DESCRIPTION
METHOD FOR PRODUCTION OF HYDROGEN, METHOD FOR PRODUCTION OF RE-FORMULATED GASOLINE, AND METHOD FOR PRODUCTION OF AROMATIC HYDROCARBON Technical Field
[0001] The present invention relates to a process for producing hydrogen, a process for producing reformulated gasoline and a process for producing aromatic hydrocarbons. Background Art
[0002] Hydrogen is considered to be a highly clean fuel because its combustion produces only water. As production processes for hydrogen there may be mentioned electrolysis of water, and modification or partial oxidation of hydrocarbons. The most widely employed hydrogen production processes are steam reforming and catalytic reforming of hydrocarbons for petroleum refining. Of these, catalytic reforming of hydrocarbons is a process that uses light hydrocarbons such as naphtha as starting materials, and employs a catalyst to produce primarily hydrogen, along with high octane gasoline bases known as reformulated gasoline (for example, see Non-patent document 1).
[0003] Incidentally, reduction of CO2 emissions, especially from transportation fuel, is a major issue for combating global warming. Increasing interest is being directed toward a resource known as biomass with the goal of reducing CO2 emissions. That is, since plant-derived biomass absorbs CO2 by photosynthesis during the course of the plant growth, plant-derived biomass is a "carbon neutral" material

which is unlinked to atmospheric CO2 increase, as a result of natural life cycles. In other words, the CO2 emitted by combustion of biomass fuels is not counted as CO2 emission because it is equivalent to the C02 that has been fixed by plants. Consequently, biomass is thought to have tremendous potential toward achieving future C02 reduction. For example, biomass will take on a major significance if used for production of fuel oils such as light oil or kerosene light oil. The use of biomass especially for production of gasoline should lead to reduction in CO2 emissions from gasoline. Considering the wide use of gasoline, this is expected to provide a major contribution toward preventing global warming.
[0004] Also, hydrogen is an indispensable fuel for high energy efficiency fuel cells, and the demand for hydrogen is predicted to increase in the future. Yet technology for production of hydrogen with reduced CO2, i.e. reduced LCA-CO2 in consideration of natural life cycles, has not yet been established.
[Non-patent document 1] The Japan Petroleum Institute, "Petroleum Refining Processes", Kodansha, March 20, 1999, 101-119 Disclosure of the Invention Problems to be Solved by the Invention
[0005] It is therefore an object of the present invention, which has been accomplished in light of the circumstances described above, to provide a process for production of hydrogen with adequately effective LCA-CO2 reduction. In addition to this process for producing hydrogen, it is another object to provide a process for production of reformulated gasoline and a process for production of aromatic

hydrocarbons, also with adequately effective LCA-CO2reduction. Means for Solving the Problems
[0006] As a result of much diligent research directed toward achieving the object stated above, the present inventors have found that by using animal or plant derived fat or oil components as biomass, it is possible to produce hydrogen, reformulated gasoline and aromatic hydrocarbons without requiring special operating conditions or extra equipment investment, and the invention has been completed upon this find. [0007] Specifically, the invention is a process for producing hydrogen comprising a hydrocracking step in which a first stock oil or a second stock oil derived from the first stock oil is subjected to hydrocracking by contact with a hydrocracking catalyst in a hydrogen-containing atmosphere to obtain a first naphtha, and a catalytic reforming step in which the first naphtha or a second naphtha derived from the first naphtha is subjected to catalytic reforming to obtain hydrogen, wherein the first stock oil contains an animal or vegetable oil-derived fat or oil component, the animal or vegetable oil-derived fat or oil component including fractions with boiling points of 230°C or higher, the hydrocracking catalyst comprises at least one metal selected from the group consisting of metals belonging to Group 6A and Group 8 of the Periodic Table and an acidic inorganic oxide, and the first naphtha includes a fraction with boiling point of 100-120°C. [0008] This process for producing hydrogen allows production of hydrogen from biomass fat or oil components contained in the first stock oil, and can therefore serve as adequately effective means for LCA-CO2 reduction. In addition, the process for producing hydrogen

can yield LPG, gasoline, kerosene and light oil in addition to hydrogen, and in particular it is possible to efficiently and reliably produce liquid fractions such as gasoline, kerosene and light oil.
[0009] The process for producing hydrogen of the invention preferably also comprises a hydrotreatment step in which, prior to the aforementioned catalytic reforming step, the naphtha fraction containing the first naphtha is contacted with a second hydrogenation catalyst in a hydrogen-containing atmosphere, to obtain the second naphtha with a lower sulfur content and oxygen content than the naphtha fraction, wherein the second hydrogenation catalyst comprises at least one metal selected from the group consisting of metals belonging to Group 6 A and Group 8 of the Periodic Table, and the second naphtha has a sulfur content of no greater than 1.0 ppm by weight and an oxygen content of no greater than 10 ppm by weight.
[0010] The process for producing hydrogen of the invention preferably also comprises a hydrogenating pre-processing step in which, prior tc the hydrocracking step, the first stock oil is contacted with the firsi hydrogenation catalyst in a hydrogen-containing atmosphere to obtain i. second stock oil, wherein the first hydrogenation catalyst comprises a. least one metal selected from the group consisting of metals belonging to Group 6 A and Group 8 of the Periodic Table.
[0011] The hydrocracking conditions in the hydrocracking step of the process for producing hydrogen of the invention are preferably a hydrogen pressure of 6-20 MPa, a LHSV of 0.2-1.5 h-1 and a hydrogen/oil ratio of 200-2000 NL/L. [0012] The invention further provides a process for producing

reformulated gasoline wherein reformulated gasoline is obtained by the
catalytic reforming step described above. Since the reformulated
gasoline is obtained using a biomass fat or oil component as the starting
material, the amount of CO2 is adequately minimized in consideration
of natural life cycles, even though CO2 is generated by using the fuel.
In addition, the obtained reformulated gasoline is not inferior in quality
to gasoline produced from ordinary petroleum crude.
[0013] The invention still further provides a process for producing
aromatic hydrocarbon whereby C6-8 aromatic hydrocarbons are
obtained in the catalytic reforming step described above. This process
allows the aromatic hydrocarbons to be obtained with adequately
minimized CO2 emission, since biomass fat or oil components are used
as starting materials.
Effect of the Invention
[0014] According to the invention it is possible to provide a process
for producing hydrogen which is adequately effective for LCA-C02
reduction.
Brief Description of the Drawings
[0015] Fig. 1 is a flow chart showing an example of a hydrogen
production apparatus suitable for carrying out the process for producing
hydrogen of the invention.
Fig. 2 is a flow chart showing another example of a hydrogen
production apparatus suitable for carrying out the process for producing
hydrogen of the invention.
Explanation of Symbols
[0016] 100, 200: Hydrogen production apparatus, 110, 114, 120, 126,

