CN102649562B - Method for dehydrogenation of CO gas raw material in virtue of catalytic oxidation - Google Patents
Method for dehydrogenation of CO gas raw material in virtue of catalytic oxidation Download PDFInfo
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Abstract
The invention relates to a method for dehydrogenation of CO gas raw material in virtue of catalytic oxidation and mainly solves the technical problems of difficulty in temperature control and low hydrogen desorbing rate in the reaction process of dehydrogenation of the CO gas raw material in virtue of catalytic oxidation. In the invention, by taking gas containing hydrogen and CO as raw material, and under the conditions that the reaction temperature is 80 to 260 DEG C, the volume space velocity is 100 to 10,000 hours<-1>, the molar ratio of oxygen/ hydrogen is (0.5-10) : 1, and the reaction pressure is 0.008 below zero to 5.0 MPa, the invention adopts the technical scheme that the raw material is sequentially in contact with the precious metal catalyst in an upper insulating catalyst layer, an isothermal catalyst bed and a lower insulating catalyst layer inside a combined reactor, and the hydrogen in the raw material is oxidized into water. Therefore, the problems are well solved, and the method can be used in industrial production of dehydrogenation reaction of the CO gas raw material in virtue of catalytic oxidation.
Description
Technical field
The present invention relates to a kind of CO gas raw material carries out dehydrogenation method by catalyzed oxidation, particularly, about adopting combination cot reactor, be practically applicable to CO gas raw material and undertaken in the reaction process of dehydrogenation by catalyzed oxidation.
Background technology
Barkite is important Organic Chemicals, in a large number produces various dyestuffs, medicine, important solvent, extraction agent and various intermediate for fine chemistry industry.Enter 21 century, barkite is subject to international extensively attention as degradable environment-friendly engineering plastics monomer.In addition, barkite ordinary-pressure hydrolysis can obtain oxalic acid, and normal pressure ammonia solution can obtain high-quality slow chemical fertilizer oxamyl.Barkite can also be used as solvent, produces medicine and dyestuff intermediate etc., for example, carry out various condensation reactions with fatty acid ester, hexamethylene phenyl methyl ketone, amido alcohol and many heterogeneous ring compounds.It can also synthesize at the chest acyl alkali that is pharmaceutically used as hormone.In addition, barkite low-voltage hydrogenation can be prepared very important industrial chemicals ethylene glycol, and ethylene glycol mainly relies on petroleum path to prepare at present, and cost is higher, and China needs a large amount of import ethylene glycol every year, 2007 years nearly 4,800,000 tons of import volumes.
The production line of tradition barkite utilizes oxalic acid to prepare with alcohol generation esterification, and production technique cost is high, and energy consumption is large, seriously polluted, and prepared using is unreasonable.Become the focus of domestic and international research and adopt carbon monoxide coupling technology to produce barkite.
As everyone knows, carbon monoxide can be from various containing separation and Extraction the gas mixture of carbon monoxide, and the industrial unstripped gas that can be used for separating carbon monoxide comprises: the tail gas of synthetic gas, water-gas, semi-water gas and Steel Plant, calcium carbide factory and Yellow Phosphorous Plant that Sweet natural gas and oil transform etc.The main method of existing CO separating-purifying is pressure swing adsorption process, You Duo company of China has developed pressure-variable adsorption and has separated carbon monoxide new technology, especially the high-efficiency adsorbent of exploitation, carbon monoxide is had to high loading capacity and selectivity, can solve a difficult problem of isolating high-purity carbon monooxide from nitrogen or the high unstripped gas of methane content, can design and build up large-scale carbon monoxide tripping device.However, by this technology isolated carbon monoxide from synthetic gas, taking into account under the prerequisite of carbon monoxide yield, the content of its hydrogen can reach more than 1% under normal circumstances.And research shows that the existence of hydrogen can cause follow-up CO coupling reaction catalyst activity decreased, until reaction cannot be carried out, therefore, exploitation carbon monoxide selects dehydrogenation technical meaning great.
