JP4740564B2 - Hydrogen purification method - Google Patents

Description

  The present invention relates to a method for purifying hydrogen by separating and removing an impurity gas other than hydrogen from a hydrogen-containing mixed gas containing hydrogen as a main component to obtain higher-purity hydrogen, and particularly uses a hydrogenation reaction and a dehydrogenation reaction. The present invention also relates to a novel method for purifying hydrogen that separates and removes impurity gases other than hydrogen from a hydrogen-containing mixed gas.

  Hydrogen with clean secondary energy is expected to be used in the future, and the development of hydrogen utilization technologies such as fuel cells is spurring. Industrial production of this hydrogen includes water electrolysis, coke gasification, coal complete gasification, petroleum gasification, natural gas modification, coke oven gas modification, petroleum refining waste gas modification, iron Many methods are known such as hydrogen recovery from off-gas generated in steelworks, refineries, ethylene plants, etc., including the reaction of water and steam, decomposition of methanol and ammonia, and brine electrolysis. In addition, hydrogen gas produced by these methods often contains a relatively large amount of impurity gas other than hydrogen, such as methane, carbon dioxide gas, nitrogen gas, carbon monoxide, etc. In this way, impurity gases other than hydrogen are separated and removed from a hydrogen-containing mixed gas containing a relatively large amount of impurity gas, and used as higher-purity hydrogen.

  And, as a purification method of hydrogen to obtain higher-purity hydrogen by separating and removing impurity gases other than hydrogen from this hydrogen-containing mixed gas, generally, PSA (pressure swing adsorption), TSA ( Methods such as temperature swing adsorption), cryogenic separation, membrane separation, etc. are employed. In particular, in order to obtain high-purity hydrogen in which 90 mol% exceeds 99.9 mol% in some cases, zeolite, activated carbon, silica gel, PSA using an adsorbent such as carbon molecular sieve is widely used as the most effective method (for example, published by Toshicho Kawai, “Pressure Swing Adsorption Technology Assembly” Industrial Technology Association in 1986).

  However, this PSA requires a pressure / vacuum resistant device, and it is difficult to desorb strongly adsorbed impurities, and because all the impurity gases other than hydrogen are adsorbed to the adsorbent, it is proportional to the increase in throughput. As a result, the amount of adsorbent used is increased, the size of the apparatus is increased, and the scale and resistance of the switching valve for switching between the adsorption area and the regeneration area of the adsorbent are limited. There is a problem that the product hydrogen recovery rate is generally 85 mol% or less and 80 mol% or less in actual operation because hydrogen is inevitably left in the adsorption region due to partial adsorption.

Published by Toshicho Kawai, “Pressure Swing Adsorption Technology Collection”, Industrial Technology Association, 1986

  Therefore, the present inventors have provided a method for purifying hydrogen, which obtains higher-purity hydrogen by separating and removing impurity gases other than hydrogen from a relatively low-grade hydrogen-containing mixed gas containing hydrogen as a main component. High-purity hydrogen that can be used in industrial applications such as chemical processes and fuel cell applications that require ultra high purity of 99.99 mol% or more can be easily obtained, and continuous processing is possible. However, as a result of intensive studies on a hydrogen purification method that can easily cope with large-scale treatment of hydrogen-containing mixed gas by scaling up, the hydrogen-containing mixed gas and the aromatic compound are brought into contact in the presence of a hydrogenation catalyst. A so-called reaction component is applied in combination with a hydrogenation reaction of an aromatic compound and a dehydrogenation reaction of a hydrogenated aromatic compound in which the hydrogenated aromatic compound obtained by the hydrogenation reaction is brought into contact with a dehydrogenation catalyst. The present inventors have found that higher purity hydrogen can be obtained by separating and removing impurity gases other than hydrogen from the hydrogen-containing mixed gas by the separation method.

  Therefore, the object of the present invention is to easily obtain high-purity hydrogen that is easy to operate, and is capable of continuous processing, and can easily cope with large-scale processing of hydrogen-containing mixed gas by scaling up. An object of the present invention is to provide a novel method for purifying hydrogen.

That is, the present invention is a method for purifying hydrogen that obtains higher-purity hydrogen by separating and removing impurity gases other than hydrogen from a hydrogen-containing mixed gas containing hydrogen as a main component, and in the presence of a hydrogenation catalyst, The mixed gas and the aromatic compound are brought into contact with each other to hydrogenate the aromatic compound, and the hydrogenated reaction mixture obtained by the hydrogenation reaction is gas-liquid separated to recover the hydrogenated aromatic compound, and then dehydrogenated. the hydrogenated aromatic compound in the presence of a catalyst allowed dehydrogenation, purification process der of hydrogen you recover hydrogen generated by dehydrogenation reaction,

As the dehydrogenation catalyst of the above SL, surface area 150 meters ² / g or more, a pore volume 0.55 cm ³ / g or more, an average pore diameter 90～300A, and pores having a pore diameter 90～300A is the total pore volume Particle size with at least one catalytic metal selected from platinum, palladium, rhodium, iridium and ruthenium being 10 % or less on a porous γ-alumina support having an occupation ratio (pore diameter 90 to 300%) of 60% or more The dehydrogenation catalyst supported by the above is used . As a result, in the dehydrogenation reaction of the hydrogenated aromatic compound, it is possible to achieve a conversion rate of 90% or more of the hydrogenated aromatic compound at a relatively low reaction temperature of 290 to 350 ° C., and a high reaction selectivity of 98% or more. And can be stably operated over a long period of time.

