption tower includes high-pressure adsorption, displacement, vacuum desorption and pressure increase processes, and cyclic operation, and the opening and closing time sequence of each exhaust valve, i.e. the cutting and reuse time sequence of the exhaust gas stream, is adjusted according to the composition and purity requirements of each separation product.

[0063] The reduction operation is required before the new catalyst is loaded and put into dehydrogenation operation, and the last step of the regeneration process is also the reduction operation; the reduction operation conditions are as follows: using a hydrogen stream with a hydrogen molar content higher than 98% and an olefin molar content lower than 0.1%, an airspeed of 500-1000h -1 , a pressure of 0.02-0.2MPa, a bed initial temperature of 480-500℃, a bed maximum temperature not higher than 520℃, and the water vapor molar content in the outlet gas stream is reduced to below 0.1%, and then extended for 2-4h.

[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. -1, the pressure 0-0.1 MPa, the bed initial temperature 450-500°C; during the carbon burning process, the maximum bed temperature is controlled to be not more than 520°C, until the difference between the O2 molar content in the outlet gas stream and the inlet gas stream is reduced to be below 0.1%, and then extended for 2-4 h. During the oxidation operation, an oxygen-containing gas stream with an O2 molar content of 10-20% obtained by mixing dry air stream into nitrogen stream or directly using dry air stream is used, the space velocity is 500-2000 h -1 , the pressure 0.02-0.2 MPa, the bed initial temperature 480-520°C, the gas stream inlet temperature 500-520°C, the oxidation treatment time 6-10 h. After the oxidation operation is completed, the reactor is changed to use nitrogen stream and high-purity nitrogen stream with a temperature of 480-500°C for purging, the space velocity is 500-2000 h -1 , until the O2 molar content in the outlet gas stream is reduced to be below 0.2%, and then the reduction of the catalyst is performed according to the above reduction operation conditions. The reactor or series loaded with the catalyst JB, after the in-situ regeneration treatment under the chlorine-free conditions is performed, when the dehydrogenation of isopentane is performed under the setting and operation conditions of the dehydrogenation reactor group mode A, B and C, the long-term operation effects of isopentane single-pass conversion rate of 30% or even above 35%, isopentene selectivity of 90% or above, methanation rate of 80% or above in the presence of CO2, and catalyst regeneration cycle of 300 h or above or even 450 h or above can be obtained; especially, the problem of isomerization side reaction to generate n-pentane is lighter than that of the catalyst JB after the in-situ regeneration with the oxygen-chlorination operation. After the in-situ regeneration under the chlorine-free conditions, the catalyst JB, the carbon deposition is basically burned out, and the isopentane dehydrogenation reaction performance and CO2 hydrogen-assisted methanation performance are greatly restored (slightly lower than the effect level in the first operation cycle).

[0066] Example 2

[0067] The present embodiment designs another process for preparing isoamylene from isoamylane dehydrogenation, which is basically the same as the process of embodiment 1, the difference is that the isoamylene separation device includes a rectification system in addition to the cooling separation system and the pressure swing adsorption separation system, wherein the PSA-CH device in the pressure swing adsorption unit is only provided with an alkylene separation unit PSA-CH-1 (without PSA-CH-2 unit); the rectification unit includes a rectification column D1 (theoretical tray 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-removing column D1, the isoamylane liquid stream Y-2 is separated at the top, and the isoamylene-rich liquid stream Z containing n-pentane and diene hydrocarbons is obtained at the bottom. The isoamylene-rich liquid stream Z is gasified and then passes through the PSA-CH-1, and is separated into three parts: n-pentane gas stream Z-11, isoamylene gas stream Z-21, and diene hydrocarbon-containing mixed hydrocarbon gas stream Z-31, wherein the isoamylene gas stream Z-21 and the isoamylene liquid stream Z-21L obtained by liquefaction thereof are the target products of isoamylane dehydrogenation in the present embodiment, and the molar purity can reach more than 98% or even more than 99%; the n-pentane gas stream Z-11 and the diene hydrocarbon-containing mixed hydrocarbon gas stream Z-31 can be further utilized or liquefied and sold. The isoamylene separation device can achieve the same overall effect as in embodiment 1.

[0068] In embodiments 1 and 2, the isoamylane dehydrogenation device can achieve a long-term running effect of isoamylane single-pass conversion rate of more than 30% or even more than 35%, isoamylene selectivity of more than 90%, and catalyst regeneration cycle of more than 300h; the isoamylene separation device can obtain isoamylene with molar purity of more than 98% or even more than 99%, and hydrogen; the service life of catalyst HA and dehydrogenation catalyst JB should be more than 3 years.

[0069] Embodiment 3

[0070] In the present embodiment, a Pt-Sn-K / alumina type alkane dehydrogenation catalyst HA (2.1L) is prepared, which is a spherical shape with an average outer diameter of Φ2.8mm; it is a spherical catalyst containing zirconium prepared and improved on the basis of the ratio and method of CN103212411A: taking high-temperature zirconium-aluminum composite oxide containing ZrO21.1wt% as the carrier, and loading Pt 0.35wt%, Sn 0.7wt%, and K 0.83wt% (calculated value in the catalyst) in terms of elements; it is prepared by the following steps:

[0071] (H-1) taking the pore volume 0.62mL / g, the specific surface area 118m 2 / g, spherical alumina with size Φ2.8 mm, dried at 400℃ for 3h, as carrier (room temperature, 1200g, tap density 2310ml, pore volume and specific surface area measured by low temperature N2 adsorption method), was placed in a sealable small rotating drum (volume 12L), sprayed evenly with 74.5ml of zirconium oxychloride aqueous solution with concentration 1.453mol / L, the spraying liquid amount was 90% of the saturated water absorption volume of the carrier (69.1ml / 100g), the spraying liquid time was 25min, the rotating drum was continuously rotated for 20min until the surface of the small balls was completely dry, the product was discharged, bagged, and placed for 15h for homogenization, then was placed in a tray and dried in an oven at 130℃ for 8h, then was calcined in a muffle furnace at 900℃ for 3h, when the temperature was decreased to 350℃, the product was discharged and placed in a 5L stainless steel tank which was previously replaced with nitrogen, sealed and connected to nitrogen to maintain positive pressure in the tank (to prevent moisture absorption), and then was cooled to room temperature, to obtain zirconium modified alumina carrier, the calculated ZrO2 content of which was 1.1wt%;

[0072] (H-2) 1100g of the zirconium modified alumina carrier obtained in step (H-1) was placed in a rotating drum, and 626ml of Pt-Sn-acid-K aqueous solution (temperature 26℃) was sprayed evenly thereon, the spraying liquid amount was 85% of the saturated water absorption volume of the zirconium modified alumina carrier (67.1L / 100g), the spraying liquid time was 25min, the rotating drum was continuously rotated for 28min until the surface of the small balls was completely dry, the product was discharged, bagged, and placed for 12h for homogenization, then was placed in a tray and dried in an oven at 130℃ for 8h, then was calcined in a muffle furnace at 550℃ for 3h, and the product was cooled and discharged, to obtain an alkane dehydrogenation catalyst HA, tap density 2110ml;

[0073] The Pt-Sn-acid-K aqueous solution in step (H-2) was prepared by mixing chloroplatinic acid, stannous chloride, hydrochloric acid, citric acid, and potassium nitrate into an aqueous solution with the desired concentration, the preparation method was as follows: the required amount of chloroplatinic acid and hydrochloric acid were mixed with an appropriate amount of water to form an aqueous solution, the required amount of stannous chloride, potassium nitrate, and citric acid were added and stirred until dissolved, water was added to the desired total liquid volume and stirred evenly, to obtain 630ml of Pt-Sn-acid-K aqueous solution; the concentration of chloroplatinic acid was 0.0327mol / L, the concentration of stannous chloride was 0.107mol / L, the concentration of potassium nitrate was 0.387mol / L, the concentration of HCl introduced by hydrochloric acid was 0.075mol / L, and the concentration of citric acid was 0.3mol / L (the concentrations of each element or substance were all calculated values), and the pH value was about 1.

[0074] The above preparation process was repeated again, to obtain 4.2L of alkane dehydrogenation catalyst HA. The obtained alkane dehydrogenation catalyst HA had uniform color on the outer surface and cross section, indicating that Pt was uniformly distributed in the catalyst particles.