214, 226: reaction column, 118, 124, 218: separation column, 112, 116,
122, 128, 216, 228: catalyst layer.
Best Modes for Carrying Out the Invention
[0017] Preferred embodiments of the invention will now be described in detail.
[0018] The process for producing hydrogen according to a preferred embodiment of the invention is a process for producing hydrogen comprising a hydrogenating pre-processing step in which a first stock oil is contacted with a first hydrogenation catalyst in a hydrogen-containing atmosphere to obtain a second stock oil, a hydrocracking step in which the second stock oil is subjected to hydrocracking by contact with a hydrocracking catalyst in a hydrogen-containing atmosphere to obtain a first naphtha, a hydrotreatment step in which the naphtha fraction including the first naphtha is contacted with a second hydrogenation catalyst in a hydrogen-containing atmosphere to obtain a second naphtha having a lower sulfur content and oxygen content than the naphtha fraction, and a catalytic reforming step in which the second naphtha is subjected to catalytic reforming to obtain hydrogen. [0019] According to this embodiment, an oil including fractions with boiling points of 230°C or higher and containing animal or vegetable oil-derived fat or oil components is used as the first stock oil. As examples of animal and vegetable oil-derived oils there may be mentioned beef tallow, rapeseed oil, soybean oil, palm oil and the like. The fat or oil component for this embodiment is not particularly restricted so long as it includes fractions with boiling points of 230°C or higher. The fat or oil component may be a single type or a mixture of

more than one type, or it may be the waste oil obtained after using such fats or oils. From the viewpoint of further promoting LCA-C02 reduction, a vegetable oil-derived fat or oil component is preferred. From the viewpoint of fatty acid alkyl chain carbon number and reactivity, a fat or oil component derived from rapeseed oil, soybean oil and/or palm oil is preferred. The concept of "animal or vegetable oil-derived fat or oil component" throughout the present specification encompasses not only naturally and artificially produced animal and vegetable oils, but also fat and oil components produced using such animal and vegetable fats and oils as starting materials. The fat and oil components used for the invention may also contain additives included to maintain and improve the quality and performance of various fat and oil products.
[0020] These fat and oil components generally have a fatty acid triglyceride structure. However, they also include fat and oil components that have been processed to fatty acids or to esters such as fatty acid methyl esters. Production of fatty acids or fatty acid esters from animal and vegetable oil-derived fat or oil components generates C02. From the viewpoint of promoting LCA-C02 reduction, therefore, the animal or vegetable oil-derived fat or oil component preferably consists mainly of components with a triglyceride structure. Specifically, the proportion of compounds with a triglyceride structure in the animal or vegetable oil-derived fat and oil components is preferably 80 mol% or greater, more preferably 85 mol% or greater and even more preferably 90 mol% or greater. [0021] The fat or oil components are not particularly restricted so long

as they include fractions with boiling points of 230°C or higher. However, the fat or oil components preferably include fractions with boiling points of 250°C or higher, more preferably include fractions with boiling points of 300°C or higher, and even more preferably include fractions with boiling points of 360°C or higher. If the fat or oil components do not include fractions with boiling points of 230°C and higher, the production of gas fractions such as LPG in addition to hydrogen will increase in the process for producing hydrogen of this embodiment. This will tend to reduce the yield of the naphtha fraction and heavier fractions (hereinafter referred to as "liquid fractions"). As will be explained in more detail below, the animal or vegetable oil-derived fat or oil components may be pre-processed by hydrogenation before the hydrocracking treatment.
[0022] The oxygen content of the first stock oil is preferably in the range of 0.1-13 wt%, more preferably in the range of 0.2-12 wt.% and even more preferably in the range of 0.5-11 wt%. If the oxygen content is below 0.1 wt%, the amount of fat or oil components in the first stock oil will be lower, tending to lessen the LCA-C02 reduction effect. The oxygen content is preferably not greater than 13 wt% in order to avoid the need for equipment to treat water by-product and to avoid reduced catalyst power due to interaction between water and the catalyst carrier. Throughout the present specification, the oxygen content of the first stock oil, etc. is measured by a publicly known method using an ordinary elemental analyzer. For example, the oxygen in a sample to be measured may be converted to CO on platinum-carbon, or further converted to €02= and then measured using

a thermal conductivity detector.
[0023] The first stock oil may consist of a blend of the animal or vegetable oil-derived fat or oil component with a petroleum-based hydrocarbon fraction. The hydrocarbon fraction used may be a fraction obtained by any common petroleum refining treatment. As a specific example, there may be used a fraction corresponding to a prescribed boiling point range obtained from an atmospheric distillation apparatus or vacuum distillation apparatus, or a fraction corresponding to a prescribed boiling point range obtained from a hydrodesulfurizer, hydrocracker, bottom oil direct desulfurizer or fluidized catalytic cracker. These may be fractions corresponding to prescribed boiling point ranges from one apparatus, or fractions corresponding to prescribed boiling point ranges from different apparatuses. The petroleum-based hydrocarbon fraction preferably includes fractions with boiling points of 340°C or higher, and more preferably includes fractions with boiling points of 700°C or higher. If the petroleum-based hydrocarbon fraction does not include fractions with boiling points of 340°C and higher, excessive decomposition in the hydrocracking step will tend to lower the liquid fraction yield. If the petroleum-based hydrocarbon fraction includes heavy fractions with boiling points of 700°C or higher, those fractions will promote formation of carbonaceous substances on the catalyst, thus covering the catalyst active site and tending to lower the catalytic activity. Throughout the present specification, "boiling point", "boiling point range" and "running point" refer to the values measured according to the method specified in JIS-K2254, "Distillation Test Method" or ASTM-