Furthermore, as the upper Symbol dehydrogenation catalyst, it is preferable to use a dehydrogenation catalyst, wherein the catalyst metal and the alkali metal is supported, thereby, in the dehydrogenation reaction of hydrogenated aromatic compounds, referred to two hundred and ninety to three hundred and fifty ° C. The conversion of the hydrogenated aromatic compound can be achieved at a relatively low reaction temperature of 90% or more, and it has a high reaction selectivity of 98% or more and can be stably operated over a long period of time.

  In the present invention, the hydrogen-containing mixed gas containing hydrogen as a main component is not particularly limited in its origin, and usually has a hydrogen content of 30 mol% to 90 mol%, preferably 50 mol% to 70 mol%. There should be. When the hydrogen content of this hydrogen-containing mixed gas is lower than 30 mol%, the hydrogenation reaction is a reaction that is regulated by equilibrium, so the equilibrium conversion rate of hydrogenation is lowered and the reaction conditions for obtaining a sufficient recovery rate On the other hand, when it exceeds 90 mol%, hydrogenation becomes easy in equilibrium, but the purity improvement is low, so that there is a problem that it is difficult to have economic efficiency.

  If this hydrogen-containing gas contains a catalyst poison gas that is a catalyst poison of a hydrogenation catalyst, such as carbon monoxide gas (CO) or a sulfur element-containing gas such as hydrogen sulfide or mercaptan, these are included in advance. The catalyst poison gas is preferably separated and removed as much as possible, preferably 10 ppm or less, more preferably 1 ppm or less. Such a catalyst poison gas separation and removal operation is not particularly limited, and depending on the type of the catalyst poison gas, for example, an adsorption operation using an adsorbent, a chemical absorption operation using an absorbing solution, or a reaction removal operation using zinc oxide or the like. This method can be adopted.

  In the present invention, the aromatic compound to be reacted with the hydrogen-containing mixed gas in the presence of a hydrogenation catalyst is not only stable but also hydrogenated to become a stable hydrogenated aromatic compound. Although not particularly limited, it is preferably a monocyclic aromatic compound such as benzene, toluene or xylene, a bicyclic aromatic compound such as naphthalene, tetralin or methylnaphthalene, or a tricyclic aromatic compound such as anthracene. More preferred are monocyclic aromatic compounds such as benzene, toluene and xylene, and bicyclic aromatic compounds such as naphthalene, tetralin and methylnaphthalene, and more preferred are benzene and toluene. , Naphthalene and methylnaphthalene, which can be used alone or as a mixture of two or more It is also possible.

  And as a hydrogenation catalyst used for the hydrogenation reaction of this aromatic compound, those conventionally used for this type of hydrogenation reaction can be used as they are, for the production of conventional hydrogenated aromatics. Depending on the type of aromatic compound and the specifications of the hydrogenated aromatic product, for example, alumina or silica is used as a carrier, and platinum (Pt), palladium (Pd), nickel (Ni), etc. are supported as active metals. The prepared catalyst is preferably used.

  Also, the reaction conditions for the hydrogenation reaction are the same as the conventional hydrogenation reaction of this type. For example, when benzene is used as the aromatic compound, the reaction temperature is usually 200 to 400 ° C. and the reaction pressure is 25 to 50 atm. Done on condition. Further, the reaction method may be a gas phase fixed bed method or a liquid phase slurry bed method using Raney nickel as a hydrogenation catalyst. Furthermore, the reaction is divided into two stages and the reaction is performed at a relatively high temperature in the previous stage. It is also possible to adopt a method in which the reaction proceeds to the vicinity of the equilibrium conversion rate by lowering the temperature in the latter stage while proceeding quickly.

  In the present invention, the hydrogenation reaction mixture obtained by the hydrogenation reaction is then gas-liquid separated to recover the hydrogenated aromatic compound. The hydrogenated aromatic compound recovered here is the product of the hydrogenation reaction of the aromatic compound, and, like this aromatic compound, is itself stable and dehydrogenated and stable aromatic compound. As well as monocyclic hydrogenated aromatic compounds such as cyclohexane, methylcyclohexane and dimethylcyclohexane corresponding to the above aromatic compounds, and bicyclic hydrogenated aromatic compounds such as tetralin, decalin and methyldecalin. Compounds, and tricyclic hydrogenated aromatic compounds such as tetradecahydroanthracene.