At present, the dehydrogenation catalyst of report mainly contains Pd/Al both at home and abroad
2o
3, carbon monoxide Mo/Al
2o
3deng, also there is the dehydrogenating agent based on manganese series metal oxide, but being generally used for the dehydrogenation of the non-reducing gas such as High Purity Nitrogen, high purity oxygen and carbonic acid gas, these catalyzer or dehydrogenating agent purify.And under existing for CO reducing gas, catalyzer is low to the decreasing ratio of hydrogen, the rate of loss of CO is high.As adopt method and the catalyzer of the disclosed catalytic oxidative dehydrogenation of document CN97191805.8, and be raw material at the CO mixed gas for hydrogen content 10%, 220 DEG C of temperature of reaction, volume space velocity 3000 hours
-1, oxygen/hydrogen mol ratio is 0.6: 1, and under the condition that reaction pressure is 0.5MPa, the rate of loss of CO is up to 1.5%, and in reaction effluent, the content of hydrogen is up to 1000ppm.
For the strong exothermal reaction of CO oxydehydrogenation, in order to improve reaction efficiency, need to shift out reaction heat in reaction simultaneously, conventionally adopt tubular heat exchange, tube side passes into water or steam is removed the reaction heat in reaction tubes, but in shell and tube-type reactor design, in the full bed heat exchange of catalyzer, coolant temperature is constant.In the time that moving heat, need strengthening is conventionally subject to increasing the restriction of heat transfer area structure.Therefore adopt and improve charging tolerance gentle body linear velocity, prevent " overtemperature " and " temperature runaway " outward in time reaction heat is taken out of to reactor, but high linear gas velocity can bring the residence time too short conventionally, and then cause reacting the problem such as incomplete.
Task of the present invention is the shortcoming that overcomes above-mentioned prior art, and the CO realizing under wide in range condition oxidative dehydrogenation is provided.
Summary of the invention
Technical problem to be solved by this invention is to carry out dehydrogenation reaction process for CO gas raw material by catalyzed oxidation in previous literature technology, temperature control difficulty, exist hydrogen decreasing ratio low, the technical problem that CO rate of loss is high, provides a kind of new CO gas raw material to carry out the method for dehydrogenation by catalyzed oxidation.The method is carried out dehydrogenation for CO gas raw material process by catalyzed oxidation, temperature control is even, has hydrogen decreasing ratio high, the advantage that CO rate of loss is low.
In order to solve the problems of the technologies described above, the technical solution used in the present invention is as follows: a kind of CO gas raw material carries out the method for dehydrogenation by catalyzed oxidation, taking the gas of hydrogen and CO as raw material, is 80～260 DEG C in temperature of reaction, and volume space velocity is 100～10000 hours
-1, oxygen/hydrogen mol ratio is 0.5～10: 1, reaction pressure is under the condition of-0.08～5.0MPa, raw material successively with combined reactor in upper adiabatic catalyst layer, noble metal catalyst contact in Isothermal Catalyst bed and lower adiabatic catalyst layer, hydrogen in raw material is oxidized to water, wherein combined reactor is substantially by feed(raw material)inlet (1), feed(raw material)inlet (2), a distributing chamber of gas (26), a distributing chamber of gas (27), lower adiabatic catalyst layer (30), bundle of reaction tubes outer tube (5), bundle of reaction tubes inner tube (28), Isothermal Catalyst bed (7), upper adiabatic catalyst layer (31), gas quadratic distribution chamber (24), collection chamber (13), porous gas collection plate (11) and product outlet (12) composition, catalyst bed (7) is divided into the first heat exchange block (22) according to the mobile direction order of reaction gas, the second heat exchange block (19) and the 3rd heat exchange block (16), on the top of reactor upper tubesheet (4), upper adiabatic catalyst layer (31) is set, in the bottom of reactor lower tubesheet (10), lower adiabatic catalyst layer (30) is set.
Bundle of reaction tubes inner tube (28) is set in isothermal catalyst bed (7) in technique scheme, and bundle of reaction tubes inner tube (28) is connected with a distributing chamber of gas (27) with the distributing chamber of gas (26) in collection chamber (13) by inlet gas connecting hose (29).Porous gas collection plate (11) is positioned at collection chamber (13), and is connected with product outlet (12).The first heat exchange block (22) is connected with first district's heat transferring medium entrance (21) with first district's heat transferring medium outlet (23), the second heat exchange block (19) is connected with Second Region heat transferring medium outlet (20) with Second Region heat transferring medium entrance (8), is connected with the 3rd heat exchange block (16) Yu tri-district's heat transferring medium entrance (15) He tri-district's heat transferring medium outlets (17); Between the first heat exchange block (22) and the second heat exchange block (19), separate by the first subregion dividing plate (6), between the second heat exchange block (19) and the 3rd heat exchange block (16), separate by the second subregion dividing plate (9).The first subregion dividing plate (6) is preferably 1/8～1/3 of reactor length apart from the lower distance of reactor cover plate (25); The lower distance of second subregion dividing plate (9) distance the first subregion dividing plate (6) is preferably 1/8～1/3 of reactor length.The height of lower adiabatic catalyst layer (30) preferably Isothermal Catalyst bed (7) height 1/20～1/3; The height of upper adiabatic catalyst layer (31) preferably Isothermal Catalyst bed (7) height 1/20～1/3.
