Hydrogenation catalyst and preparation method thereof, and method and system for continuously producing isopentenol

Abstract

The invention provides a hydrogenation catalyst and a preparation method thereof, and a method and a system for continuously producing isopentenol. The hydrogenation catalyst comprises a carrier, an active component and an auxiliary component, wherein the active component and the auxiliary component are loaded on the carrier; the method for continuously producing the isopentenol comprises the following steps: feeding a raw material flow containing hydrogen and 3-methyl-3-butene-1-ol into a reactor from an inlet in the bottom of the reactor, and obtaining a reacted material flow at an outlet in the top of the reactor; carrying out gas-liquid separation to obtain circulating hydrogen and a liquid-phase material flow; one part of the circulating hydrogen is circulated to the reactor, and the rest part is moved out of the reaction zone; a portion of the liquid phase stream is recycled to the reactor, and the remaining portion is collected as a product. When the catalyst is used for continuously producing the isopentenol, the conversion rate of the 3-methyl-3-butene-1-ol is 30%, the selectivity of the isopentenol can be up to 96%, and the catalyst can continuously and stably run for more than 1200 hours.

Description

Technical Field

[0001] The present invention relates to the technical field of preparation of isopentenol, and more particularly to a hydrogenation catalyst and a preparation method thereof, and a method and system for continuously producing isopentenol. Background Art

[0002] Isopentenol, i.e. 3-methyl-2-butene-1-ol, is an important organic fine chemical intermediate, and is widely used in the fields of pesticides, vitamins, water reducing agents, synthetic spices, etc. It is a double bond shift that 3-methyl-3-butene-1-ol isomerizes to generate isopentenol, which can be understood as a two-step reaction of hydrogenation and dehydrogenation. Therefore, the isomerization reaction is carried out under the hydrogen state, and no matter the raw material 3-methyl-3-butene-1-ol or the product isopentenol is likely to be hydrogenated to become isopentane, isopentanol, and there is also isoprene generated by dehydration, and it is difficult to obtain a higher product yield.

[0003] Chinese invention patent CN110759810A discloses a complex catalyst with a chelate structure formed by a bisphosphine ligand and a transition metal palladium compound, which catalyzes the double bond isomerization of 3-methyl-3-butene-1-ol and inhibits the formation of by-products such as isopentane, isopentanol, and isoprene, thereby improving the conversion rate of raw materials and the selectivity of products. The reaction process is characterized by using an intermittent reaction, and the catalyst 1,3-bis(diphenylphosphine)propane-palladium acetate chelate complex is a homogeneous catalyst.

[0004] Chinese invention patent CN109721470A discloses a method for continuously preparing isopentenol, which is as follows: 3-methyl-3-butene-1-ol is used as a raw material, sodium methoxide, sodium ethoxide, and sodium isopentenol are dissolved in the raw material as catalysts, the reaction pressure is preferably 8-12 MPa, and the reaction temperature is preferably 170-240° C. The amount of catalyst used is large and the reaction conditions are high temperature and high pressure, and the catalyst needs to be separated later.

[0005] Chinese invention patent CN102675048A discloses a method for synthesizing isopentenol using a composite catalyst, using 3-methyl-3-butene-1-ol as raw material, a kettle reaction, Raney nickel and solid superacid SO4 2- / TiO2, SO4 2- / ZrO2, SO4 2- / SiO2 composite catalyst, hydrogenation, so that the double bond of 3-methyl-3-butene-1-ol isomerizes to form isopentenol.

[0006] Chinese invention patent CN101544538A discloses a kettle-type preparation of isopentenol, in which 3-methyl-3-butene-1-ol is isomerized in the presence of a catalyst in a hydrogen-containing atmosphere to produce isopentenol, the reaction pressure is normal pressure, the reaction temperature is 50-90°C, the reaction time is 25-40 minutes, the catalyst is supported by Al2O3, and the active component metal Pd is loaded, the Pd content in the catalyst is 0.1-2.0wt%, the catalyst particle size is 80-200 mesh, and the weight ratio of the raw material to the catalyst is 30-100:1.

[0007] In summary, the currently disclosed methods for preparing isopentenol from 3-methyl-3-butene-1-ol are mostly batch reactions in kettles, and some of them involve catalyst separation problems, and cannot prepare isopentenol continuously and highly selectively. There is still much room for improvement in the catalysts and reaction processes for preparing isopentenol. Summary of the invention

[0008] In order to solve the technical problems existing in the prior art, the present invention provides a hydrogenation catalyst and a preparation method thereof, and a method and system for continuously producing isopentenol.

[0009] The invention prepares a modified palladium aluminum oxide catalyst, adjusts the ratio of auxiliary agents, performs catalyst crystallization treatment in an impregnation process, and obtains highly dispersed active components after drying and calcining.

[0010] The prior art for preparing 3-methyl-3-butene-1-ol is a kettle reaction, which cannot continuously prepare isopentenol with high selectivity, and some of them also involve the problem of catalyst separation. The system of the present invention adopts a bubbling bed process, the catalyst is fixedly loaded in the reaction tube, the material enters the reaction tube from the bottom, and the catalyst separation is not required. Isopentenol can be continuously produced, the conversion rate of 3-methyl-3-butene-1-ol is 30%, the selectivity of isopentenol can reach up to 96%, and the catalyst can operate continuously and stably for more than 1200 hours.