The hydrogenated aromatic compound thus obtained is then dehydrogenated in the presence of a dehydrogenation catalyst, and the hydrogen generated during the dehydrogenation reaction is recovered as purified hydrogen.
As the dehydrogenation catalyst used for this purpose, those conventionally used in this type of dehydrogenation reaction can be used as they are, but the dehydrogenation reaction of this hydrogenated aromatic compound is carried out at 290 to 350 ° C. From the viewpoint of extremely efficient operation at a relatively low reaction temperature of 90% or more and a reaction selectivity of 98% or more and stable operation over a long period of time, the following dehydrogenation catalyst is preferable. Should be used. That is, the hydrogenation aromatic compound dehydrogenation catalyst has a surface area of 150 m ² / g or more, a pore volume of 0.55 cm ³ / g or more, an average pore diameter of 90 to 300 mm, and a pore diameter of 90 to the total pore volume. At least one catalyst metal selected from platinum, palladium, rhodium, iridium and ruthenium on a porous γ-alumina support in which the proportion of pores of ~ 300Å (pore size 90 to 30090) is 60% or more It is preferable to use a dehydrogenation catalyst A on which is supported, or a surface area of 150 m ² / g or more, a pore volume of 0.55 cm ³ / g or more, an average pore diameter of 90 to 300 mm, and fine with respect to the total pore volume. At least one selected from platinum, palladium, rhodium, iridium and ruthenium is used for the porous γ-alumina carrier in which the proportion of pores having a pore size of 90 to 300 mm (pore size 90 to 300 mm) is 60% or more. It is preferable to use a dehydrogenation catalyst B in which the catalyst metal and the alkali metal supported.

Hereinafter, dehydrogenation catalysts A and B suitable for use in the dehydrogenation reaction of the hydrogenated aromatic compound of the present invention will be described in detail.
First, the porous γ-alumina support used as the catalyst support has a surface area of 150 m ² / g or more, preferably 200 m ² / g or more, and a pore volume of 0.55 cm ³ / g or more, preferably 0.65 cm ^3. The average pore diameter is 90 to 300 mm, preferably 100 to 200 mm, and the pore diameter 90 to 300 mm occupancy is 60% or more, preferably 80% or more. When the surface area is less than 150 m ² / g, the activity after the catalyst is not sufficient, and when the pore volume is less than 0.55 cm ³ / g, it is difficult to uniformly carry the active metal component. As for the average pore diameter, when the average pore diameter is smaller than 90 mm, the surface area increases, but the pore volume decreases, and conversely, when the average pore diameter is larger than 300 mm, the surface area decreases and the pore volume increases. As a result of comprehensively considering the correlation, an average pore diameter of 90 to 300 mm is appropriate. Furthermore, if the pore diameter 90-300% occupancy is less than 60%, the effect on the catalyst performance is reduced. The reason for using such an alumina support having specific physical properties is that the distribution of pores is uniformly controlled, and the use of an alumina support in which the pore size is concentrated in the range of 90 to 300 mm throughout the support allows platinum to be used. This is because the dispersion state is improved by uniformly impregnating with alkaline metal.

  Such a porous γ-alumina support is obtained by, for example, filtering and washing a slurry of aluminum hydroxide produced by neutralization of an aluminum salt, as disclosed in Japanese Patent Publication No. 6-72,005. Is obtained by baking at 400 to 800 ° C. for about 1 to 6 hours. Preferably, the pH value of the alumina hydrogel is alternated between the alumina hydrogel dissolution pH region and the boehmite gel precipitation pH region. And a pH swing step of growing an alumina hydrogel crystal by adding an alumina hydrogel-forming substance upon pH change from at least one of the pH ranges to the other pH range. . The porous γ-alumina support obtained through this pH swing process is excellent in the uniformity of pore distribution, there is little variation in physical properties even in the alumina support pellets after molding, and the physical properties of each pellet are stable. It is excellent in that it is.

  The catalyst metal supported on the porous γ-alumina carrier is one or more metals selected from platinum, palladium, rhodium, iridium and ruthenium, preferably platinum, and the supported amount thereof. For example, when the catalyst metal is platinum, it is 0.3 wt% or more and 2.0 wt% or less, preferably 0.5 wt% or more and 1.0 wt% or less. If the supported amount of platinum is less than 0.3% by weight, there is a problem that the activity is low. On the other hand, if it is more than 2.0% by weight, the particle size of platinum becomes large, the selectivity is lowered and sintering is performed. The problem that it is easy to deteriorate arises.

  In addition, when impregnating and supporting noble metals such as platinum and palladium on an alumina carrier, the degree of dispersion of the noble metals after firing and supporting may vary depending on the pH of the impregnating aqueous solution, but the optimum pH range is 1.8 to The range is 3.0. When the pH value of the impregnating solution is lower than 1.8, the degree of dispersion of the noble metals after loading becomes low, and when the pH value is higher than 3.0, the degree of dispersion also decreases. This is presumed that the adsorption force of the metal compound molecules on the alumina support differs depending on the pH value at the time of impregnation, and has a great influence on sintering during particle growth. Further, the degree of dispersion of the noble metal tends to decrease due to the subsequent supporting of the alkali metal, but by adjusting the pH value during impregnation of the noble metal to the range of 1.8 to 3.0, the alkali metal is dispersed. A decrease in the degree of dispersion of the noble metal due to the loading of can be minimized. As a result, the degree of dispersion of the noble metal after the catalyst can be highly dispersed to 70% or more, more preferably 80% or more. As described above, the size of the noble metal particles having a high degree of dispersion is 10 mm or less, and when the degree of dispersion is 70%, the size is 7 mm or less.