D86.
10024] When a petroleum-based hydrocarbon fraction is blended with the first stock oil, the blending ratio of the hydrocarbon fraction is preferably 10-99 vol%, more preferably 30-99 vol% and even more preferably 60-98 vol% with respect to the total volume of the first stock oil. If the blending ratio of the petroleum-based hydrocarbon fraction is below the aforementioned lower limit, equipment may be required for treatment of water by-product. If the blending ratio of the petroleum-based hydrocarbon fraction exceeds the aforementioned upper limit, the LCA-CO2 reduction effect will tend to be lower.
[0025] In the hydrogenating pre-processing step for this embodiment, the first stock oil is contacted with the first hydrogenation catalyst in a hydrogen-containing atmosphere and preferably in a pressurized atmosphere, to obtain the second stock oil. The hydrogenating preprocessing step can reduce the oxygen content of the second stock oil to below that of the first stock oil.
[0026] The respective volumes of the first hydrogenation catalyst used for the hydrogenating pre-processing step and the hydrocracking catalyst described hereunder may be set as desired for satisfactory hydrocracking activity in the hydrocracking step. The volume ratio of the first hydrogenation catalyst with respect to the total catalyst volume is preferably 10-90 vol% and more preferably 25-75 vol%. If the volume ratio is below the aforementioned lower limit, it will tend to be difficult to reduce the oxygen content in the second stock oil obtained by treatment of the first stock oil with the first hydrogenation catalyst. If the volume ratio exceeds the aforementioned upper limit, the

hydrocracking reaction may be hindered in the hydrocracking step. [0027] The oxygen content of the second stock oil obtained from the hydrogenating pre-processing step is preferably reduced to not greater than 40 wt% and more preferably reduced to not greater than 30 wt% of the oxygen content of the first stock oil. When the second stock oil is contacted with the hydrocracking catalyst, the oxygen in the second stock oil poisons the active site of the catalyst. Thus if the oxygen content exceeds 40 wt% of the oxygen content of the first stock oil, the hydrocracking activity will tend to be reduced.
[0028] A catalyst other than the first hydrogenation catalyst and hydrocracking catalyst and/or a filler may also be used if necessary. For example, guard catalysts, demetallizing catalysts or inert fillers may be used either alone or in combination for the purpose of trapping the scale portion flowing in with the first stock oil, or supporting the first hydrogenation catalyst and hydrocracking catalyst in separated sections of the catalyst bed. A catalyst with hydrogenating activity may also be used in a stage after the hydrocracking catalyst for the purpose of stabilizing hydrogenation of the decomposition product obtained from the hydrocracking step.
[0029] The reaction temperature in the hydrogenating pre-processing step and hydrocracking step may be set as desired to achieve the desired crack-per-pass for the heavy fraction of the first stock oil and to achieve the prescribed fractions at the intended yields. Also, the reaction temperature in the hydrogenating pre-processing step and the reaction temperature in the hydrocracking step may each be set so as to limit the oxygen content of the second stock oil obtained from the hydrogenating

pre-processing step to below the upper limit mentioned above. Trie first hydrogenation catalyst and hydrocracking catalyst may be packed into separate reactors, or they may be packed together into the same reactor. When the first hydrogenation catalyst and hydrocracking catalyst are packed into the same reactor, the mean temperature for the entire reactor is set in a range of usually 330-480°C, preferably 350-450°C, and more preferably 360-430°C, in order to satisfactorily promote the reaction, and produce hydrogen at the prescribed yield and gasoline, kerosene and light oil with the prescribed properties. If the mean temperature is below the aforementioned lower limit the reaction may not proceed adequately, while if it exceeds the aforementioned upper limit, hydrocracking may proceed excessively, resulting in a reduced liquid fraction yield.
[0030] The active metal of the first hydrogenation catalyst comprises at least one metal selected from the group consisting of metals belonging to Group 6A and Group 8 of the Periodic Table. The active metal preferably comprises two or more metals selected from the group consisting of metals belonging to Group 6A and Group 8. As examples of such active metals there may be mentioned Co-Mo, Ni-Mo, Ni-Co-Mo and Ni-W, and these metals may also be converted to sulfides for use in the hydrogenating pre-processing. [0031] The carrier used for the first hydrogenation catalyst is preferably a porous inorganic oxide. Specifically, the carrier may be, for example, an alumina-containing porous inorganic oxide. As carrier constituent components that are inorganic oxides other than alumina there may be mentioned silica, titania, zirconia and boria. The carrier

is preferably a complex oxide comprising alumina and one or more selected from the group consisting of the other constituent components mentioned above. Phosphorus may also be included as another component in the carrier.
[0032] The total content of components other than alumina in the carrier is preferably 1-20 wt% and more preferably 2-15 wt% based on the total weight of the carrier. If the content is less than 1 wt%, the surface area of the catalyst may be reduced, thereby lowering the activity. If the content is greater than 20 wt% the acidic substance concentration of the carrier will tend to increase, leading to activity reduction as a result of coke production. When phosphorus is included as a constituent component of the carrier, its content is preferably 1-8 wt% and more preferably 2-5 wt% in terms of oxide (P2O5). [0033] There are no particular restrictions on the starting materials used as precursors of silica, titania, zirconia, boria and the like as carrier constituent components other than alumina, and usually a solution containing silicon, titanium, zirconium or boron may be used. For example, silicic acid, water glass, silica sol and the like may be used for silicon, titanium sulfate, titanium tetrachloride or various alkoxide salts may be used for titanium, zirconium sulfate or various alkoxide salts may be used for zirconium, and boric acid may be used for boron. As phosphorus starting materials there may be used phosphoric acid or phosphoric acid alkali metal salts and the like.
[0034] The starting materials for carrier constituent components other than alumina are preferably added to the first hydrogenation catalyst starting material at some stage prior to firing of the carrier during

preparation of the first hydrogenation catalyst. For example, the constituent component starting materials may be pre-added to an aluminum aqueous solution to obtain an aluminum hydroxide gel containing the constituent components. Alternatively, the constituent component starting materials may be added to a previously prepared aluminum hydroxide gel. The constituent component starting materials may instead be added in a step in which water or an acidic aqueous solution is added to and kneaded with a commercially available alumina intermediate or boehmite powder. Preferred among these methods is addition of the constituent component starting materials to aluminum oxide during the stage of preparing the aluminum hydroxide gel. The mechanism by which the carrier constituent components other than alumina exhibit their effects has not yet been fully elucidated, but the present inventors conjecture that the constituent components form complex oxides with aluminum. Presumably this increases the carrier surface area and leads to some kind of interaction of the carrier with the active metal, thereby affecting the catalytic activity. [0035] The content of these active metals in the first hydrogenation catalyst, when W and/or Mo are used as active metals for example, is a total loading weight of W and Mo of preferably 12-35 wt% and more preferably 15-30 wt% in terms of oxides (W03, MoO3) with respect to the catalyst weight. If the loading weight is below the aforementioned lower limit, the number of catalyst active sites will be reduced and the catalytic activity will thus tend to be lower. If the loading weight is above the aforementioned upper limit, the active metals may not effectively disperse, also leading to reduced catalytic activity.