[0011] One of the objects of the present invention is to provide a hydrogenation catalyst, which comprises a carrier and an active component and an auxiliary component loaded on the carrier; the carrier is preferably at least one of aluminum oxide and silicon dioxide; the active component preferably comprises palladium, and optionally, also comprises at least one of ruthenium and nickel; the auxiliary component is preferably at least one of the lanthanides, more preferably lanthanum oxide and cerium oxide.

[0012] The raw powder of at least one of aluminum oxide and silicon dioxide is used as raw material, a binder is added, and the binder is at least one of water, nitric acid and phosphoric acid. The mixture is mechanically mixed, kneaded and extruded. The wet carrier is dried at 80-120° C. for 8-16 hours and calcined at 700-800° C. for 3-6 hours to obtain a carrier.

[0013] In a preferred embodiment of the present invention,

[0014] Taking the total mass of the hydrogenation catalyst as 100%, the mass fraction of the carrier is 75-99.3wt.%, preferably 97.5%-99.3wt.%; the mass fraction of the active component is 0.1-5wt.%, preferably 0.2-1wt.%; the mass fraction of the auxiliary component is 0.5-20wt.%, preferably 0.5-2wt.%; wherein the mass fraction of palladium is not less than 0.1wt.%, preferably not less than 0.2wt.%.

[0015] In a preferred embodiment of the present invention,

[0016] The mass ratio of lanthanum oxide to cerium oxide is (1-9):1, preferably (1-2):1; and / or,

[0017] The content of chlorine in the hydrogenation catalyst is less than 1% of the total mass of the hydrogenation catalyst. If the hydrogenation catalyst contains chlorine, the product will contain dissolved chlorine, and the product isopentenol will avoid the introduction of chloride ions.

[0018] In a preferred embodiment of the present invention,

[0019] The specific surface area of ​​the hydrogenation catalyst is 80 to 230 m 2 / g, preferably 100 to 200 m 2 / g; and / or,

[0020] The average pore size of the hydrogenation catalyst is 5 to 20 nm, preferably 8 to 16 nm; and / or,

[0021] The pore volume of the hydrogenation catalyst is 0.3 to 0.8 mL / g, preferably 0.5 to 0.8 mL / g.

[0022] A second object of the present invention is to provide a method for preparing a hydrogenation catalyst, comprising:

[0023] The hydrogenation catalyst is first prepared by loading the components including the active component and the auxiliary component on the carrier to obtain a catalyst precursor, then subjecting the obtained catalyst precursor to crystallization treatment in a crystallization kettle, then transferring it to an oven for drying, and finally calcining it in an air atmosphere.

[0024] In a preferred embodiment of the present invention,

[0025] The crystallization temperature is 40 to 80° C. and the crystallization time is 4 to 24 hours; and / or,

[0026] The drying temperature is 80 to 120°C and the drying time is 4 to 12 hours; and / or,

[0027] The calcination temperature is 350 to 550° C. and the calcination time is 4 to 12 hours; and / or,

[0028] The precursor of the active component is a chlorine-free active component precursor, preferably at least one of palladium nitrate and palladium acetate; and / or,

[0029] The precursor of the auxiliary agent component is a soluble salt of lanthanum or cerium, preferably at least one of lanthanum nitrate and cerium nitrate;

[0030] Preferably,

[0031] The crystallization temperature is 50 to 70° C. and the crystallization time is 6 to 12 hours; and / or,

[0032] The drying temperature is 80-100°C and the drying time is 8-12 hours; and / or,

[0033] The calcination temperature is 400-500°C, and the calcination time is 4-8 hours.

[0034] The third object of the present invention is to provide a method for continuously producing isopentenol, comprising the following steps:

[0035] (1) feeding a raw material stream including hydrogen and 3-methyl-3-butene-1-ol into a reactor from a bottom inlet of the reactor, and obtaining a post-reaction stream at a top outlet of the reactor;

[0036] (2) separating the post-reaction logistics obtained in step (1) into gas and liquid to obtain circulating hydrogen and liquid logistics;

[0037] (3) circulating a portion of the circulating hydrogen obtained in step (2) to the reactor, and removing the remaining portion from the reaction zone;

[0038] (4) A portion of the liquid phase stream obtained in step (2) is recycled to the reactor, and the remaining portion is collected as the product.

[0039] In step (3), low-boiling isoprene will be produced during the isomerization process, and part of the recycled hydrogen will be removed from the reaction zone. The removed part can be used to collect low-boiling components, which can increase added value if separated later;

[0040] In step (4), after the product enters the gas-liquid separator, it enters the collection tank through a pipeline, a branch is separated from the pipeline, a circulation pump is connected, and the circulation volume is controlled by setting parameters of the circulation pump.