[0075] Example 4

[0076] A Pt-Sn-K / alumina type alkane dehydrogenation catalyst JB (2.1 L) was prepared in this example, which was spherical with an average outer diameter of Φ 2.8 mm: a magnesium modified alumina containing 1.5 wt% of MgO was used as the carrier, and the loading of the components was 0.3 wt% of Pt, 0.7 wt% of Sn, and 0.83 wt% of K (calculated value in the catalyst) by elements, and was prepared by the following steps:

[0077] (J-1) A spherical alumina with a pore volume of 0.62 mL / g, a specific surface area of 118 m 2 / g, a size of Φ 2.8 mm, and which had been previously baked at 400°C for 3 h was used as the carrier (room temperature, 1200 g, tap density volume of 2310 mL), and was placed in a rotating drum (the same as that used in step H-1 of Example 2), and was uniformly sprayed with 497 mL of a magnesium bicarbonate aqueous solution with a concentration of 1.453 mol / L, the spraying liquid amount was 60% of the saturated water absorption volume of the carrier (69.1 mL / 100 g), the spraying liquid time was 25 min, and the rotating drum was continuously rotated for 6 min until the surface of the small balls was completely dry, and the material was discharged, bagged, and placed for 10 h of homogenization treatment in a sealed state, was loaded into a tray and was heated to 73°C in an oven and a flow of air, and was treated at a constant temperature for 4 h, and was then heated to 130°C and was dried for 8 h, and was then baked at 900°C for 3 h in a muffle furnace, and was discharged when the temperature was reduced to 350°C and was loaded into a stainless steel tank (the same as that used in step H-1 of Example 2, which had been previously replaced with nitrogen), was sealed and was maintained at a positive pressure in the tank by passing nitrogen gas (to prevent moisture absorption), was placed and cooled to room temperature, and a magnesium modified alumina carrier was obtained, the calculated value of the MgO content of which was 1.5 wt%;

[0078] (J-2) 1100 g of the magnesium modified alumina carrier obtained in step (J-1) was placed in a rotating drum, and was uniformly sprayed with 626 mL of a Pt-Sn-EDTA-K aqueous solution (temperature of 57°C), the spraying liquid amount was 85% of the saturated water absorption volume of the magnesium modified alumina carrier (67.0 mL / 100 g), the spraying liquid time was 25 min, and the rotating drum was continuously rotated for 22 min until the surface of the small balls was completely dry, and the material was discharged, bagged, and placed for 12 h of homogenization treatment in a sealed state, was loaded into a tray and was dried at 130°C in an oven for 8 h, and was then baked at 550°C for 3 h in a muffle furnace, and was cooled and discharged, and an alkane dehydrogenation catalyst JB was obtained, the tap density volume of which was 2120 mL;

[0079] The Pt-Sn-EDTA-K aqueous solution in step (J-2) is prepared by mixing diamminonitrito platinum (Pt(NH3)2(NO2)2), potassium stannate (K2SnO3), EDTA (ethylenediaminetetraacetic acid), nitric acid, potassium hydroxide, and potassium nitrite into an aqueous solution with a desired concentration. The preparation method is as follows: (A) 20.2 g of potassium stannate trihydrate is added to 300 mL of a potassium hydroxide aqueous solution with a pH value of 11, and stirred until dissolved. Then, 20.3 g of EDTA powder and 46.1 mL of a nitric acid aqueous solution (concentration: 3.22 mol / L) are added. The solution is heated to about 60°C and stirred for 30 min until clear. Then, 15.8 mL of a potassium hydroxide aqueous solution (concentration: 4.6 mol / L) is added to adjust the pH value to 6.3. Then, 3.1 g of potassium nitrite is added and stirred until dissolved to supplement the total amount of K. Finally, the solution is stirred to obtain a Sn-EDTA-K aqueous solution. The molar ratio of K2SnO3, EDTA, and HNO3 is 1:1.03:2.2. (B) The Sn-EDTA-K aqueous solution in step (A) is controlled at a temperature of 60°C. Then, 234.6 mL of a diamminonitrito platinum aqueous solution (Pt concentration: 0.0752 mol / L) at 60°C is added. Water is added to the total liquid amount of 630 mL and stirred to obtain a Pt-Sn-EDTA-K aqueous solution. The concentration of Pt is 0.028 mol / L, the concentration of Sn is 0.107 mol / L, the concentration of K is 0.387 mol / L, and the concentration of EDTA is 0.11 mol / L (as required, the concentrations of each element or substance are calculated values), and the pH value is about 6.5.

[0080] The above preparation process is performed again to obtain 4.2 L of the alkane dehydrogenation catalyst JB. The obtained alkane dehydrogenation catalyst JB has uniform color on the outer surface and cross section, indicating that Pt is uniformly distributed in the catalyst particles.

[0081] Catalyst evaluation example 1

[0082] The alkane dehydrogenation catalyst HA prepared in Example 2 is subjected to performance evaluation tests for preparing isoamylene by dehydrogenation of isoamylane under the conditions of a fixed bed and a quasi-adiabatic reactor in a specially designed dehydrogenation pilot plant. The dehydrogenation pilot plant is arranged in the production site of a certain isobutane dehydrogenation device, and the connection mode B has been used for a long time in the process development of the isobutane dehydrogenation process for preparing isobutene 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 structure, and heaters I, II and III, pipe mixers I and II, two high temperature ovens A and B with design temperature of 600℃ and with internal circulating fans, and two condensers A and B; the pipe mixer I, the heater I and the reactor I are installed in the high temperature oven A, the pipe mixer II, the heater II and the reactor II are installed in the high temperature oven B, and the heater III and the reactor III are also installed in the high temperature oven B when the connection mode C is adopted;

[0084] The reactors I, II and III all have a round plate shaped lower head, a cylindrical filling section (inner diameter 65mm, wall thickness 2mm), an upper flange and a thermowell (the upper flange and the lower head are respectively welded with a gas inlet interface and a gas outlet interface with inner diameter of 8mm and wall thickness of 1mm, the thermowell (outer diameter 6mm) is welded to the lower head with the central axis coinciding with the cylindrical filling section and extending upward to the top of the filling section, the gas flow direction in the reactor / catalyst bed is vertically downward), and are densely filled with alumina porcelain balls (outer diameter 3mm) 200mL, catalyst HA 1000mL (catalyst bed height about 300mm, height to diameter ratio 4.6), alumina porcelain balls (outer diameter 3mm) 200mL from bottom to top respectively, and four thermocouples are inserted into the thermowell for detecting the inlet gas temperature and the upper, middle and lower three section temperatures of the catalyst bed (three thermocouples for detecting the catalyst bed temperature, the ends of which are at the height of 50mm, 150mm and 250mm from bottom to top in the catalyst bed), and the unfilled space at the top of the reactor is about 220mL; the whole outer surface of the reactor is tightly covered with 25-30mm thick alumina fiber felt as heat insulation material, and the alumina fiber felt is fixed by wrapping with quartz gauze; the reactors I, II and III all basically have the adiabatic dehydrogenation reaction conditions;

[0085] The heaters I, II and III are all shell-and-tube heat exchangers without shell side, and the heat exchange areas are 1.3m 2 , 0.6m 2 , 0.6m 2 respectively; during the dehydrogenation operation, the mixed raw gas stream or the intermediate reaction gas stream is heated and warmed to the required reactor inlet temperature by adjusting the oven temperature; during the stable operation, the temperature difference between the ovens A and B and the gas inlet section temperature of the reactors I and II is within 2℃; the specifications of the pipe mixers I and II are the same, and the inner volumes are both 320mL, which are installed vertically with the stream entering from the bottom interface, and the pipe mixer I also has the function of gasifying isopentane;

[0086] The condensers A and B are the same specification tube-and-shell heat exchangers, and the heat exchange areas are both 1m 2, the shell side volume and the maximum liquid holdup are 12 L and 0.6 L respectively; the condenser A is a pre-cooling heat exchanger and uses circulating water as the cooling medium (circulating temperature 20-30℃), the condenser B uses the vehicle engine anti-freezing solution (glycol-water solution) with a freezing point of -40℃ as the cooling medium (circulating temperature can be adjusted in the range of -30℃ to -25℃); the condenser A and B are used in series to cool the reactor outlet gas stream to -22℃ to -20℃; an adjustable back pressure valve is arranged in the connecting pipeline between the condenser A and B to adjust the system pressure; the condenser B discharges the uncondensed gas stream and the condensed liquid stream, the uncondensed gas stream is continuously detected for the volume flow rate and temperature and sampled for chemical composition detection (once in 1-3h) and then sent to the flare network, the condensed liquid stream is sampled for chemical composition detection (once in 2-5h) and then sent to the storage tank (50 L, placed in the temperature-controlled ice cabinet at -30℃, weighed, emptied or replaced regularly) for temporary storage; the condenser B and the condensed liquid stream pipeline are respectively externally insulated by foamed plastic sheets;

[0087] The two side walls of the high temperature ovens A and B have low position openings for connecting pipelines and thermocouples (K type armored, outer diameter 1 mm) and are blocked by alumina cotton; the gas outlet interface of the reactor I and the gas outlet interface of the reactor II when the connecting mode C is used are respectively connected to high temperature gas sampling tubes which extend to the outside of the oven through the side wall openings of the oven and have needle valves at the ends which can withstand short-term high temperature to detect the composition of the reactor outlet gas stream;