  As described above, the dehydrogenation catalyst A obtained by supporting a specific catalytic metal on a specific porous γ-alumina carrier is itself dehydrogenated for the dehydrogenation reaction of the hydrogenated aromatic compound of the present invention. Although it is suitable as a catalyst, it is more preferable to use an alkali metal together with a specific catalyst metal to form the dehydrogenation catalyst B. The reason for supporting the alkali metal in the dehydrogenation catalyst B is to mask the acid sites on the alumina and suppress the decomposition reaction on the alumina surface. The residual acid point on alumina is considered to vary depending on the amount of noble metal supported, and the amount of alkali metal necessary for masking decreases as the amount of noble metal supported increases. Therefore, if a highly dispersed state of platinum is realized, a certain level of performance is ensured even when masking with an alkali metal is not performed. However, hydrogenated aromatic compounds should be recovered after dehydrogenation and used again as aromatic compounds for hydrogenation reactions, and it is necessary to reduce the loss due to decomposition as much as possible. Is more desirable.

  Further, the alkaline metal supported on the porous γ-alumina support specifically includes Group 1A and Group 1 of the periodic table including lithium, sodium, potassium, rubidium, cesium, beryllium, magnesium, calcium, strontium, and barium. Group 2A metal element, preferably potassium, and the supported amount is, for example, 0.001 wt% or more and 1.0 wt% or less, preferably 0.005 wt% when the alkaline metal is potassium. The content is 0.5% by weight or less. When the amount of the alkaline metal supported is less than 0.001% by weight, there is a problem that the effect is not substantially obtained. On the other hand, when the amount is more than 1.0% by weight, there is a problem that the activity is reduced due to excess. is there.

The dehydrogenation catalyst used in the present invention is obtained by impregnating the above porous γ-alumina carrier with the above catalyst metal solution, drying and calcining it to obtain a catalyst metal-supported calcined product, which is not reduced. It is manufactured by impregnating and drying an alkaline metal solution in the state, and then directly carrying out final hydrogen reduction without firing the obtained dried alkaline metal supported product.
When an alkali metal is not supported, it can be prepared by an ordinary method in which a reduction operation is performed after impregnating a catalyst metal such as platinum and drying followed by air calcination.

  Here, the catalyst metal compound solution impregnated in the porous γ-alumina support includes various complex compounds such as chloride, bromide, ammonium salt, carbonyl compound, amine, ammine complex and acetylacetonato complex of the catalyst metal. For example, when the catalyst metal is platinum, chloroplatinic acid, acetylacetonatoplatinum, platinum platinum ammonium salt, bromoplatinic acid, platinum dichloride, platinum tetrachloride hydrate, carbonyl platinum dichloride Examples thereof include platinum compounds such as dichloride and dinitrodiamine platinate. When the catalyst metal is supported, the porous γ-alumina carrier is impregnated with the solution of the above-mentioned catalyst metal compound, and is preferably 50 ° C. or more and 200 ° C. or less, preferably 0.5 hours or more and 48 hours or less. And then baking is preferably performed at a temperature of 350 ° C. to 600 ° C., 0.5 hour to 48 hours, more preferably 350 ° C. to 450 ° C. for 0.5 hour to 5 hours. It is good.

  In addition, as an alkaline metal compound used when an alkali metal is supported on a catalyst metal-supported fired product obtained by supporting a catalyst metal on a porous γ-alumina carrier, an alkali metal chloride, bromide, iodide, Examples thereof include nitrates, sulfates, acetates, propionates and the like, preferably those which are soluble in water and / or soluble in organic solvents such as acetone, such as potassium chloride, potassium bromide, potassium iodide. , Potassium nitrate, potassium sulfate, potassium acetate, potassium propionate, rubidium chloride, rubidium bromide, rubidium iodide, rubidium nitrate, rubidium sulfate, rubidium acetate, rubidium propionate, lithium chloride, lithium bromide, lithium iodide, lithium nitrate , Lithium sulfate, lithium acetate, lithium propionate, cesium chloride , Cesium bromide, cesium iodide, cesium nitrate, cesium sulfate, cesium acetate, cesium propionate, magnesium chloride, magnesium bromide, magnesium iodide, magnesium nitrate, magnesium sulfate, magnesium acetate, magnesium propionate, calcium chloride, odor Calcium iodide, calcium iodide, calcium nitrate, calcium sulfate, calcium acetate, calcium propionate and the like can be mentioned.

  Further, when the alkaline metal is supported on the catalyst metal-supported fired product, after impregnating the alkaline metal compound solution, the temperature is from room temperature to 200 ° C. and from 0.5 hour to 48 hours, preferably from 50 ° C. to 150 ° C. It is preferable to dry under the drying conditions of not more than 0 ° C. and not less than 0.5 hours and not more than 24 hours, more preferably not less than 80 ° C. and not more than 120 ° C. and not less than 0.5 hours and not more than 5 hours.