[0041] In a preferred embodiment of the present invention,

[0042] Step (1),

[0043] The hydrogen in the feed stream is at least one of newly added hydrogen and recycled hydrogen in step (2);

[0044] The 3-methyl-3-butene-1-ol in the raw material stream is at least one of the newly added 3-methyl-3-butene-1-ol and the 3-methyl-3-butene-1-ol in the circulating liquid phase stream of step (4);

[0045] The feed stream may or may not contain isopentenol (the fresh feed stream does not contain isopentenol), and the isopentenol is the isopentenol in the circulating liquid phase stream of step (4).

[0046] In a preferred embodiment of the present invention,

[0047] Step (1),

[0048] The molar ratio of the hydrogen to the 3-methyl-3-butene-1-ol is (0.001-1):1, preferably (0.03-0.1):1;

[0049] The liquid phase volume space velocity of the 3-methyl-3-butene-1-ol is 0.1h -1 ~10h -1 , preferably 2h -1 ~10h -1 , more preferably 3h -1 ~9h -1 The raw materials are fed from the bottom of the reactor. The reaction time is the time from the bottom contacting the catalyst to leaving the catalyst. The space velocity here is the volume space velocity (h -1 ),

[0050]

[0051] The reactor is a bubbling reactor, more preferably a bubbling fixed bed reactor;

[0052] The reaction pressure is 0.1 MPa to 6.0 MPa, preferably 0.1 to 1.0 MPa;

[0053] The reaction temperature is 40°C to 100°C.

[0054] In a preferred embodiment of the present invention,

[0055] The mass ratio of the circulating hydrogen gas circulated to the reactor in step (3) to the circulating hydrogen gas removed from the reaction zone is (90-100):(0-10), preferably (95-98):(2-5); and / or,

[0056] The mass ratio of the liquid phase material recycled to the reactor in step (4) to the liquid phase flow collected as the product is (50-98): (2-50), preferably (50-70): (30-50).

[0057] In a preferred embodiment of the present invention,

[0058] The reactor is filled with a hydrogenation catalyst;

[0059] The hydrogenation catalyst is preferably the above-mentioned hydrogenation catalyst or the hydrogenation catalyst obtained by the above-mentioned preparation method.

[0060] The fourth object of the present invention is to provide a system for continuously producing isopentenol, the system comprising a reactor, a gas-liquid separator, and a condenser, wherein a hydrogen inlet is arranged at the bottom of the reactor and connected to a hydrogen inlet pipeline; a raw material inlet is arranged at the lower part of the side of the reactor and connected to the raw material inlet pipeline; a discharge port is arranged at the top of the reactor, the discharge port is connected to a pipeline at the top of the gas-liquid separator, and a condenser is arranged on the pipeline; the reaction stream of the reactor is discharged from the top and sent to the gas-liquid separator through the condenser, a gas phase outlet is arranged at the top of the gas-liquid separator, and a liquid phase outlet is arranged at the bottom; the gas phase outlet is connected to the hydrogen inlet pipeline, and an outlet for transferring hydrogen is arranged on the pipeline; the liquid phase outlet is connected to the raw material inlet pipeline, and a product outlet is arranged on the pipeline;

[0061] Preferably, a gas distributor is arranged at the bottom of the reactor; the inlet hydrogen enters the reactor through the gas distributor;

[0062] The reactor is a bubbling bed reactor, and the catalyst is fixedly loaded in the reaction tube of the reactor;

[0063] After gas-liquid separation, the circulating hydrogen is connected to the gas distributor via a pipeline, part of the liquid is connected to the reactor feed port via a pipeline as circulating material, and part is collected as a product.

[0064] A fifth object of the present invention is to provide isopentenol obtained by the above-mentioned preparation method.

[0065] Compared with the prior art, the present invention has the following beneficial effects:

[0066] The invention prepares a modified palladium aluminum oxide catalyst, adjusts the ratio of auxiliary agents, performs catalyst crystallization treatment in an impregnation process, and obtains highly dispersed active components after drying and calcining.

[0067] The prior art for preparing 3-methyl-3-butene-1-ol is a kettle reaction, which cannot continuously prepare isopentenol with high selectivity, and some of them also involve the problem of catalyst separation. The system of the present invention adopts a bubbling bed process, the catalyst is fixedly loaded in the reaction tube, the material enters the reaction tube from the bottom, and the catalyst separation is not required. Isopentenol can be continuously produced, the conversion rate of 3-methyl-3-butene-1-ol is 30%, the selectivity of isopentenol can reach up to 96%, and the catalyst can operate continuously and stably for more than 1200 hours. BRIEF DESCRIPTION OF THE DRAWINGS

[0068] Figure 1 It is a schematic diagram of the system of the present invention;

[0069] 1-reactor, 2-gas-liquid separator, 3-condenser, 4-gas distributor, 5-fresh feed, 6-fresh hydrogen, 7-circulating feed, 8-circulating hydrogen, 9-relaxing hydrogen, 10-product;

[0070] Figure 2 This is a TEM image of the catalyst prepared in Example 1 after reduction;

[0071] Figure 3 The pore size and volume distribution diagram of the catalyst prepared in Example 1;

[0072] Figure 4 This is a TEM image of the catalyst prepared in Comparative Example 2 after reduction;

[0073] Figure 5 This is the pore size and volume distribution diagram of the catalyst prepared in Comparative Example 2. DETAILED DESCRIPTION

[0074] The present invention is described in detail below in conjunction with specific drawings and embodiments. It is necessary to point out that the following embodiments are only used to further illustrate the present invention and cannot be understood as limiting the scope of protection of the present invention. Some non-essential improvements and adjustments made to the present invention by those skilled in the art based on the content of the present invention still fall within the scope of protection of the present invention.