[0088] During the evaluation test of the catalyst HA, the following three equipment configurations and connecting modes B, C and A were used in sequence according to the overall flow direction of the reaction stream:

[0089] Connecting mode B: the combined stream of isopentane liquid stream, hydrogen stream and carbon dioxide stream I → pipeline mixer I → heater I → reactor I → supplement of carbon dioxide stream II → pipeline mixer II → heater II → reactor II → condenser A → back pressure valve → condenser B; this connecting mode B is used for testing the performance of the catalyst HA in the dehydrogenation reactor group setting and operation mode B described in Example 1;

[0090] Connecting mode C: the combined stream of isopentane liquid stream, water stream and hydrogen stream → pipeline mixer I → heater I → reactor I → heater II → reactor II → heater III → reactor III → condenser A → back pressure valve → condenser B; this connecting mode C is used for testing the performance of the catalyst HA in the dehydrogenation reactor group setting and operation mode C described in Example 1;

[0091] Connection mode A: combined stream of isopentane stream, hydrogen stream and carbon dioxide stream I → pipe mixer I → heater I → reactor I → condenser A → back pressure valve → condenser B; the connection mode A is used for testing the performance of catalyst HA in the dehydrogenation reactor group setting and operation mode A described in Example 1;

[0092] The feed of reactor I during the dehydrogenation operation is configured as: hydrogen stream at a certain pressure (hydrogen pipe network from isobutane dehydrogenation production device, molar purity of more than 98%, hydrocarbon content of less than 350 mg / Nm 3 , unsaturated hydrocarbon content of less than 100 mg / Nm 3 , dew point of less than -18℃), carbon dioxide stream I (steel cylinder gasification, molar purity of more than 99.98%, hydrocarbon content of less than 3 mg / Nm 3 ) controlling the flow rate through mass flow meters, isopentane stream pressurized and flow rate controlled through plunger flow meter, and then combined to enter pipe mixer I installed in high temperature oven A, and the obtained reaction raw gas stream is warmed through heater I and then enters reactor I; the main composition of the isopentane used is: isopentane 99.2wt%, C4 and below components 0.31wt%, C6 and above components 0.16wt%, n-pentane 0.11wt%, total olefins 0.04wt%, and total sulfur 0.8 mg / Nm 3 ;

[0093] The materials of the reactors, heaters, coolers, pipe mixers, sampling tubes and connecting pipes, and the like are 310S heat-resistant and corrosion-resistant stainless steel.

[0094] According to the flow rates and chemical compositions of the uncondensed gas stream and the condensed liquid stream discharged from condenser B, the isopentane conversion rate, carbon dioxide conversion rate, isopentene selectivity and methane generation of the dehydrogenation reaction are calculated. The chemical compositions of the high-temperature gas stream at the outlet of reactor I (2-5h measured once, with a sampling temperature of about 100℃), the uncondensed gas stream of condenser B (sampling temperature of room temperature) and the condensed liquid stream (sampling temperature of -30℃) are detected by two gas chromatographs.

[0095] After the reactors I and II are filled with new catalyst, assembled according to connection mode B, and tested for tightness under 0.5 MPa nitrogen at room temperature, the catalyst is first reduced; the reduction operation conditions are: hydrogen stream (molar purity of more than 98%, hydrocarbon content of less than 350 mg / Nm 3 , unsaturated hydrocarbon content of less than 100 mg / Nm 3 , dew point of less than -18℃) from the hydrogen pipe network of the isobutane dehydrogenation production device at 750 NL / h (the gas hourly space velocities of the catalyst bed layers in the two reactors are both 750 h -1The back pressure valve in the connecting pipe between the condensers A and B was set at 0.05 MPa, the temperature of the high temperature ovens A and B was set at 5°C / min and 490°C respectively, and the hydrogen flow was introduced when the temperature started to rise. The condenser A was used (circulating water) while the condenser B was not used (normal temperature and no circulation of antifreeze). The upper section of the catalyst bed of the reactors I and II reached 480°C after 1.2 hours, and the lower section of the catalyst bed also reached 480°C. Then the reduction was continued for 3.5 hours. During the reduction, the temperature of the upper, middle and lower sections of the two catalyst beds did not exceed 516°C.

[0096] After the reduction was completed, the condenser B was used (cooling medium circulating temperature -30°C), the isopentane liquid flow and the carbon dioxide gas flow were introduced, the temperature of the high temperature ovens A and B was adjusted, and the hydrogen flow was adjusted. Then the dehydrogenation reaction evaluation of mode B was started.

[0097] The catalyst HA was subjected to the process test of mode B under the operation conditions listed in Table-1. In Table-1, the isopentane mass space velocity was 2h -1 The isopentane liquid flow was 2000 g / h, which was equivalent to the gasified isopentane flow of 620.9 NL / h. The CO2 molar content was the calculated value. The CO2 molar content in the inlet gas flow of the reactor I was adjusted by adjusting the flow of the carbon dioxide gas flow I, and the CO2 molar content in the inlet gas flow of the reactor II was adjusted by adjusting the flow of the carbon dioxide gas flow II based on the calculated flow and CO2 content of the outlet gas flow of the reactor I. The CO2 molar content in the two inlet gas flows was controlled to ensure that the upper, middle and lower sections of the catalyst beds of the reactors I and II did not exceed 550°C, and the isopentene yield was also considered.

[0098] Table-1 Dehydrogenation reaction evaluation operation conditions of the first running cycle of the catalyst HA

[0099]

[0100] Under the operation conditions listed in Table-1, the temperature difference between the inlet gas stream and the oven temperature of the two reactors is less than 3℃; the specific isopentane dehydrogenation reaction effect includes: during the evaluation test at the time, the maximum temperature of the catalyst bed of the two reactors can be controlled below 550℃; the temperature of the outlet section of the catalyst bed of reactor I can be controlled above 500℃, and the temperature of the outlet section of the catalyst bed of reactor II can be controlled above 530℃; the overall reaction effect of stably obtaining an isopentane total conversion rate of 30% or more (the operating time of an isopentane total conversion rate of 35% or more is 206h), an isopentene selectivity of 90% or more, and a total CO2 methanation rate of 80% or more can be stably obtained, wherein reactor I can stably obtain an isopentane conversion rate of 20% or more, an isopentene selectivity of 90% or more, and a CO2 methanation rate of 83% or more; the isopentane total conversion rate is mainly related to the overall temperature condition of the catalyst bed of reactor II and is not too related to the total amount of carbon dioxide gas stream, wherein the isopentane conversion rate of reactor I is closely related to the temperature of the lower section of the catalyst bed, i.e. the outlet section temperature. When the evaluation is up to 478h, the isopentane total conversion rate is reduced to below 30%, and the system pressure drop, i.e. the bed pressure drop of the two reactors, rises to 0.009MPa, indicating that a significant amount of carbon deposition has occurred.

[0101] At the end of the 480th hour, the isopentane and carbon dioxide feeds were turned off, the cooling medium circulation of condenser B was turned off (the water circulation of condenser A was not stopped), the back pressure valve was fully opened, and the first running cycle was ended. The doors and gas inlets and outlets of the high temperature ovens A and B were opened for about 20 minutes, the temperatures of the ovens A and B were reduced to about 490°C, the temperatures of the ovens A and B were set to 480°C, the hydrogen stream was switched to a 99.9% nitrogen stream with a flow rate of 2000 NL / h (the nitrogen stream was mixed through the hydrogen flow meter and the gas inlet pipeline), and the six temperatures of the catalyst beds of the two reactors were reduced to below 490°C after 1 hour of replacement. The first in-situ regeneration of the catalysts of the two reactors was started, including carbon burning, oxychlorination, and reduction operations. During the carbon burning operation, 80 NL / h of dry air (dew point -32°C) was first mixed through the flow meter of the carbon dioxide stream I and the gas inlet pipeline, and the molar content of O2 in the inlet gas stream of reactor I was about 0.76%. The O2 content in the outlet gas stream of reactor I was detected every 30 minutes. According to the detection value of the O2 content in the outlet gas stream of reactor I, the appropriate flow rate of dry air was mixed through the flow meter of the carbon dioxide stream II and the gas inlet pipeline to make the molar content of O2 in the inlet gas stream of reactor II about 0.7%. When the detection value of the O2 molar content in the outlet gas stream of reactor I rose to 0.7%, the molar content of O2 in the outlet gas stream of reactor II was detected (sampling from the exhaust pipe of condenser B). After the molar content of O2 in the outlet gas stream of reactor II reached 0.7%, the carbon burning was continued for another 3 hours, and the maximum temperature of the catalyst bed during the carbon burning was 515°C.After the end of the carbon burning operation, the oven A, B temperature was set to 520°C, the flow rate of the dry air through the carbon dioxide flow meter and the inlet gas line during the dehydrogenation operation was increased to 500 NL / h, the flow rate of the 99.9% nitrogen stream through the hydrogen flow meter and the inlet gas line was decreased to 500 NL / h (the O2molar content in the inlet gas stream of the reactor I was 10%), the pressure control of the back pressure valve was 0.05 MPa, the flow rate of the chloroacetic acid aqueous solution (water, chlorine molar ratio 6.5) with a mass concentration of 44.5% and a flow rate of 6.6 g / h was added through the isopentane flow pump during the dehydrogenation operation, the chloroacetic acid aqueous solution was stopped after 5 h (the amount of chloroacetic acid added was 0.5 wt% of the total amount of catalyst in the two reactors based on chlorine, and the temperature of the lower part of the catalyst bed in the two reactors reached more than 515°C after about 1 h); then the oven A, B temperature was set to 530°C, the oxygen chlorination operation was completed after 5 h of treatment; then the dry air was stopped, the oven A, B door and the inlet and outlet ports were opened for about 20 min, the oven A, B temperature was reduced to about 490°C, the oven A, B temperature was set to 480°C, the flow rate of the 99.9% nitrogen stream was increased to 2000 NL / h, and the O2molar content in the outlet gas stream of the reactor II was reduced to less than 0.1% after 1 h. The 99.9% nitrogen stream was switched to the pipe network hydrogen stream 750 NL / h (the gas hourly space velocity of the catalyst bed in the two reactors was 750 h. -1 ), and the catalyst reduction was completed after 5 h.