[0036] When Co and/or Ni are used as active metals, the total loading weight of the Co and Ni is preferably 1.5-18 wt% and more preferably 2-15 wt% in terms of oxides (CoO, NiO) with respect to the catalyst weight. If the loading weight is less than 1.5 wt%, a cocatalyst effect will not be easily achieved and the activity will tend to be reduced. If the loading weight is greater than 18 wt%, the active metals will not effectively disperse, also leading to reduced catalytic activity. [0037] In the hydrocracking step of this embodiment, the second stock oil obtained from the hydrogenating pre-processing step is then subjected to hydrocracking by contact with a hydrocracking catalyst in a hydrogen-containing atmosphere and preferably a pressurized atmosphere, to obtain a first naphtha.
[0038] The hydrocracking conditions in the hydrocracking step are preferably a hydrogen pressure of 6-20 MPa, a liquid space velocity (LHSV) of 0.2-1.5 h_1 and a hydrogen/oil ratio of 200-2000 NL/L, more preferably a hydrogen pressure of 8-17 MPa, a LHSV of 0.2-1.1 h"1 and a hydrogen/oil ratio of 300-1800 NL/L, and even more preferably a hydrogen pressure of 10-16 MPa, a LHSV of 0.3-0.9 h"1 and a hydrogen/oil ratio of 350-1600 NL/L. When a hydrogenating preprocessing step is performed, the conditions for both the hydrogenating pre-processing step and hydrocracking step are preferably a hydrogen pressure of 6-20 MPa,, a liquid space velocity (LHSV) of 0.2-1.5 h"1 and a hydrogen/oil ratio of 200-2000 NL/L, more preferably a hydrogen pressure of 8-17 MPa, a LHSV of 0.2-1.1 h-1 and a hydrogen/oil ratio of 300-1800 NL/L, and even more preferably a hydrogen pressure of 10-16 MPa, a LHSV of 0.3-0.9 h"1 and a hydrogen/oil ratio of 350-1600 NL/L.

[0039] Here, "LHSV (liquid hourly space velocity)" means the volume flow rate of the stock oil under standard conditions (25°C, 101.325 kPa) per volume of the catalyst layer packed with the catalyst. The unit "h" ]" represents inverse hours. The term. "NL" as the unit of hydrogen volume in "NL/L" as the unit ordinarily used for hydrogen/oil ratio represents the hydrogen volume (L) in a normal state (0°C, 101325 Pa). The reaction temperature is the mean temperature of the catalyst layer. [0040] All of these conditions affect the activity in the hydrocracking reaction. For example, if the hydrogen pressure and hydrogen/oil ratio are below the aforementioned lower limits, the reactivity will tend to be reduced and activity will tend to decline rapidly. If the hydrogen pressure and hydrogen/oil ratio are above the upper limits mentioned above, major equipment investment may be necessary for a compressor or the like. A lowr LHSV will tend to favor the reaction, but if it is below the aforementioned lower limit a very large reaction column volume may be necessary, tending to increase the equipment investment. If the LHSV is above the aforementioned upper limit, the reaction may be inhibited.
[0041] The form of the reactor in which the first hydrogenation catalyst and/or hydrocracking catalyst are packed may be a fixed bed system. Specifically, the hydrogen may be introduced in a crosscurrent or cocurrent flow with respect to the first stock oil or second stock oil. The reactor may be a single one or a combination of different reactors. When multiple reactors are combined, one reactor may be in a crosscurrent flow while the other reactors are in a cocurrent flow. The most common form of reactor is in a downflow, in which case a gas-

liquid cocurrent flow system may be employed. The interior of a single reactor may also have a structure that is partitioned into multiple catalyst beds.
[0042] The product oil obtained by hydrocracking the second stock oil in the reactor in the hydrocracking step of this embodiment may if necessary be fractionated into prescribed fractions via an additional gas-liquid separation step, rectification step or the like. When moisture is produced during the reaction or sulfur is present in the first stock oil, hydrogen sulfide may be generated. To counter this, gas-liquid separation equipment or other by-product gas removal equipment may be installed between multiple reactors or within the flow route of the product oil. The product oil may be fractionated in this manner to obtain the first naphtha.
[0043] The hydrogen gas used in the hydrogenating pre-processing step and hydrocracking step for this embodiment is introduced through the inlet port of the first reactor together with the first stock oil either before or after it passes through a heating furnace. However, hydrogen gas may also be introduced between catalyst beds or between multiple reactors in order to control the temperature in the reactors and keep the hydrogen pressure as constant as possible throughout all of the reactor interiors. Hydrogen gas introduced into a system in this manner is referred to as "quench hydrogen". In this case, the proportion of quench hydrogen with respect to the total volume including the quench hydrogen and the hydrogen introduced with the stock oil is preferably at least 10-60 parts by volume and more preferably at least 15-50 parts by volume. If the quench hydrogen proportion is below the