[0075] The raw materials used in the examples are all conventional commercially available raw materials.

[0076] Test method:

[0077] TEM test: A Tecnai 20S-TWIN transmission electron microscope from FEI, USA, was used, with an operating voltage of 200 kV. The catalyst was reduced with NaBH4 before the test, and then the sample was dispersed with ethanol.

[0078] Nitrogen physical adsorption test: The specific surface area and pore structure of the sample were determined by N2 physical adsorption, using the Tristar-3000 automatic adsorption instrument of Micromeritics, USA. The sample was vacuum degassed at 200°C for 2h, and the catalyst sample was degassed at 400°C to remove physically adsorbed water molecules and impurities in the sample, and then N2 physical adsorption was performed in liquid nitrogen. The specific surface area was calculated by the Brunauer-Emmett-Teller (BET) method, and the pore size distribution was calculated by the Barrett-Joyner-Halenda (BJH) method.

[0079] Figure 1 The system schematic diagram of the present invention; wherein, 1-reactor, 2-gas-liquid separator, 3-condenser, 4-gas distributor, 5-fresh material, 6-fresh hydrogen, 7-circulating material, 8-circulating hydrogen, 9-relaxed hydrogen, 10-product; the system comprises a reactor 1, a gas-liquid separator 2, and a condenser 3; a hydrogen inlet is arranged at the bottom of the reactor 1, connected to a hydrogen inlet pipeline; a raw material inlet is arranged at the lower part of the side of the reactor, connected to a raw material inlet pipeline; a discharge port is arranged at the top of the reactor, and the discharge port is connected to a pipeline at the upper part of the gas-liquid separator 2, and a condenser 3 is arranged on the pipeline; the reaction stream of the reactor is discharged from the top, and after condensation, The reactor is sent to a gas-liquid separator, a gas phase outlet is arranged on the top of the gas-liquid separator, and a liquid phase outlet is arranged on the bottom; the gas phase outlet is connected to the hydrogen inlet pipeline, and an outlet for transferring hydrogen is arranged on the pipeline; the liquid phase outlet is connected to the raw material inlet pipeline, and a product outlet is arranged on the pipeline; preferably, a gas distributor 4 is arranged at the bottom of the reactor; the inlet hydrogen enters the reactor through the gas distributor; the reactor is a bubbling bed reactor, and the catalyst is fixedly loaded in the reaction tube of the reactor; after gas-liquid separation, the circulating hydrogen is connected to the gas distributor 4 through a pipeline, a part of the liquid is connected to the reactor feed port through a pipeline as a circulating material, and a part is collected as a product.

[0080] Example 1

[0081] Preparation of Pd-La2O3-CeO2 / Al2O3 catalyst:

[0082] Alumina powder is used as raw material, and a binder is added, wherein the binder is at least one of water, nitric acid and phosphoric acid; the mixture is mechanically mixed, kneaded and extruded. The wet carrier is dried at 90°C for 12 hours and calcined at 750°C for 4 hours to obtain the carrier;

[0083] The obtained carrier was impregnated with equal volumes of solutions containing palladium nitrate, lanthanum nitrate and cerium nitrate, and then crystallized at 50°C for 6 hours, dried at 90°C for 12 hours, and calcined at 450°C for 4 hours to obtain a 1.0wt.% Pd-0.5wt.% La2O3-0.5wt.CeO2% / Al2O3 carrier in the shape of a clover leaf.

[0084] The prepared catalyst 0.35wt.%Pd / Al2O3 was characterized by nitrogen physical adsorption, and the specific surface area of ​​the catalyst was 150.6m 2 / g, the pore volume is 0.59mL / g, and the average pore diameter of the catalyst is 15.6nm.

[0085] The loading amount of 1.0wt.% Pd-0.5wt.% La2O3-0.5wt.CeO2% / Al2O3 catalyst in the fixed bed reactor was 10mL. After the catalyst was reduced, the temperature was lowered to 60°C, the device pressure was 0.5MPa, and the feed was added and the reactor was started. The feed rate of the raw material stream of 3-methyl-3-butene-1-ol is 1.0 mL / min, the feed rate of the circulating pump is 0.5 mL / min (wherein, the mass ratio of the liquid phase material circulated to the reactor to the liquid phase stream collected as the product is 50:50), the fresh raw material and the circulating product are mixed and sent into the reactor from the bottom inlet of the reactor, and the post-reaction stream is obtained at the top outlet of the reactor; at the same time, hydrogen is introduced into the reaction tube from the gas distributor at the bottom of the reactor by fresh hydrogen and circulating hydrogen, the molar ratio of hydrogen to the 3-methyl-3-butene-1-ol is 0.045:1, the hydrogen flow rate is 1.45 mL / min, and the circulating hydrogen flow rate is 9.0 mL / min (wherein, the volume ratio of the circulating hydrogen circulated to the reactor to the circulating hydrogen removed from the reaction zone is 95:5). After sampling, the sample is analyzed by gas chromatography, and the conversion rate of 3-methyl-3-butene-1-ol is 31.2%, and the selectivity of isopentenol is 93.2%.