[0102] The oven A, B temperature was appropriately increased, and the isopentane dehydrogenation reaction evaluation test of the second running period of the catalyst HA was carried out according to the operation conditions described in Table-1. The results were basically repeated in the first running period, the maximum temperature of the catalyst bed in the two reactors could be controlled below 550°C, the temperature of the outlet section of the catalyst bed in the reactor I could be controlled above 500°C, and the temperature of the outlet section of the catalyst bed in the reactor II could be controlled above 530°C. The overall reaction effect of more than 30% total isopentane conversion (the running time of more than 35% total isopentane conversion was 202 h), more than 90% isopentene selectivity, and more than 80% total CO2methanation rate could be stably obtained, and the reactor I could stably obtain more than 20% isopentane conversion, more than 90% isopentene selectivity, and more than 85% CO2methanation rate. When the evaluation reached the 491st hour, the total isopentane conversion decreased to less than 30%, and the system pressure drop, i.e. the bed pressure drop of the two reactors, increased to 0.011 MPa, indicating that a significant amount of carbon deposition had occurred.

[0103] The catalysts of the two reactors were regenerated in situ for the second time according to the first in-situ regeneration method, and then the isopentane dehydrogenation reaction evaluation test of the third running cycle of catalyst HA was carried out according to the operation conditions and control method of the third running cycle, and then the catalysts of the two reactors were regenerated in situ for the fourth time according to the first in-situ regeneration method, and then the isopentane dehydrogenation reaction evaluation test of the fifth running cycle of catalyst HA was carried out according to the operation conditions and control method of the fourth running cycle, until the total isopentane conversion rate decreased to less than 30%. The results are basically repeated in the third running cycle. The running time of the total isopentane conversion rate of more than 30% is 443h and 457h respectively, and the running time of the total isopentane conversion rate of more than 35% is 275h and 284h respectively. -1 The results include: the temperature of the outlet section of the catalyst bed of reactor I and reactor II can be maintained at more than 510°C and more than 530°C respectively; the overall reaction effect of more than 30% total isopentane conversion rate (the running time of more than 35% total isopentane conversion rate is 281h), more than 90% isopentene selectivity, and more than 80% total CO2 methanation rate can be stably obtained, among which reactor I can stably obtain more than 20% isopentane conversion rate, more than 90% isopentene selectivity, and more than 82% CO2 methanation rate. When the evaluation is carried out to the 455th hour, the total isopentane conversion rate decreases to less than 30%, and the system pressure drop, i.e. the bed pressure drop of the two reactors, rises to 0.016MPa (which indicates that a significant amount of carbon deposition has occurred).

[0104] The catalysts of the two reactors were regenerated in situ for the second time according to the first in-situ regeneration method, and then the isopentane dehydrogenation reaction evaluation test of the third running cycle of catalyst HA was carried out according to the operation conditions and control method of the third running cycle, and then the catalysts of the two reactors were regenerated in situ for the fourth time according to the first in-situ regeneration method, and then the isopentane dehydrogenation reaction evaluation test of the fifth running cycle of catalyst HA was carried out according to the operation conditions and control method of the fourth running cycle, until the total isopentane conversion rate decreased to less than 30%. The results are basically repeated in the third running cycle. The running time of the total isopentane conversion rate of more than 30% is 443h and 457h respectively, and the running time of the total isopentane conversion rate of more than 35% is 275h and 284h respectively.

[0105] In the isopentane dehydrogenation reaction (mode B) evaluation test of the third, fourth and fifth running cycles, the CO2 molar content in the inlet gas stream of reactor I was gradually increased from 1.0wt% to 2.0wt%, and the CO2 molar content in the inlet gas stream of reactor II was gradually increased from 1.3wt% to 2.0wt%.

[0106] The fifth in-situ regeneration was performed on the catalysts in the two reactors according to the first in-situ regeneration method above, and before reduction, the heater III and the reactor III (with other connection relationships appropriately changed, the reactor III was filled with the new catalyst HA to be reduced) were connected, the dehydrogenation reaction of isopentane in mode C and the sixth running cycle was evaluated, the carbon dioxide gas streams I and II were closed, and the water stream was injected through the isopentane pipeline (pressurized and controlled by the plunger flow pump); the specific conditions included: the inlet gas stream temperature of the reactor I was 560°C, the pressure was 0.03 MPa, the isopentane mass space velocity was 2h -1 -1, the molar ratio of isopentane to H2 and water was 1:1:2 (without CO2); the inlet gas stream temperature of the reactors II and III was 550°C. The results included: the outlet section temperatures of the catalyst bed layers of the reactors I, II and III could be maintained at above 498°C, above 510°C and above 530°C respectively; the overall reaction effect of the stable isopentane total conversion rate of more than 30% (the running time of the isopentane total conversion rate of more than 35% was 295h) and the isopentene selectivity of more than 90% could be obtained, and the reactor I could stably obtain the reaction effect of the isopentane conversion rate of more than 15% and the isopentene selectivity of more than 90%. When the evaluation was performed to the 572th hour, the isopentane total conversion rate decreased to less than 30%, and the system pressure drop, i.e. the bed layer pressure drop of the two reactors, increased to 0.019 MPa (indicating that a significant amount of carbon deposition had occurred).

[0107] The sixth in-situ regeneration was performed on the reactor system according to the first in-situ regeneration method above (the main difference was that the amount of chloroacetic acid was 0.5wt% of the total amount of catalysts in the three reactors in terms of chlorine, and the chloroacetic acid aqueous solution (water, chlorine molar ratio 6.5) with a mass concentration of 44.5% and a flow rate of 6.6g / h was injected for 7.5h, wherein the catalysts in the reactors I and II were the sixth in-situ regeneration, and the catalyst in the reactor III was the first in-situ regeneration), and then the dehydrogenation reaction of isopentane in mode C and the seventh running cycle was evaluated; the dehydrogenation reaction conditions of isopentane included: the inlet gas stream temperature of the reactor I was 560°C, the pressure was 0.05 MPa, the isopentane mass space velocity was 3h -1, the molar ratio of isopentane to H2, water 1:0.5:2.5 (without CO2); the temperature of the inlet gas stream of reactors II and III 550°C. The results include: the temperature of the outlet section of the catalyst bed of reactors I, II and III can be maintained at above 498°C, above 510°C and above 530°C, respectively; the overall reaction effect of isopentane total conversion rate of above 30% (the running time of isopentane total conversion rate of above 35% is 386h) and isopentene selectivity of above 90% can be stably obtained, wherein reactor I can stably obtain the reaction effect of isopentane conversion rate of above 15% and isopentene selectivity of above 90%. At the 605th hour of the evaluation, the isopentane total conversion rate decreased to below 30%, and the system pressure drop, i.e. the bed pressure drop of the two reactors, increased to 0.015 MPa (indicating that a significant amount of carbon deposition has occurred).

[0108] The reactor system was regenerated in-situ for the seventh time according to the sixth in-situ regeneration method described above, and then the isopentane dehydrogenation reaction evaluation test of the eighth running cycle of catalyst HA was carried out according to the operating conditions and control method of the seventh running cycle (mode C), until the isopentane total conversion rate decreased to below 30%. The results basically repeated the effect of the seventh running cycle, the running time of isopentane total conversion rate of above 30% was 600h, and the running time of isopentane total conversion rate of above 35% was 373h.