  The alkaline metal-supported dried product obtained by impregnating the catalyst metal-supported fired product with an alkaline metal and drying is directly subjected to final hydrogen reduction without firing, and the reduction conditions for this hydrogen reduction are as follows. In an atmosphere of hydrogen gas, it is preferably performed at 350 ° C. to 600 ° C. and 0.5 hours to 48 hours, preferably 350 ° C. to 550 ° C. and 3 hours to 24 hours. Firing prior to the hydrogen reduction of the alkaline metal-supported dry matter causes a problem that the catalytic performance of activity, selectivity and life is lowered, and it is sufficient that the temperature during the hydrogen reduction is less than 350 ° C. On the other hand, there is a problem that platinum is not reduced. On the other hand, when the temperature exceeds 600 ° C., platinum particles are sintered at the time of reduction, and the metal dispersibility is lowered.

According to experiments, the dehydrogenation catalysts A and B obtained as described above are, for example, hydrogenated aroma under the reaction conditions of a reaction temperature of 300 ° C., normal pressure, and LHSV of the raw material solution of 2.0 h ^−1. The conversion of the group compound is usually 90% or more, preferably 95% or more, and the reaction selectivity is usually 98% or more, preferably 99% or more, at a relatively low reaction temperature of 290 to 350 ° C. The conversion rate of the hydrogenated aromatic compound can be 90% or more, and has a high reaction selectivity of 98% or more, and the hydrogenation aromatic compound can be stably dehydrogenated over a long period of time. .

  In the present invention, since the hydrogenation reaction of the aromatic compound is an exothermic reaction and the dehydrogenation reaction of the hydrogenated aromatic compound is an endothermic reaction, it is preferable that hydrogen that is an exothermic reaction in one reactor. Incorporating the hydrogenation region of the hydrogenation reaction and the dehydrogenation region of the dehydrogenation reaction that is endothermic, and configured so that heat can be transferred between the hydrogenation region and the dehydrogenation region, the hydrogenation reaction The generated reaction heat is preferably used as the reaction heat for the dehydrogenation reaction.

  Specifically, for example, by configuring the reactor like a double tube structure and operating the hydrogenation reaction, which is an exothermic reaction, at a higher temperature than the dehydrogenation reaction, which is an endothermic reaction, the hydrogenation reaction and dehydrogenation are performed. Heat exchange with the reaction may be performed.

  In the present invention, the dehydrogenation reaction mixture obtained by the dehydrogenation reaction of the hydrogenated aromatic compound is gas-liquid separated, and the hydrogen gas produced by the dehydrogenation reaction is recovered. Hydrogen gas recovered by gas-liquid separation of this dehydrogenation reaction mixture (primary recovered hydrogen gas) usually has a hydrogen concentration of 90 mol% or more, and examples of impurity gases include methane gas and carbon dioxide gas. The mixed gas used as the original hydrogen source contains the gas contained as a component other than hydrogen, the hydrogenated aromatic compound used for the reaction separation, and / or the vapor of the aromatic compound.

  Since this primary recovered hydrogen gas has a hydrogen concentration increased to 90 mol% or more, it can be used as it is as a product hydrogen gas for applications such as naphtha hydrogen reforming in refineries. One of the means for further improving the purity of the recovered hydrogen is a method of performing gas-liquid separation operation by applying pressure. This is because the amount of aromatics or the like present in the vapor pressure is reduced by pressurization. In general, it is considered that the recovered hydrogen is often used under pressure, so the hydrogen purity is improved by gas-liquid separation by pressurizing the pressure, and the pressure is maintained while maintaining the pressure. Is also a good method from an economic point of view. In other words, the hydrogen purity of the component that has undergone the dehydrogenation reaction can be selected economically from both the pressure condition using the recovered hydrogen and the hydrogen purity that can be used. However, for use in applications such as fuel cells, a purity with a hydrogen concentration of 99.99 mol% or more is required. When such high-purity hydrogen is produced, it is effective to use not only gas-liquid separation under pressure but also an adsorption column or the like because only high pressure is required to produce a high-purity hydrogen. In addition to the adsorption operation, an absorption operation can be exemplified. Even when these operations are combined, the amount of impurities is already very small compared to the case where hydrogen purification is performed by PSA or the like instead of the present invention, so the time until saturation adsorption or saturation absorption is reached. It can be taken for a long time, and the amount of adsorbent to be used is small, and scale-up is easy.

  According to the present invention, a reaction separation method comprising a hydrogenation reaction of an aromatic compound and a dehydrogenation reaction of a hydrogenated aromatic compound is applied to a relatively low-grade hydrogen-containing mixed gas containing hydrogen as a main component. Since hydrogen is purified, impurity gases other than hydrogen can be separated and removed efficiently with relatively simple operations, and impurities in the recovered hydrogen gas can also be dehydrogenated from hydrogenated aromatic compounds. Because it is limited to the specific components accompanying the fuel, it is extremely easy to further purify and purify, and fuel that requires ultra high purity of 99.99 mol% or more including industrial uses such as various chemical processes. High-purity hydrogen that can also be used for battery applications and the like can be easily obtained. In addition, continuous processing is possible, and large-scale treatment of a hydrogen-containing mixed gas can be easily handled by scaling up.