[0086] Example 2

[0087] The difference from Example 1 is that the molar ratio of hydrogen to the 3-methyl-3-butene-1-ol is 0.001;

[0088] Except for the above differences, other conditions of Example 2 are the same as those of Example 1.

[0089] After the product was sampled, gas chromatography analysis was performed, and the conversion rate of 3-methyl-3-butene-1-ol was 8.0%, and the selectivity of isopentenol was 92.9%.

[0090] Example 3

[0091] The difference from Example 1 is that the molar ratio of hydrogen to the 3-methyl-3-butene-1-ol is 0.1;

[0092] Except for the above differences, other conditions of Example 2 are the same as those of Example 1.

[0093] After the product was sampled, gas chromatography analysis was performed, and the conversion rate of 3-methyl-3-butene-1-ol was 40.5%, and the selectivity of isopentenol was 87.2%.

[0094] Example 4

[0095] The preparation of the carrier is the same as in Example 1.

[0096] The obtained carrier was impregnated with equal volumes of solutions containing palladium nitrate, lanthanum nitrate and cerium nitrate, and then crystallized at 70°C for 12 hours, dried at 100°C for 8 hours, and calcined at 400°C for 8 hours to obtain a 0.5wt.% Pd-0.5wt.% La2O3-0.5wt.CeO2% / Al2O3 carrier in the shape of a clover leaf.

[0097] The loading amount of 0.5wt.% Pd-0.5wt.% La2O3-0.5wt.CeO2% / Al2O3 catalyst in the fixed bed reactor was 10mL. After the catalyst was reduced, the temperature was lowered to 60°C, the device pressure was 0.5MPa, and the feed was added and the reactor was started. The feed rate of the raw material stream of 3-methyl-3-butene-1-ol is 1.0 mL / min, the feed rate of the circulating pump is 0.7 mL / min (wherein, the mass ratio of the liquid phase material circulated to the reactor to the liquid phase stream collected as the product is 70:30), the fresh raw material and the circulating product are mixed and sent into the reactor from the bottom inlet of the reactor, and the post-reaction stream is obtained at the top outlet of the reactor; at the same time, hydrogen is introduced into the reaction tube from the gas distributor at the bottom of the reactor by fresh hydrogen and circulating hydrogen, the molar ratio of hydrogen to the 3-methyl-3-butene-1-ol is 0.045:1, the hydrogen flow rate is 1.18 mL / min, and the circulating hydrogen flow rate is 9.0 mL / min (wherein, the volume ratio of the circulating hydrogen circulated to the reactor to the circulating hydrogen removed from the reaction zone is 98:2). After sampling, the sample is analyzed by gas chromatography, and the conversion rate of 3-methyl-3-butene-1-ol is 30.8%, and the selectivity of isopentenol is 92.9%.

[0098] Example 5

[0099] The preparation of the catalyst is the same as in Example 4;

[0100] The loading amount of 0.5wt.% Pd-0.5wt.% La2O3-0.5wt.CeO2% / Al2O3 catalyst in the fixed bed reactor was 10mL. After the catalyst was reduced, the temperature was lowered to 100°C, the device pressure was 0.1MPa, and the feed was added and the reactor was started. The feed rate of the raw material stream of 3-methyl-3-butene-1-ol is 1.0 mL / min, the feed rate of the circulating pump is 0.7 mL / min (wherein, the mass ratio of the liquid phase material circulated to the reactor to the liquid phase stream collected as the product is 70:30), the fresh raw material and the circulating product are mixed and sent into the reactor from the bottom inlet of the reactor, and the post-reaction stream is obtained at the top outlet of the reactor; at the same time, hydrogen is introduced into the reaction tube from the gas distributor at the bottom of the reactor by fresh hydrogen and circulating hydrogen, the molar ratio of hydrogen to the 3-methyl-3-butene-1-ol is 0.045:1, the hydrogen flow rate is 1.45 mL / min, and the circulating hydrogen flow rate is 9.0 mL / min (wherein, the volume ratio of the circulating hydrogen circulated to the reactor to the circulating hydrogen removed from the reaction zone is 95:5). After sampling, the sample is analyzed by gas chromatography, and the conversion rate of 3-methyl-3-butene-1-ol is 38.4%, and the selectivity of isopentenol is 88.2%.

[0101] Example 6

[0102] The preparation of the catalyst is the same as in Example 4.