[0109] The reactor system was regenerated in-situ for the eighth time according to the sixth in-situ regeneration method described above, and then the isopentane dehydrogenation reaction evaluation of the ninth running cycle of reactor I was carried out according to mode A, with the gas outlet of reactor I of oven A connected to condenser A (the high-temperature gas sampling tube was still used, the carbon dioxide gas stream II inlet system, oven B and the built-in pipeline mixer II, heater II / III, reactor II / III were not used), and the water injection pump and pipeline were disconnected; the specific conditions include: the temperature of the inlet gas stream of reactor I 540°C, the pressure 0.05 MPa, the isopentane mass space velocity 2h -1, the molar ratio of isopentane to H2 was 1:1, the overall temperature of the catalyst bed of reactor I was controlled to the highest level possible but the highest temperature was not more than 550℃, and the temperature of the outlet section was maintained at more than 530℃ as much as possible by adjusting the flow rate of the carbon dioxide gas stream I. The results included that the temperature of the outlet section of the catalyst bed of reactor I was maintained at more than 530℃ for a long time, the evaluation time of isopentane conversion of more than 30% was 452h (the operation time of the total isopentane conversion of more than 35% was 233h), and the reaction effects of isopentene selectivity of more than 90% and CO2 methanation rate of more than 85% were obtained at the same time. The main control method was to gradually increase the CO2 molar content of the inlet gas stream of reactor I from 1.7% to 2.2%. After the evaluation for 454h, the reaction effects of isopentane conversion of more than 30% and isopentene selectivity of more than 90% could not be obtained at the same time, and the bed pressure drop of reactor I increased to 0.08MPa (indicating that a significant amount of carbon deposition had occurred).

[0110] In the above nine operation cycles, the main dehydrogenation products of isopentane and the side reaction conditions included that in the main dehydrogenation product isopentene, the molar proportion of 2-methyl-2-butene was 70-82%, and the molar proportion of 3-methyl-1-butene was 1-3%; the main side reaction was the isomerization reaction to generate n-pentane, the generation proportion of n-pentane in the isopentane side reaction products was about 80% (calculated by carbon), and the molar ratio of n-pentane to its dehydrogenation product (various straight-chain pentenes) was higher than 20:1; the secondary side reaction was the cracking reaction to generate carbon four or lower low-carbon hydrocarbons, the proportion in the isopentane side reaction products was less than 15% (calculated by carbon); the proportion of carbon five diene hydrocarbons in the isopentane side reaction products was less than 6% (calculated by carbon) and mainly 3-methyl-1,2-butadiene, and 2-methyl-1,3-butadiene (isoprene) was less; the proportion of carbon nine or higher hydrocarbons in the isopentane side reaction products was less than 2.7% (calculated by carbon).

[0111] From the overall reaction effect of this evaluation example 1, it can be seen that the catalyst HA has moderate CO2 hydrogenation methanation performance and high isopentane dehydrogenation reaction performance (isopentane dehydrogenation to isopentene) under the operation conditions, the heat release of the methanation reaction of CO2 significantly improves the conversion rate of isopentane in the dehydrogenation reaction, and does not substantially increase the side reaction; in the above evaluation process, the isopentane conversion rate of more than 30% is stably obtained under the above-mentioned lower reaction temperature conditions, the isopentene selectivity is stably maintained at more than 90%, and the regeneration period is more than 350h; the in-situ regeneration does not substantially change or reduce the reaction performance of the catalyst HA; the conversion rate of isopentane in the dehydrogenation reaction is limited to a certain extent by the thermodynamic reaction equilibrium. The above-mentioned three equipment configurations and connection modes A, B and C correspond to the configurations and operation modes A, B and C in the embodiment 1, respectively, wherein the reactor conditions in the modes B and C are relatively easy to achieve and have certain operation flexibility, and the operation flexibility of the reactor conditions in the mode A is relatively narrow but can be achieved and is relatively simplified.

[0112] Catalyst evaluation example 2

[0113] In another set of dehydrogenation pilot plant devices which are completely the same as those used in the catalyst evaluation example 1, the Pt-Sn-K / alumina type alkane dehydrogenation catalyst JB prepared in the embodiment 3 is subjected to the process test of preparing isopentene from isopentane dehydrogenation under the conditions of fixed bed and quasi-adiabatic reactor in this evaluation example 2.

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

[0115] After the system (connection mode B, reactors I and II are used) of the first running period is assembled, qualified under the condition of 0.5MPa nitrogen gas at room temperature, the catalyst is reduced, and the reduction operation conditions are completely the same as those in the catalyst evaluation example 1; during the reduction process, the temperatures of the upper, middle and lower sections of the two catalyst beds do not exceed 519℃.

[0116] Immediately after reduction of catalyst JB, condenser B (cold medium circulation temperature -30°C) was started, the isopentane liquid stream and the carbon dioxide gas stream were switched on, the temperature conditions of high temperature ovens A, B and the hydrogen flow rate were adjusted, the dehydrogenation reaction evaluation was started, and the dehydrogenation reaction evaluation of isopentane was carried out for ten operation cycles in the same way as in catalyst evaluation example 1, but the in-situ regeneration method was used under the conditions of no chlorine and with chlorine. First, catalyst JB was subjected to process test under the operating conditions listed in Table-2, mode B. The CO2molar content in the inlet gas stream of reactors I, II was controlled, and the upper, middle and lower three sections of the catalyst bed of reactors I, II were also controlled so that the temperature did not exceed 550°C, and the isopentene yield was also considered.

[0117] Table-2 Dehydrogenation reaction evaluation operating conditions of catalyst JB in the first operation cycle

[0118]

[0119] Under the operating conditions listed in Table-2, the difference between the temperature of the inlet gas stream of the two reactors and the temperature of the oven was less than 3°C; the specific isopentane dehydrogenation reaction effect included: during the evaluation test at the time, the maximum temperature of the catalyst bed of the two reactors could be controlled below 550°C; the temperature of the outlet section of the catalyst bed of reactor I could be controlled above 500°C, and the temperature of the outlet section of the catalyst bed of reactor II could be controlled above 530°C; the overall reaction effect of stable isopentane total conversion rate of more than 30% (isopentane total conversion rate of more than 35% for 389 hours), isopentene selectivity of more than 90%, and total CO2methanation rate of more than 80% was obtained, among which reactor I could stably obtain isopentane conversion rate of more than 20%, isopentene selectivity of more than 90%, and CO2methanation rate of more than 85%; the isopentane total conversion rate was mainly related to the overall temperature of the catalyst bed of reactor II and was not much related to the total amount of carbon dioxide gas stream, among which the isopentane conversion rate of reactor I was closely related to the temperature of the lower section of the catalyst bed of the reactor, i.e. the outlet section. When the evaluation reached 633h, the isopentane total conversion rate decreased to less than 30%, and the system pressure drop, i.e. the bed pressure drop of the two reactors, increased to 0.012MPa (indicating that a significant amount of carbon deposition had occurred).

[0120] At the end of the 635th evaluation, the isopentane and carbon dioxide feeds were turned off, the cooling medium circulation of condenser B was turned off (the water circulation of condenser A was not stopped), the back pressure valve was fully opened, the first running cycle was ended, the doors and inlet and outlet ports of the high temperature ovens A and B were opened for about 20 min, the temperatures of the ovens A and B were reduced to about 490°C, the temperatures of the ovens A and B were set to 480°C, the hydrogen stream was switched to a 99.9% nitrogen stream with a flow rate of 2000 NL / h (99.9% nitrogen was added through the hydrogen flow meter and the inlet pipe), and the six temperatures of the catalyst beds of the two reactors were reduced to below 490°C after 1 h of replacement. The first in-situ regeneration of the catalysts of the two reactors was started, including carbon burning, oxidation, and reduction operations. During the carbon burning operation, 80 NL / h of dry air (dew point -32°C) was first added through the flow meter of carbon dioxide stream I and the inlet pipe (about 0.7% molar content of O2 was added to the inlet stream of reactor I, the O2 content in the outlet stream of reactor I was detected once every 30 min; according to the detection value of the O2 content in the outlet stream of reactor I, appropriate flow rate of dry air was added through the flow meter of carbon dioxide stream II and the inlet pipe to make the molar content of O2 in the inlet stream of reactor II about 0.7%; when the detection value of the molar content of O2 in the outlet stream of reactor I rose to 0.65%, the molar content of O2 in the outlet stream of reactor II was detected (sampling from the exhaust pipe of condenser B), and after the molar content of O2 in the outlet stream of reactor II reached 0.65%, the carbon burning was continued for another 3 h, and the maximum temperature of the catalyst bed during the carbon burning was 513°C. After the completion of the carbon burning operation, the temperatures of the ovens A and B were set to 510°C, the flow rate of dry air added through the flow meter of carbon dioxide stream I and the inlet pipe was increased to 500 NL / h, the flow rate of the 99.9% nitrogen stream added through the hydrogen flow meter and the inlet pipe was reduced to 500 NL / h (the molar content of O2 in the inlet stream of reactor I was 10%), the pressure of the back pressure valve was controlled to 0.05 MPa, and the oxidation operation was completed after 7 h of treatment; then the dry air was stopped, the doors and inlet and outlet ports of the high temperature ovens A and B were opened for about 20 min, the temperatures of the ovens A and B were reduced to about 490°C, the temperatures of the ovens A and B were set to 480°C, the flow rate of the 99.9% nitrogen stream was increased to 2000 NL / h, and the molar content of O2 in the outlet stream of reactor II was reduced to below 0.1% after 1 h; the 99.9% nitrogen stream was switched to the pipe network hydrogen stream 750 NL / h (the gas hourly space velocities of the catalyst beds in the two reactors were both 750 h -1 ), and the reduction of the catalyst was completed after 5 h.