Hereinafter, based on the flow sheet shown in the accompanying drawings, an embodiment of the present invention will be specifically described using the calculated equilibrium value of each operation.
In FIG. 1, a hydrogenation reactor 1 for performing a hydrogenation reaction contains a hydrogen-containing mixed gas adjusted to 67 mol% hydrogen (HYD) and 33 mol% methane (MET) and 60 mol% toluene (TOL) as an aromatic compound. In the hydrogenation reactor 1, a hydrogenation reaction of toluene is performed in the presence of a hydrogenation catalyst under the conditions of a temperature of 330 ° C. and a pressure of 75 kg / cm ² A. At this time, the equilibrium conversion rate of hydrogen is 95.4 mol%.

Next, the hydrogenation reaction mixture (stream B) extracted from the hydrogenation reactor 1 is introduced into the gas-liquid separator 2, and gas-liquid separation is performed under normal pressure to separate off-gas (stream C). A hydrogenated liquid component (stream D) mainly composed of methylcyclohexane (MCH) and unreacted toluene and containing 4 × 10 ⁻³ mol% hydrogen and 4 × 10 ⁻⁵ mol% methane as a dissolved component is recovered. .

  About this hydrogenated liquid component (flow D), it heats to 300 degreeC with the preheater 3, introduce | transduces into a dehydrogenation reactor 4 as a high temperature hydrogenated liquid component (stream E), the conditions of temperature 300 degreeC and normal pressure The dehydrogenation catalyst of the above-mentioned dehydrogenation reactor 4 was made to contact below and the dehydrogenation catalyst of the methylcyclohexane was contacted. The equilibrium conversion rate of methylcyclohexane at this time is 96.4 mol%, corresponding to 89.2 mol% of the hydrogenated hydrogen conversion rate, and the hydrogen recovery rate at the outlet of the dehydrogenation reactor 4 is It is 85.2 mol%.

The dehydrogenation reaction mixture (stream F) extracted from the dehydrogenation reactor 4 is introduced into the gas-liquid separator 5 and separated into impurities (stream H) by gas-liquid separation at 30 ° C. and atmospheric pressure. At the same time, primary recovered hydrogen gas (flow G) is obtained.
The obtained primary recovered hydrogen gas (stream G) has a hydrogen purity of 92.7 mol%, and contains 6.7 mol% of toluene for vapor pressure and 0.37 mol% of methane as impurities.

When the purity of the primary recovered hydrogen gas (flow G) thus obtained is further improved by gas-liquid separation by pressurization, 10 kg / cm ² A (absolute pressure 0.981 MPa) , 20 kg / cm ² A ( Absolute pressure 1.962 MPa) or 50 kg / cm ² A (absolute pressure 4.905 MPa) (kg / cm ² A indicates absolute pressure) will be used for the description. Primary recovered hydrogen gas (flow G) is introduced into the gas-liquid separator 7 through the 10-atmospheric pressure booster 6, and is introduced into the gas-liquid separator 9 through the 20-atmospheric pressure booster 8. It is introduced into the gas-liquid separator 11 through the machine 10, and is pressurized at 30 ° C. under the conditions of 10 kg / cm ² A, 20 kg / cm ² A, or 50 kg / cm ² A, respectively. Stream J, stream L, or stream N) is separated and purified hydrogen gas (stream I, stream K, or stream M) is obtained.

Purified hydrogen gas (stream I, stream K, or stream M) obtained by gas-liquid separation under pressure has a hydrogen purity of 98.9 mol% (stream I), 99.2 mol% (stream K), and 99.4 mol% (stream M) and the hydrogen recovery is 85.1 mol% in all cases.
Below, the temperature (° C.), pressure (kg / cm ² A), mass flow rate (kg / h), molar composition (mol%) and molar flow rate (kgmol / h) of each flow A to flow N in this embodiment Are summarized in Table 1.

The following example demonstrates that a reasonable reaction result can be obtained using an actual catalyst with respect to the equilibrium calculated value by the flow sheet in FIG.
(Hydrogenation reaction)
A commercially available Pt (0.5 wt%) / Al as a hydrogenation catalyst is placed in a stainless steel reaction tube with an inner diameter of 12.6 mmφ x 300 mm equipped with a 1 / 8-inch thermocouple protection tube at the center of the reaction tube cross section. ₂ O ₃ catalyst 10 cc is packed so that the center of the catalyst layer is located in the center of the length direction of the reaction tube, and further 1 mmφ spherical α-alumina beads 10 cc are packed as a preheating layer above the catalyst layer, A pressurized flow type hydrogenation reaction test apparatus (hydrogenation reactor) for the hydrogenation reaction was used.