aforementioned lower limit, the reaction may be inhibited at the reaction site in later stages. If the quench hydrogen proportion exceeds the aforementioned upper limit, the reaction may be inhibited near the reactor inlet port.
[0044] The hydrocracking catalyst comprises at least one metal selected from the group consisting of metals belonging to Group 6A and Group 8 of the Periodic Table, and preferably it comprises two or more metals selected from the group consisting of metals belonging to Group 6A and Group 8. As specific examples there may be mentioned Co-Mo, Ni-Mo, Ni-Co-Mo and Ni-W. Ni-Mo, Ni-Co-Mo and Ni-W are preferred among these. These metals are converted to sulfides for use during the hydrocracking, in the same manner as for the first hydrogenation catalyst.
[0045] The hydrocracking catalyst preferably comprises an acidic inorganic oxide, with the inorganic oxide serving as the carrier. Preferred inorganic oxides are complex oxides containing two or more compounds selected from the group consisting of silica, alumina, boria, zirconia, magnesia and zeolite. For example, such complex oxides are preferably one or more selected from the group consisting of silica-alumina, titania-alumina, boria-alumina, zirconia-alumina, titania-zirconia-alumina, silica-boria-alumina, silica-zirconia-alumina, silica-titania-alumina and silica-titania-zirconia-alumina, more preferably one or more selected from the group consisting of silica-alumina, boria-alumina, zirconia-alumina, titama-zirconia-alumina, silica-boria-alumina, silica-zirconia-alumina and silica-titania-alumina, and even more preferably silica-alumina and/or silica-zirconia-alumina. These

complex oxides most preferably further contain zeolite. When the carrier contains alumina, the proportion of alumina to the other components may be any desired proportion with respect to the carrier. However the alumina content is preferably no greater than 96 wt% and more preferably no greater than 90 wt% with respect to the carrier weight. If the alumina content is greater than 96 wt% of the carrier weight, the amount of acidic substances will be reduced, and this may prevent the prescribed hydrocracking activity from being exhibited. [0046] When the hydrocracking catalyst contains zeolite, components of the zeolite crystal scaffold may include alumina, titania, boria, gallium and the like in addition to silica. Among these, zeolite containing silica and alumina, i.e. aluminosilicate, is preferred. Numerous types of zeolite crystal structures are known, such as faujasite, beta, mordenite and pentasil types. Faujasite, beta and/or pentasil types are preferred and faujasite and/or beta types are more preferred, from the viewpoint of exhibiting more satisfactory hydrocracking activity for this embodiment.
[0047] Such zeolites may have adjusted alumina contents in a
stoichiometric ratio of the starting material at the start of synthesis.
Alternatively, such zeolites may be subjected to specific hydrothermal
treatment and/or acid treatment. Particularly preferred is
ultrastabilized zeolite Y that has been ultrastabilized by hydrothermal treatment and/or acid treatment. Ultrastabilized zeolite Y has newly formed pores with pore diameters of 20-100 A in addition to the original zeolite microporous structure consisting of "micropores" with pore diameters of 20 A and smaller. It is believed that ultrastabilized zeolite

Y provides suitable reaction sites for conversion of the oxygen in the fat or oil component. The volume of the pores with pore diameters of 20-1.00 A (pore volume) is preferably at least 0.03 mL/g and more preferably at least 0.04 mL/g. The "pore volume" referred to here can in most cases be determined by mercury porosimetry. [0048] The hydrothermal treatment conditions for synthesis of the zeolite may be any publicly known conditions.
[0049] The silica/alumina molar ratio, as a property of the ultrastabilized zeolite Y, is preferably 10-120, more preferably 15-70 and even more preferably 20-50. If the silica/alumina molar ratio is higher than 120, the amount of acidic substances will be reduced, tending to lower the hydrocracking activity. If the silica/alumina molar ratio is less than 10, the acidity may be too strong and accelerated coke production reaction may drastically reduce the activity. [0050] The zeolite content of the carrier is preferably 2-80 wt% and more preferably 4-75 wt% with respect to the carrier weight. If the zeolite content is below the aforementioned lower limit, the hydrocracking activity may not be readily exhibited. If the zeolite content is above the aforementioned upper limit, the acidity may be too strong, tending to promote coke production.
[0051] There are no particular restrictions on the method of adding the active metals to the first hydrogenation catalyst and hydrocracking catalyst, and any publicly known method used for production of ordinary desulfurization catalysts may be applied. For example, a method of impregnating the catalyst carrier with a solution containing salts of the active metals will usually be appropriate. Other suitable

methods include the equilibrium adsorption, pore filling and incipient wetness methods. For example, the pore filling method involves first measuring the pore volume of the carrier and then impregnating it with an equal volume of the metal salt solution. The impregnation method in this case is not particularly restricted, and may be a method that is suitable for the metal loading weight and the physical properties of the catalyst carrier.
[0052] The first naphtha obtained from the hydrocracking step contains fractions with boiling points of 100-120°C. The first naphtha may be used alone or in a blend with other naphthas, to obtain the naphtha fraction as the stock oil for the hydrotreatment step described hereunder. As examples of other naphthas there may be mentioned straight run naphtha, as well as cracked naphtha obtained from a catalytic cracker, naphtha produced by hydrodesulfurization, condensates, rafiinates from aromatic solvent extractors, and the like. 100531 The reaction product obtained from the hydrocracking step includes the first naphtha, lighter gas fractions such as light propane, and heavier oil fractions. The first naphtha is therefore separated as necessary into hydrogen gas, the light gas fraction and the heavy oil fraction by publicly known methods.
[0054] In the hydrotreatment step of this embodiment, the naphtha fraction containing the first naphtha obtained from the hydrocracking step is contacted with a second hydrogenation catalyst in a hydrogen-containing atmosphere, and preferably a pressurized atmosphere, to obtain a second naphtha with a lower sulfur content and oxygen content than the naphtha fraction. This will allow the sulfur and oxygen

contents of the stock oil (second naphtha) of the catalytic reforming step described hereunder to be reduced to the prescribed concentrations. The second hydrogenation catalyst used in the hydrotreatment step preferably comprises at least one metal selected from the group consisting of metals belonging to Group 6 A and Group 8 of the Periodic Table. This can further reduce the sulfur and oxygen contents of the second naphtha.
[0055] The second hydrogenation catalyst used may be an ordinary hydrogenation catalyst. As examples of active metals for the second hydrogenation catalyst there may be mentioned Co-Mo, Ni-Mo and Ni-Co-Mo, and these metals may also be converted to sulfides for use in the hydrotreatment step. There are no particular restrictions on the method of adding the active metals to the second hydrogenation catalyst, and any publicly known method used for production of ordinary desulfurization catalysts may be applied. For example, a method of impregnating a catalyst carrier with a solution containing salts of the active metals will usually be appropriate. Other suitable methods include the equilibrium adsorption, pore filling and incipient wetness methods. For example, the pore filling method involves first measuring the pore volume of the carrier and then impregnating it with an equal volume of the metal salt solution. The impregnation method in this case is not particularly restricted, and may be a method that is suitable for the metal loading weight and the physical properties of the catalyst carrier.
[0056] The carrier used for the second hydrogenation catalyst may usually be an inorganic porous carrier composed mainly of alumina.