[0103] The loading amount of 0.5wt.% Pd-0.5wt.% La2O3-0.5wt.CeO2% / Al2O3 catalyst in the fixed bed reactor was 10mL. After the catalyst was reduced, the temperature was lowered to 50°C, the device pressure was 0.1MPa, and the feed was added and the reactor was started. The feed rate of the raw material stream of 3-methyl-3-butene-1-ol is 0.66 mL / min, the feed rate of the circulating pump is 0.46 mL / min (wherein, the mass ratio of the liquid phase material circulated to the reactor to the liquid phase stream collected as the product is 70:30), the fresh raw material and the circulating product are mixed and sent into the reactor from the bottom inlet of the reactor, and the post-reaction stream is obtained at the top outlet of the reactor; at the same time, hydrogen is introduced into the reaction tube from the gas distributor at the bottom of the reactor by fresh hydrogen and circulating hydrogen, the molar ratio of hydrogen to the 3-methyl-3-butene-1-ol is 0.045:1, the hydrogen flow rate is 0.96 mL / min, and the circulating hydrogen flow rate is 6.0 mL / min (wherein, the volume ratio of the circulating hydrogen circulated to the reactor to the circulating hydrogen removed from the reaction zone is 95:5). After sampling, the sample is analyzed by gas chromatography, and the conversion rate of 3-methyl-3-butene-1-ol is 29.6%, and the selectivity of isopentenol is 95.0%.

[0104] Example 7

[0105] The preparation of the catalyst is the same as in Example 4;

[0106] The loading amount of 0.5wt.% Pd-0.5wt.% La2O3-0.5wt.CeO2% / Al2O3 catalyst in the fixed bed reactor was 10mL. After the catalyst was reduced, the temperature was lowered to 50°C, the device pressure was 0.1MPa, and the feed was added and the reactor was started. The feed rate of the raw material flow of 3-methyl-3-butene-1-ol is 1.5 mL / min, the feed rate of the circulating pump is 1.05 mL / min (wherein, the mass ratio of the liquid phase material circulated to the reactor to the liquid phase flow collected as the product is 70:30), the fresh raw material and the circulating product are mixed and sent into the reactor from the bottom inlet of the reactor, and the post-reaction flow is obtained at the top outlet of the reactor; at the same time, hydrogen is introduced into the reaction tube from the gas distributor at the bottom of the reactor by fresh hydrogen and circulating hydrogen, the molar ratio of hydrogen to the 3-methyl-3-butene-1-ol is 0.045:1, the hydrogen flow rate is 2.17 mL / min, and the circulating hydrogen flow rate is 13.5 mL / min (wherein, the volume ratio of the circulating hydrogen circulated to the reactor to the circulating hydrogen removed from the reaction zone is 95:5). After sampling, the sample is analyzed by gas chromatography, and the conversion rate of 3-methyl-3-butene-1-ol is 25.6%, and the selectivity of isopentenol is 94.6%.

[0107] Example 8

[0108] Preparation of catalyst:

[0109] Alumina powder is used as raw material, and a binder is added, wherein the binder is at least one of water, nitric acid and phosphoric acid; the mixture is mechanically mixed, kneaded and extruded. The wet carrier is dried at 90°C for 12 hours and calcined at 750°C for 4 hours to obtain the carrier;

[0110] The obtained carrier was impregnated with an equal volume of a solution containing palladium acetate, and the modifiers were lanthanum nitrate and cerium nitrate (the mass ratio of lanthanum oxide to cerium oxide was 2:1), and then crystallized at 50°C for 6 hours, dried at 90°C for 12 hours, and calcined at 450°C for 4 hours. The shape of the carrier was clover.

[0111] The prepared catalyst 0.35wt.%Pd-0.5wt.%La-0.25wt.%Ce / Al2O3 was characterized by nitrogen physical adsorption. The specific surface area of ​​the catalyst was 198.5m 2 / g, pore volume is 0.63mL / g, and the average pore size of the catalyst is 12.7nm. TEM of the reduced catalyst is as follows Figure 2 As shown; the catalyst pore size and volume distribution is as follows Figure 3As shown. After the reduction of Pd / Al2O3 modified with La and Ce dual additives, it can be seen from the TEM image that the Pd grains are small and evenly dispersed, with almost no Pd agglomeration; combined with the pore size and pore volume distribution diagram, the average pore size distribution of the modified catalyst is concentrated at 10nm, showing a single-pore distribution. For hydrogenation reactions with multiple reaction pathways, the single-pore distribution is more conducive to the desorption of reaction products and the highly selective acquisition of target products.

[0112] The loading amount of Pd-La-Ce / Al2O3 catalyst in the fixed bed reactor was 10 mL. After the catalyst was reduced, the temperature was lowered to 75°C, the device pressure was 0.5 MPa, and the feed was added and the reactor was started. The feed rate of the raw material stream of 3-methyl-3-butene-1-ol is 1.0 mL / min, the feed rate of the circulating pump is 0.7 mL / min (wherein, the mass ratio of the liquid phase material circulated to the reactor to the liquid phase stream collected as the product is 70:30), the fresh raw material and the circulating product are mixed and sent into the reactor from the bottom inlet of the reactor, and the post-reaction stream is obtained at the top outlet of the reactor; at the same time, hydrogen is introduced into the reaction tube from the gas distributor at the bottom of the reactor by fresh hydrogen and circulating hydrogen, the molar ratio of hydrogen to the 3-methyl-3-butene-1-ol is 0.045:1, the hydrogen flow rate is 1.45 mL / min, and the circulating hydrogen flow rate is 9.0 mL / min (wherein, the volume ratio of the circulating hydrogen circulated to the reactor to the circulating hydrogen removed from the reaction zone is 95:5). After sampling, the sample is analyzed by gas chromatography, and the conversion rate of 3-methyl-3-butene-1-ol is 31.6%, and the selectivity of isopentenol is 96.0%.