In technique scheme, reaction conditions is preferably: temperature of reaction is 120～240 DEG C, and volume space velocity is 500～6000 hours
-1, oxygen/hydrogen mol ratio is 0.5～4: 1, and reaction pressure is 0.01～2.0MPa, and noble metal catalyst active ingredient is preferably selected from palladium or platinum, and carrier is preferably selected from aluminum oxide.
Because catalyzed reaction is carried out on catalyzer and not according to front and back phase uniform velocity, general reactor front portion is from balanced remote, speed of response is fast, emit reaction heat also many, rear portion approaches balance with reaction, speed of response slows down, emit reaction heat also few, if the same before and after the temperature of refrigerant, if reduce like this coolant temperature, strengthen heat transfer temperature difference and move heat, reach the heat request that moves of top or anterior high speed of response and strong reaction heat, reactor lower part or rear portion reaction heat reduce, move heat be greater than reaction heat cause temperature of reaction decline, speed of response is further slowed down until catalyst activity is following with regard to stopped reaction, therefore be difficult to accomplish that front and rear part reacts the way making the best of both worlds of all carrying out under optimal reaction temperature.The present invention is directed to this fundamental contradiction, break through the existing refrigerant with same temperature, and adopt the different sections of reactor to adopt differing temps refrigerant to solve, make the size that in reaction, heat exchange is shifted out by reaction heat need design, multiple districts before and after specifically can being divided into by reaction gas flow direction order in catalyst layer, carry out indirect heat exchange by refrigerant by heat transfer tube.On the other hand, the present invention is for the reaction heat of catalyzer, also adopt inner tube is set in catalyst bed, and counter-current flow unstripped gas, unstripped gas is carried out to preheating has saved energy consumption on the one hand, optimized reaction bed temperature distribution simultaneously, thereby realize the equiblibrium mass distribution of full bed temperature, in addition, the present invention, at entrance and the outlet section of reactor, has all adopted adiabatic reactor, this is for the efficiency of maximized performance catalyzer, farthest reduce the loss of CO, and remove comparatively up hill and dale the hydrogen in raw material, useful effect is provided.
The present invention's Fig. 1 shown device, adopt subregion heat exchange, accurately control temperature, adopt the sleeve structure of inner and outer tubes catalyzer to be carried out to the catalyst loading method of heat exchange and multiple-hearth structure simultaneously, adopting precious metal palladium or platinum Supported alumina is catalyzer, be 80～260 DEG C in temperature of reaction, volume space velocity is 100～10000 hours
-1oxygen/hydrogen mol ratio is 0.5～10: 1, reaction pressure is under the condition of-0.08～5.0MPa, raw material contacts with the noble metal catalyst in upper adiabatic catalyst layer, Isothermal Catalyst bed and lower adiabatic catalyst layer in combined reactor successively, hydrogen in raw material is oxidized to water, containing in the gas raw material of CO, the volumn concentration of hydrogen is to be greater than under 0～15% condition, the decreasing ratio of hydrogen can reach 100%, the rate of loss of CO can be less than 0.5%, has obtained good technique effect.
Brief description of the drawings
Fig. 1 is the combined reactor schematic diagram that the present invention adopts.
In Fig. 1,1 and 2 is feed(raw material)inlets, the 3rd, reactor upper cover, the 4th, upper tubesheet, the 5th, bundle of reaction tubes outer tube, 6 is first subregion dividing plates, the 7th, catalyst bed, the 8th, reactor tank body, 9 is second subregion dividing plates, the 10th, lower tubesheet, the 11st, porous gas collection plate, the 12nd, product outlet, the 13rd, collection chamber, the 14th, reactor lower cover, 15 Shi tri-district's heat transferring medium entrances, 16 is the 3rd heat exchange blocks, 17 Shi 3rd district heat transferring medium outlets, the 18th, Second Region heat transferring medium entrance, 19 is second heat exchange blocks, the 20th, the outlet of Second Region heat transferring medium, 21 is first district's heat transferring medium entrances, 22 is first heat exchange blocks, 23 is the heat transferring medium outlets of the firstth district, the 24th, gas quadratic distribution chamber, the 25th, reactor cover plate, 26 and 27 is distributing chamber of gas, the 28th, bundle of reaction tubes inner tube, the 29th, inlet gas connecting hose, the 30th, lower adiabatic catalyst layer, the 31st, upper adiabatic catalyst layer.