[0121] The set temperature of the ovens A and B was then increased again, and the second run of the catalyst JB was then subjected to the isopentane dehydrogenation reaction evaluation test in mode B. The specific conditions included: the inlet gas stream temperature of reactor I was 540°C, the pressure was 0.03 MPa, the isopentane mass space velocity was 2 h -1 , the molar ratio of isopentane to H2 was 1:0.5; the inlet gas stream temperature of reactor II was 540°C; the overall temperature of the catalyst beds of reactors I and II was controlled to an appropriate high level by adjusting the flow rates of the carbon dioxide gas streams I and II, but the highest temperature was always no more than 550°C. The results included: the outlet section temperatures of the catalyst beds of reactors I and II could be maintained at above 510°C and above 530°C, respectively; the overall reaction effect of stably obtaining an isopentane total conversion rate of above 30% (the running time of an isopentane total conversion rate of above 35% was 377 h), an isopentene selectivity of above 90%, and a total CO2 methanation rate of above 80% was achieved, in which reactor I could stably obtain an isopentane conversion rate of above 20%, an isopentene selectivity of above 90%, and a CO2 methanation rate of above 82%. When the evaluation was performed to the 506th hour, the isopentane total conversion rate decreased to below 30%, and the system pressure drop, i.e., the bed pressure drop of the two reactors, increased to 0.015 MPa (indicating that a significant amount of carbon deposition had occurred).

[0122] The second in-situ regeneration of the catalysts of the two reactors was performed according to the first in-situ regeneration method (chlorine-free regeneration) of this evaluation example 2, and then the third run of the catalyst JB was subjected to the isopentane dehydrogenation reaction evaluation test in mode B according to the operating conditions and control method of the second run, until the isopentane total conversion rate decreased to below 30%. Then the third in-situ regeneration of the catalysts of the two reactors was performed according to the first in-situ regeneration method (chlorine-free regeneration) of this evaluation example 2, and the fourth run of the catalyst JB was subjected to the isopentane dehydrogenation reaction evaluation test according to the operating conditions and control method of the second run, until the isopentane total conversion rate decreased to below 30%. The results were basically repeated the effect of the second run, and the running times of the isopentane total conversion rate of above 30% were 491 h and 505 h, respectively, and the running times of the isopentane total conversion rate of above 35% were 364 h and 372 h, respectively.

[0123] During the isopentane dehydrogenation reaction (mode B) evaluation test of the above-mentioned second, third, and fourth runs, the molar content of CO2 in the inlet gas stream of reactor I was gradually increased from the initial 1.0 wt% to 2.0 wt%, and the molar content of CO2 in the inlet gas stream of reactor II was gradually increased from the initial 1.3 wt% to 2.0 wt%.

[0124] The catalysts in the two reactors were regenerated in situ for the fourth time according to the first in-situ regeneration method (chlorine-free regeneration) of this evaluation example 2, and before reduction, heater III and reactor III (with other connection relationships appropriately changed, reactor III was filled with new catalyst JB and was reduced) were connected, and then the dehydrogenation reaction evaluation of isopentane was carried out in mode C and the fifth running cycle, the carbon dioxide gas streams I and II were closed, and the water stream was injected through the isopentane pipeline (pressurized and controlled by a plunger flow pump); the specific conditions included: the inlet gas stream temperature of reactor I was 560°C, the pressure was 0.03 MPa, the isopentane mass space velocity was 2.5 h -1 , the molar ratio of isopentane to H2 and water was 1:0.5:3.0 (without CO2); the inlet gas stream temperature of reactors II and III was 550°C. The results included: the outlet section temperatures of the catalyst bed layers of reactors I, II and III could be maintained at above 495°C, above 510°C and above 530°C, respectively; the overall reaction effect of stable isopentane total conversion of above 30% (the running time of isopentane total conversion of above 35% was 419 h) and isopentene selectivity of above 90% was obtained, and reactor I could stably obtain the reaction effect of isopentane conversion of above 15% and isopentene selectivity of above 90%. When the evaluation was carried out to the 641st hour, the isopentane total conversion decreased to below 30%, and the system pressure drop, i.e. the bed layer pressure drop of the two reactors, increased to 0.017 MPa (indicating that a significant amount of carbon deposition had occurred).

[0125] The catalysts in the two reactors were regenerated in situ for the fifth time according to the first in-situ regeneration method (chlorine-free regeneration) of this evaluation example 2, and then the dehydrogenation reaction evaluation of isopentane was carried out in mode C and the sixth running cycle of catalyst JB according to the operating conditions and control methods of the fifth running cycle, until the isopentane total conversion decreased to below 30%, and then the reactor system was regenerated in situ according to the first in-situ regeneration method (chlorine-free regeneration, in which the catalysts in reactors I and II were regenerated in situ for the sixth time, and the catalyst in reactor III was regenerated in situ for the first time), and the dehydrogenation reaction evaluation of isopentane was carried out in the seventh running cycle of catalyst JB according to the operating conditions and control methods of the fifth running cycle, until the isopentane total conversion decreased to below 30%. The results were basically repeated the effect of the fifth running cycle, and the running time of isopentane total conversion of above 30% was 653 h and 640 h, respectively, and the running time of isopentane total conversion of above 30% was 415 h and 433 h, respectively.

[0126] The seventh in-situ regeneration (chlorine-free regeneration) of the catalyst in the reactor system was carried out according to the first in-situ regeneration method of this evaluation example 2, after which the outlet of the oven A reactor I was connected to the condenser A (the high-temperature gas flow sampling tube was continuously used, the carbon dioxide gas flow II inlet system, the oven B, and the pipeline mixer II, heater II, and reactor II were not used), the water injection pump and the pipeline were disconnected, and the dehydrogenation reaction evaluation of isopentane in mode A and the eighth running cycle of reactor I was carried out. The specific conditions included: the inlet gas flow temperature of reactor I was 540°C, the pressure was 0.05 MPa, the isopentane mass space velocity was 1.8 h -1 -1, the molar ratio of isopentane to H2 was 1:1, the overall temperature of the catalyst bed of reactor I was controlled to the highest possible level but the highest temperature did not exceed 550°C by adjusting the flow rate of carbon dioxide gas flow I, and the outlet section temperature was maintained above 530°C as much as possible. The results included: the outlet section temperature of the catalyst bed of reactor I could be maintained above 530°C for a long time, the evaluation time of isopentane conversion rate above 30% was 439h (the running time of isopentane total conversion rate above 35% was 192h), and the reaction effect of isopentene selectivity above 90% and CO2 methanation rate above 85% was obtained at the same time. The control method was mainly to gradually increase the CO2 molar content of the inlet gas flow of reactor I from 1.8% to 2.5%.

[0127] After the carbon dioxide gas flow II inlet system, the pipeline mixer II, the heater II, and the reactor II were reconnected, the catalysts in reactors I and II were regenerated in-situ (chlorine-free regeneration, in which reactor I was regenerated for the eighth time and reactor II was regenerated for the eighth time) according to the first in-situ regeneration method of this evaluation example 2. Then, the isopentane dehydrogenation reaction evaluation test of catalyst JB was continued according to the operating conditions and control method of the second running cycle, mode B, until the total isopentane conversion rate decreased to below 30%. The results were basically repeated the effect of the second running cycle, the running time of isopentane total conversion rate above 30% was 486h, and the running time of isopentane total conversion rate above 35% was 369h.