There are no particular restrictions on the method of preparing the carrier, and any method, for preparing ordinary alumina-containing carriers may be applied.
[0057] The reaction product obtained from the hydrotreatment step includes the second naphtha and lighter gas fractions such as light propane. The second naphtha is therefore separated as necessary into hydrogen gas and the lighter gas fractions by a publicly known method. [0058] In the catalytic reforming step of this embodiment, the second naphtha obtained from the hydrotreatment step is subjected to catalytic reforming to obtain hydrogen. The catalytic reforming step may employ an ordinary catalytic reformer and the same conditions used for petroleum refining.
[0059] The catalyst used for the catalytic reforming step may be a commonly employed catalytic reforming catalyst. As a specific example, there may be used an alumina carrier loaded with Pt and/or an element belonging to Group 7 A of the Periodic Table, or Pt and/or an element belonging to Group 4B of the Periodic Table. Such catalysts are said to prevent aggregation of active metal Pt particles on catalysts. Preferred combinations of Pt and the aforementioned elements are Pt-Re, Pt-Sn and Pt-Ge.
[0060] Also, loading of chlorine on the catalyst or addition of chlorine compounds to the stock oil for catalytic reforming (second naphtha) during operation may be carried out beforehand, to supply chlorine to the catalyst and maintain the dispersed state of the Pt particles. [0061] The catalyst used may be regenerated for reuse by combustion removal of the adhered coke. There are no particular restrictions on

the method of regenerating the catalyst, and any regenerating method commonly carried out with catalytic reformers in petroleum refining steps may be used. Specifically there may be mentioned fixed bed regeneration wherein operation is periodically paused and an oxygen-containing gas is introduced into the reactor with the catalyst packed in the reactor for combustion removal of the coke, cyclic regeneration wherein only one reactor is cut off and the catalyst is regenerated in the same manner as sequential regeneration, and continuous regeneration wherein a moving bed reaction system is employed and the catalyst is continuously extracted from the reactor and regenerated in the same manner in a regenerator, and then returned to the reactor. In order to improve the reformulated gasoline yield and aromatic yield in the catalytic reforming step, the regenerating conditions are preferably lower pressure and high temperature. Also, continuous regeneration is preferred as the regenerating method so as to allow adjustment for severe catalyst use conditions.
10062] The catalyst used in the catalytic reforming step will tend to be easily poisoned by sulfur compounds in the stock oil (second naphtha). Thus, the sulfur content of the second naphtha is preferably no greater than 1 ppm by weight and more preferably no greater than 0.5 ppm by weight. The "sulfur content" referred to throughout the present specification is the content measured by the method described in JIS-K2541, "Sulfur Content Test Method".
1.0063] A high oxygen content in the stock oil (second naphtha) will lead to reaction of the oxygen with chlorine in the catalyst, thereby promoting outflow of oxygen from the catalyst. This may lead to

aggregation of particles of the loaded elements such as Pt, thereby tending to reduce the catalytic activity. Consequently, the oxygen content of the second naphtha is preferably adjusted to no greater than 10 ppm by weight and more preferably no greater than 5 ppm by weight. Trace oxygen in the second naphtha is measured by a method using a thermal conductivity detector after conversion of the oxygen in the measuring sample to CO., or further conversion to C02, on platinum-

carbon, by C-NMR (nuclear magnetic resonance) or by gas
chromatography with an atomic emission detector.
[0064] In the catalytic reforming step, aromatic hydrocarbons are
produced from linear saturated hydrocarbons (paraffins) and cyclic
saturated hydrocarbons (riaphthenes) by dehydrogenation, isomerizing
dehydrogenation or cyclizing dehydrogenation. Since
dehydrogenation is an endothermic reaction, a heat supply flow may be provided by using a plurality of reactors in the catalytic reforming process and installing a heating furnace between the reactors. [0065] There are no particular restrictions on the reaction configuration for the catalytic reforming step, but in order to minimize pressure loss in the reactor, the preferred configuration is a "radial flow" in which the starting material is passed from the periphery of the reactor through the catalyst layer and extracted into a center pipe at the center of the reactor.
[0066] The reaction conditions for the catalytic reforming step may be, for example, a pressure of up to 4 MPa, a reaction temperature of 400-600°C and a recycled hydrogen/hydrocarbon proportion of 0.1-10 mol/mol. A pressure of higher than 4 MPa will disfavor cyclizing

dehydrogenation, and will tend to reduce the octane number of the
product oil or lower the aromatic yield. If the reaction temperature is
lower than 400°C the reaction will be inhibited, thereby tending to
reduce the octane number of the product oil or lower the aromatic yield.
If the reaction temperature is higher than 600°C, decomposition reaction
will be promoted, thus tending to lower the reformulated gasoline yield
or reduce the catalyst life due to accelerated coke production. A
recycled hydrogen/hydrocarbon proportion of lower than 0.1 mol/mol
will tend to accelerate coke production. A recycled
hydrogen/hydrocarbon proportion of higher than 10 mol/mol will tend to increase the recycling operation cost or the energy consumption. [0067] The catalytic reforming step may be followed by a step of separating the hydrogen, light hydrocarbons and gasoline fraction from the product oil or a step of chlorine removal, and there are no particular restrictions on the configurations of the apparatuses used for these steps. [0068] The method of separating and recovering the hydrogen obtained in the catalytic reforming step from the light gas or other inorganic gas components may be a recovery process commonly used for petroleum refining. Specifically, there may be employed a pressure swing adsorption process (PSA process) using an adsorbent such as zeolite or silica gel, or a membrane separation process utilizing a polymer or inorganic material membrane. The PSA process is applicable for a wide range of impurities and can yield high-purity hydrogen. The membrane separation process employs a relatively inexpensive apparatus and has low energy consumption. These processes may be selected as desired depending on the composition of
26