Fig. 1 Raw is introduced by feed(raw material)inlet 1 and 2, respectively through distributing chamber of gas 26 and 27, introduce bundle of reaction tubes inner tube 28 through inlet gas connecting hose 29, with after reaction heat heat exchange in Isothermal Catalyst bed 7, enter in gas quadratic distribution chamber 24, be introduced into afterwards adiabatic catalyst layer 31 and carry out initial reaction, reacted product enters in the Isothermal Catalyst bed 7 between bundle of reaction tubes outer tube 5 and bundle of reaction tubes inner tube 28 again, with catalyzer contact reacts, reacted product enters lower adiabatic catalyst layer 31 again and continues reaction, finally enter after collection chamber 13, export 12 through porous gas collection plate 11 by product and enter follow-up system.In reaction raw materials gas enters the Isothermal Catalyst bed 7 between bundle of reaction tubes outer tube 5 and bundle of reaction tubes inner tube 28, with the reaction heat in catalyzer contact reacts process, successively through the first heat exchange block (22), the second heat exchange block (19) and the 3rd heat exchange block (16), the temperature of each heat exchange block can be by entering temperature and the control of flow philosophy of heat transferring medium of each heat exchange block, in addition, unstripped gas is from bundle of reaction tubes inner tube 28 and reactant gases counter current contact process, also catalyst bed 7 heat balances are played to better promoter action, thereby reach the uniform effect of whole reactor catalyst bed tempertaure.
Below by embodiment, the present invention is further elaborated.
Embodiment
[embodiment 1]
With reaction unit shown in Fig. 1, first, second and third heat transferring medium all adopts saturation steam, just by the difference of pressure, realize the difference of temperature, thereby the control of realization response device catalyst bed temperature, adopt in addition the sleeve structure of inner and outer tubes to carry out heat exchange to catalyzer, the first subregion dividing plate of reactor is reactor length apart from the length under reactor cover plate 1/5; Length under second subregion dividing plate distance the first subregion dividing plate is reactor length 1/6.The height of the lower adiabatic catalyst layer of reactor is 1/6 of Isothermal Catalyst bed height; The height of upper adiabatic catalyst layer is 1/10 of Isothermal Catalyst bed height, the catalyzer of the palladium Supported alumina taking palladium content as 0.5% is as catalyzer, with the CO mixed gas of hydrogen content 10% be raw material, at 180 DEG C of temperature ins of reaction, volume space velocity 3000 hours
-1, oxygen/hydrogen mol ratio is 0.6: 1, and under the condition that reaction pressure is 2.5MPa, reaction result is: the rate of loss of CO is 0.30%, and in reaction effluent, the content of hydrogen is 0ppm, and reactor catalyst bed temperature is poor is less than 6 DEG C.
[embodiment 2]
With the reaction unit of Fig. 1, first, second and third heat transferring medium all adopts saturation steam, just by adopting the difference of pressure, realize the difference of temperature, thereby the control of realization response device catalyst bed temperature, adopt in addition the sleeve structure of inner and outer tubes to carry out heat exchange to catalyzer, the first subregion dividing plate of reactor is reactor length apart from the length under reactor cover plate 1/8; Length under second subregion dividing plate distance the first subregion dividing plate is reactor length 1/5.The height of the lower adiabatic catalyst layer of reactor is 1/15 of Isothermal Catalyst bed height; The height of upper adiabatic catalyst layer is 1/15 of Isothermal Catalyst bed height, and the catalyzer of the palladium Supported alumina taking palladium content as 0.2% is as catalyzer, with the CO mixed gas of hydrogen content 5% be raw material, at 200 DEG C of reaction temperature ins, volume space velocity 1000 hours
-1, oxygen/hydrogen mol ratio is 0.7: 1, and under the condition that reaction pressure is 0.1MPa, reaction result is: the rate of loss of CO is 0.18%, and in reaction effluent, the content of hydrogen is 2ppm, and reactor catalyst bed temperature is poor is less than 8 DEG C.