The catalyst layer of this hydrogenation reactor was heated to 300 ° C. in a hydrogen stream at a pressure of 75 kg / cm ² A, and a hydrogen-containing mixed gas having a composition of 67 mol% of hydrogen and 33 mol% of methane and 7,000 Ncc / h of toluene and 20 Ncc of toluene. / h (LHSV = 2.0 h ⁻¹ ) was supplied to the catalyst layer while maintaining a molar ratio of toluene and mixed gas (toluene / mixed gas) of 60: 100. At this time, the supply of the hydrogen flow for temperature increase was stopped at the same time. Although the temperature of the catalyst layer increased due to heat generation due to the progress of the hydrogenation reaction, the temperature of the catalyst layer was raised to 330 ° C. when the temperature increase was not observed, and the reaction was started.

  After 1 hour, the reaction tube outlet fraction was separated with a gas-liquid separator, the hydrogenated liquid component was sampled, and toluene and methylcyclohexane were quantified by gas chromatography. The conversion of toluene to methylcyclohexane was 34 mol%. In terms of the hydrogen conversion rate, the hydrogen conversion rate in the hydrogen-containing mixed gas was 92 mol%. The composition of this hydrogenated liquid component was 65.01 mol% toluene and 34.59 mol% methylcyclohexane.

(Preparation of porous γ-alumina support)
A γ-alumina support was produced as described in Example 1 of JP-B-6-72,005. The outline of this method is as follows. Aqueous sodium aluminate aqueous solution is added instantaneously with vigorous stirring in hot dilute sulfuric acid to obtain an aluminum hydroxide slurry suspension (pH 10), which is used as seed aluminum hydroxide and stirred. The operation of adding hot dilute sulfuric acid and sodium aluminate aqueous solution alternately for a certain time while repeating the above is repeated to obtain a filter washed cake, which is extruded and dried, and then baked at 500 ° C. for 3 hours. .

The properties of the porous γ-alumina support obtained by such pH swing operation (pH swing method) are as follows: the average pore diameter is 119 mm, the pore volume is 0.713 cm ³ / g, the surface area is 240 m ² / g, And the ratio of the pores of 90 to 300 mm in the total pore volume was 90 vol%.

(Preparation of dehydrogenation catalyst)
To 20 g of the porous γ-alumina carrier prepared as described above, 79 g of 0.4 wt% -chloroplatinic acid aqueous solution prepared so as to have a pH value of 2.0 was added, and the mixture was allowed to stand for 3 hours for impregnation. The water was removed by decantation, then dried at 120 ° C. for 3 hours, and then calcined at 400 ° C. for 3 hours in an air stream in a muffle furnace.

  The obtained fired product was cooled to room temperature in a desiccator, then added with 10 g of a 0.52 wt% -potassium nitrate aqueous solution, allowed to stand for 3 hours and then impregnated, and then water was removed by an evaporator, and then at 120 ° C. for 3 hours. The mixture was dried and reduced at 400 ° C. for 15 hours under a hydrogen flow to prepare a dehydrogenation catalyst. This dehydrogenation catalyst had a platinum loading of 0.6% by weight and a potassium loading of 0.1% by weight.

With respect to the dehydrogenation catalyst thus obtained, the dispersion degree and particle size of platinum were measured by a pulse adsorption method of CO using a fully automatic catalyst gas adsorption amount measuring device (Okura Riken Co., Ltd .: R6015). That is, the metal surface area was calculated assuming that one molecule of CO is adsorbed with respect to the lattice constant area a ² of platinum, and the metal dispersity and particle diameter were determined with the amount of platinum supported being 0.6% by weight. As a result, the degree of dispersion was 78% and the particle size was 6.3 mm.

[Dehydrogenation reaction]
10 cc of the dehydrogenation catalyst obtained above was placed in a stainless steel reaction tube having an inner diameter of 12.6 mmφ × 300 mm equipped with a 1 / 8-inch thermocouple protection tube at the center of the reaction tube cross section. Packed so that the center is located in the center of the length of the reaction tube, and then 10cc of 1mmφ spherical α-alumina beads as a preheating layer is packed on the upper side of the catalyst layer. A reaction test apparatus (dehydrogenation reactor) was used.

The catalyst layer of this dehydrogenation reactor was heated to a central temperature of 300 ° C under a hydrogen stream (LHSV = 5.0; 50cc / hr) at a pressure of 1kg / cm ² A, and then high performance liquid chromatography (HPLC) Supply liquid component (composition: toluene 65.01 mol% and methylcyclohexane 34.59 mol%) obtained by the above hydrogenation reaction using a liquid feed pump (HPLC pump) at a rate of 50 cc / hr. The supply of industrial hydrogen was stopped.

  Since the temperature of the catalyst layer was lowered by the endothermic reaction of the dehydrogenation reaction, the output of the heater was adjusted and the temperature of the catalyst layer was controlled at 300 ° C. to initiate the reaction. After 1 hour, the fraction at the outlet of the reaction tube was separated with a gas-liquid separator, the liquid components were sampled, and when toluene and methylcyclohexane were quantified by gas chromatography, the conversion of methylcyclohexane to toluene was 95 mol%. In terms of hydrogen conversion, the hydrogen conversion in the mixed gas was 85 mol%.