gas components obtained by the catalytic reforming step, as well as the other conditions.
10069] The hydrogen produced according to this embodiment is thus produced using atmospheric C02-fixed substances as starting materials, and this is considered to be equivalent to reducing C02 emissions in each step. The hydrogen may therefore be considered environmentally friendly hydrogen.
[0070] According to this embodiment, the reformulated gasoline obtained from the catalytic reforming step may be suitably used as a gasoline base. Reformulated gasoline is rich in aromatic compounds and typically has a high research octane number (RON), and it is therefore used as a blending base for regular gasoline and high octane gasoline.
(0071] The reformulated gasoline obtained from the catalytic reforming step of this embodiment also contains abundant benzene, toluene and xylene which are important as chemical starting materials. Thus, reformulated gasoline can be suitably used as a starting material for these aromatic hydrocarbons. The C6-8 aromatic hydrocarbons in the reformulated gasoline may be fractionated into desired fractions containing the aromatic compounds in a rectifier, and then supplied as starting materials for petrochemical plants after increasing the purity of each aromatic hydrocarbon. As processes for increasing the aromatic hydrocarbon purity there may be mentioned aromatic solvent extraction processes, hydrodealkylation processes, aromatic alkyl group disproportionation processes, transalkylation processes, isomerization processes, adsorption separation processes, crystallizing separation

processes and the like. The aromatic hydrocarbons produced according to this embodiment or the various petrochemicals obtained using such compounds as starting materials are produced using atmospheric C02-fixed compounds as starting materials. [0072] As explained above, the catalytic reforming step of the present invention produces not only hydrogen but also reformulated gasoline. In addition, aromatic hydrocarbons such as benzene, toluene and xylene are also produced from reformulated gasoline.
[0073] A hydrogen production apparatus used for carrying out the process for producing hydrogen according to the invention will now be explained. Fig. 1 is a flow chart showing an example of a hydrogen production apparatus suitable for carrying out the process for producing hydrogen of the invention. The reaction column 110 of the hydrogen production apparatus 100 shown in Fig. 1 is a fixed bed reaction column. A first hydrogenation catalyst layer 112 is provided inside it for hydrogenating pre-processing. At the top of the reaction column 110 there is connected a line LI01 for supply of the oil to be treated (first stock oil) into the reaction column 110, while a line LI02 is connected upstream from the connection of the line LI 01 with the reaction column 110 for supply of hydrogen. Also, at the bottom of the reaction column 110 there is connected a line L103 for removal of reaction product containing the second stock oil from the reaction column 110. The other end of the line LI03 is connected to the top of a reaction column 114, and the reaction product containing the second stock oil is fed from the line LI03 to the reaction column 114. [0074] The reaction column 114 is a fixed bed reaction column, and a

hydrocracking catalyst layer 116 for hydrocracking is provided inside it. Also, at the bottom of the reaction column 114 there is connected a line L104 for removal of reaction product containing the first naphtha from the reaction column 114. The other end of the line LI04 is connected to the side of a separation column 118, and the reaction product containing the first naphtha is fed from the line LI04 to the separation column 118.
[0075] The separation column 118 separates the first naphtha from the reaction product produced by reaction at the reaction column 114. Separation treatment at the separation column 118 causes the light gas fraction and hydrogen gas that are lighter than the first naphtha to run out from the line LI05 connected to the top of the separation column 118. The fraction of heavier oil than the first naphtha runs out from the line LI07 connected to the bottom of the separation column 118. The first naphtha runs out from the line LI06 connected to the side of the separation column 118. The other end of the line LI06 is connected to the top of the reaction column 120.
(0076] The reaction column 120 is a fixed bed reaction column, and a second hydrocracking catalyst layer 122 for hydrotreatment is provided inside it. A line LI09 for supply of hydrogen is connected upstream from the connection of the line LI06 with the reaction column 120. Also, a line LI08 for supply of other naphtha in addition to the first naphtha is connected upstream from the connection of the line LI06 with the line LI09. The first naphtha extracted from the separation column 118 to the line L106 merges with the other naphtha from the Sine LI08 to form the naphtha fraction, and after further merging with

hydrogen from the line LI09, it is fed to the reaction column 120. At the bottom of the reaction column 120 there is connected a line LllO for removal of reaction product containing the second naphtha from the reaction column 120. The other end of the line LllO is connected to the side of a separation column 124, and the reaction product containing the second naphtha is fed from the line LllO to the separation column 124.
[0077] The separation column 124 separates the second naphtha from the reaction product produced by reaction at the reaction column 120. Separation treatment at the separation column 124 causes the light gas fraction and hydrogen gas that are lighter than the second naphtha to run from the line LI 11 connected to the top of the separation column 124. The second naphtha runs out from the line LI 12 connected to the bottom of the separation column 124. The other end of the line LI 12 is connected to the top of a reaction column 126, and the second naphtha is fed from the line LI 12 to the reaction column 126. [0078] The reaction column 126 is a fixed bed reaction column, and a catalyst layer 128 for catalytic reforming is provided inside it. At the bottom of the reaction column 126 there is connected a line LI 13 for removal of reaction product containing the hydrogen gas and reformulated gasoline from the reaction column 126. The other end of the line LI 13 is not shown but is connected to an apparatus for separation of, for example, reformulated gasoline, hydrogen gas and gas fractions lighter than the reformulated gasoline.
[0079] The present invention is not limited to the preferred embodiments described above. For example, according to another

embodiment of the invention, the hydrogenating pre-processing step may be omitted, in which case the first stock oil serves as the stock oil for the hydrocracking step. According to yet another embodiment, the hydrotreatment step may be omitted, in which case the naphtha fraction comprising the first naphtha alone or the aforementioned other naphthas is subjected to catalytic reforming in the catalytic reforming step. 10080] According to still another embodiment, both the hydrogenating pre-processing step and the hydrotreatment step may be omitted. An, example of a hydrogen production apparatus for this embodiment is shown by the flow chart in Fig. 2. The reaction column 214 of the hydrogen production apparatus 200 shown in Fig. 2 is a fixed bed reaction column, and a hydrocracking catalyst layer 216 for hydrocracking is provided inside it. At the top of the reaction column 214 there is connected a line L201 for supply of the oil to be treated (first stock oil) into the reaction column 214, while a line L202 is connected upstream from the connection of the line L201 with the reaction column 214 for supply of hydrogen. At the bottom of the reaction column 214 there is connected a line L204 for removal of reaction product containing the first naphtha from the reaction column 214. The other end of the line L204 is connected to the side of a separation column 218, and the reaction product containing the first naphtha is fed from the line L204 to the separation column 218. [0081] The separation column 218 separates the first naphtha from the reaction product produced by reaction at the reaction column 214. Separation treatment at the separation column 218 causes the light gas fraction and hydrogen gas that are lighter than the first naphtha to run