[0113] Comparative Example 1

[0114] The preparation of the catalyst is the same as in Example 8;

[0115] The difference from Example 8 is that the fresh raw material and the recycled product are mixed and fed into the reactor from the upper part of the reactor, and the hydrogen from the fresh hydrogen and the recycled hydrogen enters the reaction tube from the upper part of the gas distributor at the bottom of the reactor;

[0116] Except for the above differences, other conditions of Comparative Example 1 are the same as those of Example 8.

[0117] After sampling, the sample was analyzed by gas chromatography, and the conversion rate of 3-methyl-3-butene-1-ol was 32.0%, and the selectivity of isopentenol was 61.8%.

[0118] Comparative Example 2

[0119] Preparation of catalyst:

[0120] The preparation of the carrier is the same as in Example 8;

[0121] The obtained support was impregnated with an equal volume of a solution containing palladium acetate, then crystallized at 50°C for 6 hours, dried at 90°C for 12 hours, and calcined at 450°C for 4 hours. The support was in the shape of a clover leaf. The prepared catalyst had a 0.35wt.% Pd / Al2O3. The catalyst was characterized by nitrogen physical adsorption, and the specific surface area of ​​the catalyst was 150.6m 2 / g, pore volume 0.59mL / g, average pore size distribution between 8nm and 20nm, showing a double pore distribution. After reduction, TEM Figure 4 It can be seen that the Pd crystals are obviously agglomerated and the particles are large. Combined with the pore size and volume distribution of the catalyst Figure 5 .

[0122] A fixed bed reactor was loaded with a catalyst of 0.35 wt.% Pd / Al2O3. Except for the catalyst, the other conditions for preparing isopentenol were the same as those in Example 8. After the product was sampled, gas chromatography was used to analyze the sample. The conversion rate of 3-methyl-3-butene-1-ol was 30.4%, and the selectivity of isopentenol was 75.1%.

[0123] Compared with Comparative Example 1 (3-methyl-3-butene-1-ol conversion rate is 32.0%, isopentenol selectivity is 61.8%), the 3-methyl-3-butene-1-ol conversion rate of Example 8 is 31.6%, and the isopentenol selectivity is 96.0%. The isopentenol selectivity is higher, proving that the method for continuous production of isopentenol of the present invention is more effective.

[0124] Compared with Comparative Example 2 (3-methyl-3-butene-1-ol conversion rate is 30.4%, isopentanol selectivity is 75.1%), the 3-methyl-3-butene-1-ol conversion rate of Example 8 is 31.6%, and the isopentanol selectivity is 96.0%. The isopentanol selectivity is higher, proving that the catalyst with lanthanum oxide and cerium oxide additives has better performance.

[0125] In Examples 1, 3 to 8, the conversion rate of 3-methyl-3-butene-1-ol is 25.6 to 40.5%, and the selectivity of isopentenol is 87.2 to 96%. The present invention adopts a bubbling bed process to prepare a modified palladium alumina catalyst, the catalyst is fixedly loaded in a reaction tube, and the material enters the reaction tube from the bottom, and there is no need to separate the catalyst. Isopentenol can be continuously produced, and the selectivity of isopentenol can reach up to 96%, and the catalyst can operate continuously and stably for more than 1200 hours.

Claims

1. A hydrogenation catalyst, comprising a carrier and an active component and an auxiliary component supported on the carrier; the carrier is preferably at least one of aluminum oxide and silicon dioxide; the active component preferably comprises palladium, and optionally, further comprises at least one of ruthenium and nickel; the auxiliary component is preferably at least one of lanthanides, more preferably lanthanum oxide and cerium oxide.

2. The hydrogenation catalyst according to claim 1, characterized in that: Taking the total mass of the hydrogenation catalyst as 100%, the mass fraction of the carrier is 75-99.3wt.%, preferably 97.5%-99.3wt.%; the mass fraction of the active component is 0.1-5wt.%, preferably 0.2-1wt.%; the mass fraction of the auxiliary component is 0.5-20wt.%, preferably 0.5-2wt.%; wherein the mass fraction of palladium is not less than 0.1wt.%, preferably not less than 0.2wt.%.

3. The hydrogenation catalyst according to claim 1, characterized in that: The mass ratio of lanthanum oxide to cerium oxide is (1-9):1, preferably (1-2):1; and / or, The content of chlorine in the hydrogenation catalyst is less than 1% of the total mass of the hydrogenation catalyst.

4. The hydrogenation catalyst according to any one of claims 1 to 3, characterized in that: The specific surface area of ​​the hydrogenation catalyst is 80 to 230 m 2 / g, preferably 100 to 200 m 2 / g; and / or, The average pore size of the hydrogenation catalyst is 5 to 20 nm, preferably 8 to 16 nm; and / or, The pore volume of the hydrogenation catalyst is 0.3 to 0.8 mL / g, preferably 0.5 to 0.8 mL / g.