[embodiment 3]
With the reaction unit of Fig. 1, first, second and third heat transferring medium all adopts saturation steam, just by adopting the difference of pressure, realize the difference of temperature, thereby the control of realization response device catalyst bed temperature, adopt in addition the sleeve structure of inner and outer tubes to carry out heat exchange to catalyzer, the first subregion dividing plate of reactor is reactor length apart from the length under reactor cover plate 1/4; Length under second subregion dividing plate distance the first subregion dividing plate is reactor length 1/3.The height of the lower adiabatic catalyst layer of reactor is 1/8 of Isothermal Catalyst bed height; The height of upper adiabatic catalyst layer is 1/6 of Isothermal Catalyst bed height, and the catalyzer of the palladium Supported alumina taking palladium content as 0.2% is as catalyzer, with the CO mixed gas of hydrogen content 3% be raw material, at 240 DEG C of reaction temperature ins, volume space velocity 6000 hours
-1, oxygen/hydrogen mol ratio is 0.8: 1, and under the condition that reaction pressure is 2.0MPa, reaction result is: the rate of loss of CO is 0.15%, and in reaction effluent, the content of hydrogen is 3ppm, and reactor catalyst bed temperature is poor is less than 7 DEG C.
[embodiment 4]
With the reaction unit of Fig. 1, first, second and third heat transferring medium all adopts saturation steam, just by adopting the difference of pressure, realize the difference of temperature, thereby the control of realization response device catalyst bed temperature, adopt in addition the sleeve structure of inner and outer tubes to carry out heat exchange to catalyzer, the first subregion dividing plate of reactor is reactor length apart from the length under reactor cover plate 1/4; Length under second subregion dividing plate distance the first subregion dividing plate is reactor length 1/8.The height of the lower adiabatic catalyst layer of reactor is 1/18 of Isothermal Catalyst bed height; The height of upper adiabatic catalyst layer is 1/6 of Isothermal Catalyst bed height, and the catalyzer of the platinum Supported alumina taking platinum content as 0.1% is as catalyzer, with the CO mixed gas of hydrogen content 1% be raw material, at 260 DEG C of reaction temperature ins, volume space velocity 500 hours
-1, oxygen/hydrogen mol ratio is 0.6: 1, and under the condition that reaction pressure is 0.1MPa, reaction result is: the rate of loss of CO is 0.18%, and in reaction effluent, the content of hydrogen is 10ppm, and reactor catalyst bed temperature is poor is less than 10 DEG C.
[comparative example 1]
With reference to each step and the reaction conditions of embodiment 1, just reactor adopts insulation fix bed reactor, and reaction result is: the rate of loss of CO is 3.2%, and in reaction effluent, the content of hydrogen is 160ppm, and reactor catalyst bed temperature is poor is 20 DEG C.
[comparative example 2]
With reference to each step and the reaction conditions of embodiment 2, just reactor adopts insulation fix bed reactor, and reaction result is: the rate of loss of CO is 4.2%, and in reaction effluent, the content of hydrogen is 180ppm, and reactor catalyst bed temperature is poor is 30 DEG C.
Claims (5)
1. CO gas raw material carries out a method for dehydrogenation by catalyzed oxidation, taking the gas of hydrogen and CO as raw material, is 80～260 DEG C in temperature of reaction, and volume space velocity is 100～10000 hours
-1, oxygen/hydrogen mol ratio is 0.5～10: 1, reaction pressure is under the condition of-0.08～5.0MPa, raw material successively with combined reactor in upper adiabatic catalyst layer, noble metal catalyst contact in Isothermal Catalyst bed and lower adiabatic catalyst layer, hydrogen in raw material is oxidized to water, wherein combined reactor is substantially by feed(raw material)inlet (1), feed(raw material)inlet (2), a distributing chamber of gas (26), a distributing chamber of gas (27), lower adiabatic catalyst layer (30), bundle of reaction tubes outer tube (5), bundle of reaction tubes inner tube (28), Isothermal Catalyst bed (7), upper adiabatic catalyst layer (31), gas quadratic distribution chamber (24), collection chamber (13), porous gas collection plate (11) and product outlet (12) composition, catalyst bed (7) is divided into the first heat exchange block (22) according to the mobile direction order of reaction gas, the second heat exchange block (19) and the 3rd heat exchange block (16), on the top of reactor upper tubesheet (4), upper adiabatic catalyst layer (31) is set, in the bottom of reactor lower tubesheet (10), lower adiabatic catalyst layer (30) is set, in the Isothermal Catalyst bed (7) of described reactor, bundle of reaction tubes inner tube (28) is set, bundle of reaction tubes inner tube (28) is connected with a distributing chamber of gas (27) with the distributing chamber of gas (26) in collection chamber (13) by inlet gas connecting hose (29), the porous gas collection plate (11) of described reactor is positioned at collection chamber (13), and is connected with product outlet (12).