[0128] The catalysts in reactors I and II were regenerated for the ninth time in situ (chlorine regeneration) according to the first in situ regeneration method described in Evaluation Example 1, and then the dehydrogenation reaction of isopentane over catalyst JB was evaluated according to the operating conditions and control method of the third running cycle, mode B, of Evaluation Example 1, until the total conversion of isopentane dropped to below 30%. Thereafter, the catalysts in the two reactors were regenerated for the tenth time in situ (chlorine regeneration) according to the first in situ regeneration method described in Evaluation Example 1, and then the dehydrogenation reaction of isopentane over catalyst JB was evaluated according to the operating conditions and control method of the third running cycle, mode B, of Evaluation Example 1, until the total conversion of isopentane dropped to below 30%. The operating conditions included: the temperature of the inlet gas stream of reactor I was 540°C, the pressure was 0.05 MPa, the mass space velocity of isopentane was 2.5 h -1 -1, and the molar ratio of isopentane to H2 was 1:0.5; the temperature of the inlet gas stream of reactor II was 530°C; the overall temperature of the catalyst bed in reactors I and II was controlled to be at a proper high level but the highest temperature was always below 550°C by adjusting the flow rates of carbon dioxide gas streams I and II. The results included: the temperature of the outlet section of the catalyst bed in reactors I and II could be maintained at above 510°C and above 530°C, respectively; an overall reaction effect of total conversion of isopentane of above 30% (the running time of total conversion of isopentane of above 35% was 281 h), isopentene selectivity of above 90%, and total CO2 methanation rate of above 80% was stably obtained, wherein reactor I stably obtained a reaction effect of isopentane conversion of above 18%, isopentene selectivity of above 90%, and CO2 methanation rate of above 85%; the control method mainly included gradually increasing the molar content of CO2 in the inlet gas stream of reactor I from 1.1% to 1.8% and gradually increasing the molar content of CO2 in the inlet gas stream of reactor II from 1.3% to 2.0%; the running time of total conversion of isopentane of above 30% was 756 h and 742 h, respectively, and the running time of total conversion of isopentane of above 35% was 501 h and 515 h, respectively.

[0129] In the above eleven operation cycles of the present evaluation example 2, the main dehydrogenation product of isopentane and the side reaction conditions include: in the main dehydrogenation product isopentene, the molar proportion of 2-methyl-2-butene is 78-86%, and the molar proportion of 3-methyl-1-butene is 1-2.5%; the main side reaction is the isomerization reaction to generate n-pentane, and the generation proportion of n-pentane in the isopentane side reaction product is about 70% (calculated by carbon), and the molar ratio of n-pentane to its dehydrogenation product (various straight-chain pentenes) is higher than 20:1; the secondary side reaction is the cracking reaction to generate carbon four or lower low-carbon hydrocarbons, and the proportion in the isopentane side reaction product is less than 20% (calculated by carbon); the proportion of carbon five diene hydrocarbons in the isopentane side reaction product is less than 5% (calculated by carbon) and mainly 3-methyl-1,2-butadiene, and 2-methyl-1,3-butadiene (isoprene) is often not detected; the proportion of carbon nine or higher hydrocarbons in the isopentane side reaction product is less than 2.1% (calculated by carbon).

[0130] From the overall reaction effect of the present evaluation example 2, it can be seen that the catalyst JB has moderate CO2 hydrogenation methanation performance and high isopentane dehydrogenation reaction performance (isopentene generated by isopentane dehydrogenation) under the operation conditions, and the methanation reaction heat of CO2 significantly improves the conversion rate of isopentane in the dehydrogenation reaction, and basically does not increase the side reaction; during the evaluation process, the isopentane conversion rate of more than 30% is stably obtained under the lower reaction temperature conditions, the isopentene selectivity is stably maintained at more than 90%, and the regeneration cycle is more than 450h or even more than 700h; controlling the catalyst bed temperature to the lower level is the key condition to obtain a longer regeneration cycle; the chlorine-free and chlorine-containing in-situ regeneration methods have good regeneration effect on the catalyst JB; the conversion rate of isopentane in the dehydrogenation reaction is also obviously limited by the thermodynamic reaction equilibrium. The three equipment configurations and connection modes A, B and C in the present embodiment 2 correspond to the configurations and operation modes A, B and C in embodiment 1 respectively, wherein the reactor conditions in modes B and C are relatively easy to achieve and have certain operation flexibility, and the operation flexibility of the reactor conditions in mode A is relatively narrow but can be achieved and is relatively simplified.

[0131] Embodiment 5

[0132] The present embodiment 5 designs a preparation process of tertiary amyl alcohol, using the isopentene liquid stream Z-2L or Z-21L (molar purity of 99% or more) obtained by the process of the present embodiment 1 or 2 and water as raw materials, mixing with the recycled residual isopentene liquid stream X-1, the recycled aqueous solution stream X-31 containing tertiary amyl alcohol, and the recycled solvent liquid stream X-4 to obtain a liquid stream X, and performing liquid phase reaction of isopentene hydration to generate tertiary amyl alcohol in a fixed bed hydration reactor filled with macroporous sulfonic acid type cation resin catalyst (hydrogen type, product JQE-02 of the present company); the solvent contained in the recycled solvent liquid stream X-4 is ethylene glycol monobutyl ether; the liquid stream obtained by the hydration reaction is cut to obtain the residual isopentene liquid stream X-1, the miscible liquid stream X-20 of tertiary amyl alcohol and water, and the recycled solvent liquid stream X-4 by passing through a light component removal column D3 (theoretical plate number of about 20) and a rectification column D4 (theoretical plate number of 40-50) in sequence; the miscible liquid stream X-20 of tertiary amyl alcohol and water is separated to obtain the tertiary amyl alcohol liquid stream X-2 and the aqueous solution stream X-3 containing tertiary amyl alcohol by passing through a rectification column D5 (theoretical plate number of 40-50) with a top phase separator and azeotropic rectification method; the residual isopentene liquid stream X-1, the aqueous solution stream X-3 containing tertiary amyl alcohol, and the recycled 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, isopentene, and ethylene glycol monobutyl ether in the raw material liquid stream X is (1-2):1:(4-6), the liquid hourly space velocity is 0.6-1 h-1, the feed temperature is 40-65°C, the pressure is 0.5-1.0 MPa, and the discharge temperature is 65-80°C; the hydration reactor is also provided with an external circulation system, and 40-50% of the volume flow rate of the discharge liquid stream is recycled as the recycled material liquid stream X-5 and mixed into the liquid stream X to enter the hydration reactor again; a coil pipe for circulating water at 60-65°C is arranged in the resin catalyst bed as a heat removal temperature control component; and part of the flow rate of the residual isopentene liquid stream X-1 is sent back to the cooling separation system of the isopentene separation device to be combined and treated with the hydrocarbon-containing liquid stream L0 (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 residual isopentene liquid stream X-1 at more than 90%; the recycled solvent liquid stream X-4 is periodically subjected to fractionation purification treatment to separate and remove components other than ethylene glycol monobutyl ether, which are not conducive to improving the purity of the tertiary amyl alcohol liquid stream X-2, especially isopentene polymers and fusel alcohols. -1

[0133] ​The tertiary amyl alcohol and water mixed solution stream X-20, and the tertiary amyl alcohol-water azeotrope separated from the top of the rectification tower D4, are separated in a temperature-controlled layering device at the top of the rectification tower D4, the upper layer is a water-containing tertiary amyl alcohol solution stream X-21, and the lower layer is a tertiary amyl alcohol-containing water solution stream X-31, the water-containing tertiary amyl alcohol solution stream X-21 is used as the middle section feed of the rectification tower D4, and finally a tertiary amyl alcohol stream X-2 with a purity of more than 99 wt% or even more than 99.5 wt% is separated, the water content is less than 0.1 wt%, and after cooling, the tertiary amyl alcohol product can be packed, stored and transported.

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

Claims

1. A process for preparing isopentene by dehydrogenation of isopentane, comprising an isopentane dehydrogenation unit and an isopentene separation unit; 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; The setup and operation of dehydrogenation reactor units include any of the following modes: Mode A: Set up 2-4 reactors, each with the same catalyst loading volume, and set up a heater before the gas flow inlet of each reactor; the reactors are connected in parallel to perform isopentane dehydrogenation operation; During the stable operation of isopentane dehydrogenation, the inlet gas flow conditions for each reactor include: Temperature 520-550℃, pressure ≤0.2MPa (gauge pressure, the same below), 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℃. Mode B: Set up 2-3 parallel reactor series, each reactor series includes two reactors I and II connected in series, where reactor I is the first reactor, and the catalyst loading volume ratio in reactors I and II is 1:(0.8-1.2); Each reactor is equipped with a heating furnace before its gas flow inlet; During the stable operation of isopentane dehydrogenation, the inlet gas flow conditions for reactor I include: temperature 520-550℃, pressure ≤0.2MPa, and isopentane mass hourly space velocity (MSV) 1.5-3h⁻¹. -1 The 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℃. 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 include: 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℃. 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.

2. The process for preparing isopentene by dehydrogenation of isopentane as described in claim 1, characterized in that, The isopentane dehydrogenation unit is also equipped with a heat exchanger T1; the isopentane liquid flow and / or the isopentane-rich gas flow, the hydrogen flow and / or the hydrogen-rich gas flow and the carbon dioxide flow flow M, and the reaction gas flow P from the dehydrogenation reactor group, are subjected to 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.