5. A method for preparing the hydrogenation catalyst according to any one of claims 1 to 4, comprising: The hydrogenation catalyst is first prepared by loading the components including the active component and the auxiliary component on the carrier to obtain a catalyst precursor, then subjecting the obtained catalyst precursor to crystallization treatment in a crystallization kettle, then transferring it to an oven for drying, and finally calcining it in an air atmosphere.

6. The method for preparing a hydrogenation catalyst according to claim 5, characterized in that: The crystallization temperature is 40 to 80° C. and the crystallization time is 4 to 24 hours; and / or, The drying temperature is 80 to 120°C and the drying time is 4 to 12 hours; and / or, The calcination temperature is 350 to 550° C. and the calcination time is 4 to 12 hours; and / or, The precursor of the active component is a chlorine-free active component precursor, preferably at least one of palladium nitrate and palladium acetate; and / or, The precursor of the auxiliary agent component is a soluble salt of lanthanum or cerium, preferably at least one of lanthanum nitrate and cerium nitrate; Preferably, The crystallization temperature is 50 to 70° C. and the crystallization time is 6 to 12 hours; and / or, The drying temperature is 80-100°C and the drying time is 8-12 hours; and / or, The calcination temperature is 400-500°C, and the calcination time is 4-8 hours.

7. A method for continuously producing isopentenol, comprising the following steps: (1) feeding a raw material stream including hydrogen and 3-methyl-3-butene-1-ol into a reactor from a bottom inlet of the reactor, and obtaining a post-reaction stream at a top outlet of the reactor; (2) separating the post-reaction logistics obtained in step (1) into gas and liquid to obtain circulating hydrogen and liquid logistics; (3) circulating a portion of the circulating hydrogen obtained in step (2) to the reactor, and removing the remaining portion from the reaction zone; (4) A portion of the liquid phase stream obtained in step (2) is recycled to the reactor, and the remaining portion is collected as the product.

8. The method for continuously producing isopentenol as claimed in claim 7, characterized in that: Step (1), The hydrogen in the feed stream is at least one of newly added hydrogen and recycled hydrogen in step (2); and / or, The 3-methyl-3-butene-1-ol in the raw material stream is at least one of newly added 3-methyl-3-butene-1-ol and the 3-methyl-3-butene-1-ol in the circulating liquid phase stream of step (4); and / or, The raw material stream may or may not include isopentenol, and the isopentenol is the isopentenol in the circulating liquid phase stream of step (4).

9. The method for continuously producing isopentenol as claimed in claim 7, characterized in that: Step (1), The molar ratio of the hydrogen to the 3-methyl-3-butene-1-ol is (0.001-1):1, preferably (0.03-0.1):1; and / or, The liquid phase volume space velocity of the 3-methyl-3-butene-1-ol is 0.1h -1 ~10h -1 , preferably 2h -1 ~10h -1 , more preferably 3h -1 ~9h -1 and / or, The reactor is a bubbling reactor, more preferably a bubbling fixed bed reactor; and / or, The reaction pressure is 0.1 MPa to 6.0 MPa, preferably 0.1 to 1.0 MPa; and / or, The reaction temperature is 40°C to 100°C.

10. The method for continuously producing isopentenol as claimed in claim 7, characterized in that: The mass ratio of the circulating hydrogen gas circulated to the reactor in step (3) to the circulating hydrogen gas removed from the reaction zone is (90-100):(0-10), preferably (95-98):(2-5); and / or, The mass ratio of the liquid phase material recycled to the reactor in step (4) to the liquid phase flow collected as the product is (50-98): (2-50), preferably (50-70): (30-50).

11. The method for continuously producing isopentenol according to any one of claims 7 to 10, characterized in that: The reactor is filled with a hydrogenation catalyst; The hydrogenation catalyst is preferably the hydrogenation catalyst according to any one of claims 1 to 4 or the hydrogenation catalyst obtained by the preparation method according to claim 5 or 6.

12. A system for the method of continuously producing isopentenol according to any one of claims 7 to 11, the system comprising a reactor, a gas-liquid separator, and a condenser, wherein a hydrogen inlet is arranged at the bottom of the reactor, connected to a hydrogen inlet pipeline; a raw material inlet is arranged at the lower part of the side of the reactor, connected to a raw material inlet pipeline; a discharge port is arranged at the top of the reactor, the discharge port is connected to a pipeline at the top of the gas-liquid separator, and a condenser is arranged on the pipeline; the post-reaction material flow of the reactor is discharged from the top, and is sent to the gas-liquid separator through the condenser, a gas phase outlet is arranged at the top of the gas-liquid separator, and a liquid phase outlet is arranged at the bottom; The gas phase outlet is connected to a hydrogen inlet pipeline, and an outlet for transferring hydrogen is arranged on the pipeline; The liquid phase outlet is connected to the raw material inlet pipeline, and a product outlet is arranged on the pipeline; Preferably, a gas distributor is arranged at the bottom of the reactor; the inlet hydrogen enters the reactor through the gas distributor; and / or, The reactor is a bubbling bed reactor, and the catalyst is fixedly loaded in the reaction tube of the reactor.

13. Isopentenol prepared by the method according to any one of claims 7 to 11.

Images