2. the method that CO gas raw material carries out dehydrogenation by catalyzed oxidation according to claim 1, the the first heat exchange block (22) that it is characterized in that reactor is connected with first district's heat transferring medium entrance (21) with first district's heat transferring medium outlet (23), the second heat exchange block (19) is connected with Second Region heat transferring medium outlet (20) with Second Region heat transferring medium entrance (8), is connected with the 3rd heat exchange block (16) Yu tri-district's heat transferring medium entrance (15) He tri-district's heat transferring medium outlets (17); Between the first heat exchange block (22) and the second heat exchange block (19), separate by the first subregion dividing plate (6), between the second heat exchange block (19) and the 3rd heat exchange block (16), separate by the second subregion dividing plate (9).
3. the method that CO gas raw material carries out dehydrogenation by catalyzed oxidation according to claim 1, the first subregion dividing plate (6) that it is characterized in that reactor is reactor length apart from the length under reactor cover plate (25) 1/8～1/3; Length under second subregion dividing plate (9) distance the first subregion dividing plate (6) is reactor length 1/8～1/3.
4. the method that CO gas raw material carries out dehydrogenation by catalyzed oxidation according to claim 1, the height that it is characterized in that the lower adiabatic catalyst layer (30) of reactor is 1/20～1/3 of Isothermal Catalyst bed (7) height; The height of upper adiabatic catalyst layer (31) is 1/20～1/3 of Isothermal Catalyst bed (7) height.
5. the method that CO gas raw material carries out dehydrogenation by catalyzed oxidation according to claim 1, is characterized in that temperature of reaction is 120～240 DEG C, and volume space velocity is 500～6000 hours
-1, oxygen/hydrogen mol ratio is 0.5～4: 1, and reaction pressure is 0.01～2.0MPa, and noble metal catalyst active ingredient is selected from palladium or platinum, and carrier is selected from aluminum oxide.
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| CN114956018B (en) * | 2022-06-30 | 2024-06-28 | 重庆川茂化工科技有限公司 | Constant-temperature dehydrogenation equipment and method for high-hydrogen helium |
| CN115463618B (en) * | 2022-08-08 | 2023-11-10 | 北京鑫缘化工有限公司 | Reactor for preparing maleic anhydride by oxidizing n-butane |
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| CN2621805Y (en) * | 2003-06-03 | 2004-06-30 | 华东理工大学 | Shell external cooling-thermal insulating and combined fixed bed catalyst chember |
| CN1988950A (en) * | 2004-05-28 | 2007-06-27 | 住友化学株式会社 | Heat exchange type reactor |
| US20070264186A1 (en) * | 2004-09-09 | 2007-11-15 | Dybkjaer Ib | Process for Production of Hydrogen and/or Carbon Monoxide |
| CN101543776A (en) * | 2009-04-28 | 2009-09-30 | 华烁科技股份有限公司 | Dehydrogenation catalyst for feed gas containing carbon monoxide, preparation method and application method thereof |
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| Publication number | Priority date | Publication date | Assignee | Title |
|---|---|---|---|---|
| CN2621805Y (en) * | 2003-06-03 | 2004-06-30 | 华东理工大学 | Shell external cooling-thermal insulating and combined fixed bed catalyst chember |
| CN1988950A (en) * | 2004-05-28 | 2007-06-27 | 住友化学株式会社 | Heat exchange type reactor |
| US20070264186A1 (en) * | 2004-09-09 | 2007-11-15 | Dybkjaer Ib | Process for Production of Hydrogen and/or Carbon Monoxide |
| CN101543776A (en) * | 2009-04-28 | 2009-09-30 | 华烁科技股份有限公司 | Dehydrogenation catalyst for feed gas containing carbon monoxide, preparation method and application method thereof |
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