3. The process for preparing isopentene by dehydrogenation of isopentane as described in claim 1, characterized in that, The alkane dehydrogenation catalyst is catalyst H: a high-temperature zirconium-aluminum composite oxide containing 0.8-1.5 wt% ZrO2 is used as a support, and the supported components are Pt 0.2-0.4 wt%, Sn 0.5-1 wt%, and K 0.65-1 wt% (based on elemental composition), and it is prepared through the following steps: (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. (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. The Pt-Sn-acid-K aqueous solution described in step (H-2) is an aqueous solution prepared from chloroplatinic acid, stannous chloride, hydrochloric acid, citric acid, and potassium nitrate. The preparation method is as follows: prepare an aqueous solution from chloroplatinic acid and hydrochloric acid, add stannous chloride, potassium nitrate, and citric acid, dissolve 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.

4. The process for preparing isopentene by dehydrogenation of isopentane as described in claim 1, characterized in that, The alkane dehydrogenation catalyst is catalyst J: It uses magnesium-modified alumina containing 1-2 wt% MgO as a support, and the supported components (based on elemental composition) are Pt 0.2-0.4 wt%, Sn 0.5-1 wt%, and K 0.65-1 wt%, and is prepared through the following steps: (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. (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. The Pt-Sn-EDTA-K aqueous solution described in step (J-2) is an aqueous solution prepared from diammonium nitrite, potassium stannate, EDTA, nitric acid, potassium hydroxide, and potassium nitrite. The preparation method is as follows: (A) Add potassium stannate to a potassium hydroxide aqueous solution with a pH of 10-12 and dissolve it. Add EDTA powder and nitric acid aqueous solution and react until clear. Add potassium hydroxide aqueous solution to adjust the pH to 6.0-6.

5. Add potassium nitrite or its aqueous solution and dissolve and / or mix well to obtain the Sn-EDTA-K aqueous solution. 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 kept at a temperature of 50-70℃. Add diammonium nitrite aqueous solution at 60-70℃ and stir well to obtain the Pt-Sn-EDTA-K aqueous solution.

5. The process for preparing isopentene by dehydrogenation of isopentane as described in claim 3 or 4, characterized in that, The reduction operation performed after the reactor is loaded with new catalyst H or catalyst J and before the dehydrogenation operation, as well as the final reduction operation in the in-situ catalyst regeneration process, are all performed under the following conditions: using a hydrogen gas stream 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℃.

6. The process for preparing isopentene by dehydrogenation of isopentane as described in claim 3 or 4, characterized in that, The in-situ regeneration process performed on the reactor or series of reactors loaded with catalyst H or catalyst J after the dehydrogenation operation includes carbonization, oxychlorination, and reduction operations. When performing charcoal burning, an oxygen-containing airflow with an O2 molar content of 0.5-0.8% is used, which is obtained by mixing nitrogen gas with air, and the space velocity is 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 airflows decreases to below 0.1%. When performing oxychlorination, an oxygen-containing air stream with an O2 molar content of 4-15% is used, which is obtained by mixing a nitrogen stream with an air stream, and the space velocity is 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, as well as dichloroethane, trichloroethane, carbon tetrachloride, or tetrachloroethylene, are added to the oxygen-containing flow, or a 40-50% chloroacetic acid aqueous solution is directly added before the heater or furnace. The inlet temperature of the gas flow is controlled at 500-520℃, and the amount of chloride added is 0.4-0.6 wt% of the catalyst in the reactor, calculated as chlorine. The water-chlorine molar ratio is controlled at 5-8, and the addition time is 3-6 hours. Then, the inlet temperature of the oxygen-containing flow without adding water and chloride is controlled at 520-540℃, and the treatment is carried out for 4-6 hours, with the reactor outlet gas flow temperature above 515℃.

7. The process for preparing isopentene by dehydrogenation of isopentane as described in claim 4, characterized in that, The reactor or series loaded with catalyst J undergoes an in-situ regeneration process under chlorine-free conditions after the dehydrogenation operation is cut off, including carbonization, oxidation, and reduction operations. When performing charcoal production, an oxygen-containing airflow with an O2 molar content of 0.5-0.8% is used, which is obtained by mixing a nitrogen stream with a dry air stream, and the space velocity is 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 O2 molar content difference between the outlet and inlet airflows is reduced to below 0.1%, and then extended for another 2-4 hours. Oxidation operations are performed using an oxygen-containing airflow or dry airflow with an O2 molar content of 10-20%, obtained by mixing a nitrogen stream with a dry air stream, at a space velocity of 500-2000 h⁻¹. -1 Pressure 0.02-0.2MPa, bed initial temperature 480-520℃, airflow inlet temperature 500-520℃, oxidation treatment 6-10h.

8. The process for preparing isopentene by dehydrogenation of isopentane as described in claim 1, characterized in that, 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 isopentene dehydrogenated gas stream Q after it has been cooled by heat exchanger T1, cooling it to (-20 to -10)℃ and pressurizing it to 0.5-1MPa, separating water and obtaining a gas stream G0 containing H2, CH4, CO2, ethane, and ethylene, and a hydrocarbon-containing liquid stream L0. The hydrocarbon-containing liquid stream L0 enters the light stripping tower D0, and a low-boiling gas stream G-1 containing C4 and below components and 3-methyl-1-butene is separated from the top of the tower, and a hydrocarbon-containing liquid stream L is separated from the bottom of the tower. The pressure swing adsorption unit includes the PSA-H2 unit, which separates hydrogen from gas flow G0, and the alkane-alkene separation unit PSA-CH-1 and the alkane separation unit PSA-CH-2. The gas stream G0 is separated in the PSA-H2 unit to obtain a hydrogen gas stream G-H2 with a molar purity of over 98%, part of which is recycled to the isopentane dehydrogenation unit and the remainder is supplied externally. A low-hydrogen gas stream G-2 containing CH4, ethane, ethylene, and CO2 is also obtained and used as fuel gas in the heating furnace; The vaporized hydrocarbon-containing liquid stream L is separated into three parts by the PSA-CH-1 unit: alkane stream GL-1, isopentene stream Z-2, and heterohydrocarbon stream Z-3. The alkane stream GL-1 is further separated by the PSA-CH-2 unit to obtain two parts: isopentane stream Y-1 and n-pentane stream Z-1. The isopentane stream Y-1 is recycled in the isopentane dehydrogenation unit.

9. A process for preparing tert-amyl alcohol, characterized in that, Using isopentenylene and water obtained from the isopentenylene separation unit in the process of preparing isopentenylene by dehydrogenation of isopentenylene as described in claim 1 as raw materials, it is mixed with recycled residual isopentenylene 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 stream X is subjected to a liquid-phase reaction of isopentenylene hydration to tert-amyl alcohol in a fixed-bed hydration reactor packed with macroporous sulfonic acid type cationic resin catalyst. The circulating solvent stream X-4 contains ethylene glycol monobutyl ether as solvent. The stream obtained from the hydration reaction is then separated into residual isopentenylene stream X-1 by passing it through a light-removal tower D3 and a distillation tower D4. A mixture of tert-amyl alcohol and water, stream X-20, and a circulating solvent stream X-4 are prepared. The tert-amyl alcohol and water mixture stream X-20 is then separated into tert-amyl alcohol stream X-2 and water stream X-3 or an aqueous solution containing tert-amyl alcohol, by azeotropic distillation, extractive distillation, or extract-azeotropic distillation. The remaining isopentenyl alcohol stream X-1, water stream X-3 or the aqueous solution containing tert-amyl alcohol X-3, and circulating solvent stream X-4 are all recycled for the hydration reaction. The operating conditions of the hydration reactor include: a molar ratio of water, isopentenyl alcohol, and ethylene glycol monobutyl ether in stream X of (1-3):1:(3-8), and a liquid hourly space velocity (LISH) of 0.3-1.2 h⁻¹. -1 The feed temperature is 40-65℃, the pressure is 0.5-1.0MPa, and the discharge temperature is 65-80℃. The resulting tert-amyl alcohol liquid is cooled by X-2 to obtain the tert-amyl alcohol product.

10. The process for preparing tert-amyl alcohol as described in claim 9, characterized in that, The hydration reactor is equipped with an external circulation system, which reuses 40-70% of the volume or mass flow rate of the effluent as circulating material stream X-5 and mixes it with stream X before it re-enters the hydration reactor. The catalyst bed of the hydration reactor is equipped with a coil that internally circulates 60-65℃ circulating water. The remaining isopentenene stream X-1, which is required to be flow rate, is sent back to the cooling and separation system of the isopentenene separation unit and combined with the hydrocarbon-containing stream L0 for treatment and utilization. The total molar content of 2-methyl-2-butene and 2-methyl-1-butene in the remaining isopentenene stream X-1 is maintained above 90%. The extractant used in the extractive distillation method or extractive-azeotropic distillation method is cyclohexane or n-hexane. The circulating solvent stream X-4 is periodically fractionated and purified.

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