[Dehydrogenation catalyst life test]
In the same manner as in the above dehydrogenation reaction, a pressurized flow-type dehydrogenation reaction test apparatus (dehydrogenation reactor) was prepared, and a catalyst life test in the dehydrogenation reaction of methylcyclohexane was conducted. The dehydrogenation reaction was performed by supplying hydrogen at a reaction temperature of 300 ° C. together with methylcyclohexane having an LHSV of 2.0 h ^{−1 so} that the concentration of the supplied hydrogen in the feed was 20 mol%.

  At this time, the equilibrium conversion of methylcyclohexane was 96.0 mol%, the equilibrium conversion of methylcyclohexane 24 hours after the start of the reaction was 94.2 mol%, and the selectivity of toluene was 99.9 mol%. It was.

  Further, after 24 hours, the reaction temperature was raised to 320 ° C. and the dehydrogenation reaction was continued for 3000 hours. The equilibrium conversion rate of methylcyclohexane when the reaction temperature was raised to 320 ° C. was 99.8 mol%, and the equilibrium conversion rate of methylcyclohexane after 3000 hours was 95.2 mol%. The selectivity was maintained at 99.9 mol%.

  The method for purifying hydrogen according to the present invention is relatively easy to operate when a higher purity hydrogen is obtained by separating and removing impurity gases other than hydrogen from a relatively low-grade hydrogen-containing mixed gas. It can be continuously processed and can easily handle large-scale processing of hydrogen-containing mixed gas by scaling up, and ultra-high purity hydrogen is required for industrial applications such as various chemical processes. It can also be used for fuel cell applications and has high industrial utility value.

FIG. 1 is a flow sheet showing a preferred embodiment of the present invention.

Explanation of symbols

1 ... Hydrogenation reactor,
2, 5, 7, 9, 11 ... gas-liquid separator,
3 ... Preheater,
4 ... dehydrogenation reactor,
6 ... 10 atmosphere pressure booster, 8 ... 20 atmosphere pressure booster, 10 ... 50 atmosphere pressure booster.

Claims (8)

This is a method for purifying hydrogen by separating and removing impurity gases other than hydrogen from a hydrogen-containing mixed gas containing hydrogen as a main component to obtain higher-purity hydrogen, and the hydrogen-containing mixed gas and aromatics in the presence of a hydrogenation catalyst. The aromatic compound is hydrogenated by contacting with the compound, and the hydrogenated reaction mixture obtained by this hydrogenation reaction is gas-liquid separated to recover the hydrogenated aromatic compound, and then in the presence of a dehydrogenation catalyst. A hydrogen purification method for dehydrogenating a hydrogenated aromatic compound and recovering hydrogen generated by the dehydrogenation reaction ,
The dehydrogenation catalyst used for the dehydrogenation reaction of the hydrogenated aromatic compound has a surface area of 150 m ² / g or more, a pore volume of 0.55 cm ³ / g or more, an average pore diameter of 90 to 300 mm, and a fine pore volume. A porous γ-alumina carrier in which the proportion of pores having a pore size of 90 to 300 mm (pore size 90 to 300 mm) is 60% or more is at least one selected from platinum, palladium, rhodium, iridium and ruthenium. A method for purifying hydrogen, wherein the catalyst metal is a dehydrogenation catalyst supported with a particle size of 10 mm or less . The method for purifying hydrogen according to claim 1, wherein when recovering hydrogen generated in the dehydrogenation reaction, gas-liquid separation operation is performed by pressurizing to an absolute pressure of 0.981 MPa or more . Dehydrogenation catalyst used in the dehydrogenation of hydrogenated aromatic compounds, method for purifying hydrogen according to claim 1 or 2 together alkaline metals and the catalyst metal is dehydrogenation catalyst supported.   The method for purifying hydrogen according to any one of claims 1 to 3, wherein the aromatic compound used in the hydrogenation reaction is one or a mixture of two or more selected from benzene, toluene, naphthalene and methylnaphthalene.   The hydrogenation catalyst used for the hydrogenation reaction of the aromatic compound is a catalyst in which at least one catalyst metal selected from platinum, palladium, rhodium, iridium, ruthenium and nickel is supported on a catalyst carrier of γ-alumina or silica. The method for purifying hydrogen according to any one of claims 1 to 4.   The method for purifying hydrogen according to any one of claims 1 to 5, wherein the hydrogen-containing mixed gas has a hydrogen content of 30 to 80 mol%.   The method for purifying hydrogen according to any one of claims 1 to 6, wherein carbon monoxide and / or a sulfur compound which is a catalyst poison of a hydrogenation catalyst is previously removed from the hydrogen-containing mixed gas.   The hydrogenation region of the hydrogenation reaction, which is an exothermic reaction, and the dehydrogenation region of the dehydrogenation reaction, which is an endothermic reaction, are incorporated in one reactor, and the reaction heat generated by the hydrogenation reaction is used as the reaction heat of the dehydrogenation reaction. The method for purifying hydrogen according to any one of claims 1 to 7.

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