from the line L205 connected to the top of the separation column 218. The fraction of heavier oil than the first naphtha runs out from the line L207 connected to the bottom of the separation column 218. The first naphtha runs out from the line L206 connected to the side of the separation column 218. The other end of the line L206 is connected to the top of a reaction column 226, and the first naphtha is fed from the line L206 to the reaction column 226.
[0082] The reaction column 226 is a fixed bed reaction column, and a catalyst layer 228 for catalytic reforming of the first naphtha is provided inside it. At the bottom of the reaction column 226 there is connected a line L213 for removal of reaction product containing the hydrogen gas and reformulated gasoline from the reaction column 226. The other end of the line L213 is not shown but is connected to an apparatus for separation of, for example, reformulated gasoline, hydrogen gas and gas fractions that are lighter than the reformulated gasoline. [Examples]
[0083] The present invention will now be explained in greater detail based on examples and comparative examples, with the understanding that the invention is in no way limited to the examples. [0084] (Example 1)
Water glass No.3 was added to 1 kg of a 5 wt% sodium aluminate aqueous solution, and the mixture was placed in a vessel that had been heated to 70°C. After placing 1 kg of a 2.5% concentration aluminum sulfate aqueous solution in a separate vessel that had been heated to 70°C, it was added dropwise over a period of 15 minutes to the aforementioned sodium aluminate aqueous solution and water glass

mixture. The amount of water glass was adjusted to the prescribed silica content (listed in Table 1). The end point for the dropwise addition was defined as when the pH of the mixed solutions reached 6.9-7.5. The obtained slurry product was filtered through a filter to obtain a cake-like slurry.
[0085] The cake-like slurry was transferred to a vessel equipped with a reflux condenser, and then 300 ml of distilled water and 3 g of a 27% ammonia water solution were added and the mixture was heated and stirred at 70°C for 24 hours. Next, the slurry was placed in a kneader and heated to above 80°C, and then kneading was performed while removing the moisture to obtain a clay-like kneaded blend. The obtained kneaded blend was extruded into a 1.5 mm diameter cylinder form using an extruder, and after drying at 110°C for 1 hour, it was fired at 550°C to obtain a molded carrier.
[0086] Separately, molybdenum trioxide, nickel nitrate hexahydrate and phosphoric acid (85% concentration) were added to 150 ml of distilled water, and malic acid was added until the components thoroughly dissolved to prepare an impregnating solution. A 300 g portion of the molded carrier was impregnated with the impregnating solution by spraying. The amounts of molybdenum trioxide, nickel sulfate hexahydrate and phosphoric acid used were adjusted for the prescribed loading weights (contents listed in Table 1) as molybdenum, nickel and phosphorus oxides (M0O3, NiO, P2O5).
10087] The sample obtained by impregnation in this manner was dried at 110°C for 1 hour and then fired at 550°C in air to obtain a first hydrogenation catalyst A. The physical properties of the prepared first

hydrogenation catalyst A are shown in Table 1.
[0088] [Table 1]
(Table Removed)

[0089] Next, zeolite Y with a silica/alumina molar ratio of 5 was stabilized by an ultrastabilizing method. The zeolite was subjected to acid treatment with a IN nitric acid aqueous solution, to obtain protonated ultrastabilized zeolite Y The ultrastabilized zeolite Y had a unit lattice length of 24.33 Ǻ, a silica/alumina molar ratio of 30 and a volume of 30-100 A diameter pores that was 0.055 mL/g with respect to the zeolite weight, as measured by mercury porosimetry. [0090] The obtained ultrastabilized zeolite Y (550 g) was then added to 3 L of an ammonium nitrate aqueous solution (2N concentration) and stirred at room temperature for conversion to the ammonium form. [0091] Next 1 kg of a 5 wt% sodium aluminate aqueous solution was mixed with water glass No.3 to prepare a clay-like kneaded blend by the same method as for the aforementioned first hydrogenation catalyst carrier. The obtained kneaded blend was then extruded into a 1.5 mm diameter cylinder form using an extruder, and after drying at 110°C for I hour, it was fired at 550°C to obtain a molded carrier comprising 55 wt% zeolite. [0092] Separately, ammonium paratungstate and nickel nitrate

hexahydrate were dissolved in 150 ml of distilled water to obtain an impregnating solution. A 300 g portion of the zeolite-containing molded carrier was impregnated with the impregnating solution by spraying to obtain hydrocracking catalyst B. The amounts of ammonium paratungstate and nickel nitrate hexahydrate used were adjusted for the prescribed loading weights (contents listed in Table 1) as tungsten and nickel oxides (WO3, NiO). The physical properties of the prepared hydrocracking catalyst B are shown in Table 1. [0093] (Example 2)
A first reaction tube (20 mm inner diameter) packed with the first hydrogenation catalyst A (70 mL) and a second reaction tube (20 mm inner diameter) packed with the hydrocracking catalyst B (30 mL) were installed in series in that order in a fixed bed circulating reactor. As the first stock oil there was prepared a blend of palm oil (15°C density =z 0.916 g/mL, oxygen content =11.4 wt%, 10% running point - 588°C) and a Middle East vacuum light oil fraction (15°C density = 0.919 g/ml, sulfur content = 2.41 wt%, nitrogen content ==610 ppm by weight, 10% running point = 344°C) in a volume ratio of 20:80. Next, straight-run light oil (3 wt% sulfur content) containing added dimethyl disulfide was used for pre-sulfurization of the catalyst for 4 hours under conditions with a catalyst A and B catalyst layer mean temperature of 300°C, a hydrogen partial pressure of 6 MPa, a LHSV of 1 h"1 and a hydrogen/oil ratio 200 NL/L. Upon completion of the pre-sulfurization, the first stock oil was passed through the reactor under conditions with a catalyst A and B reaction temperature of 400°C, a hydrogen pressure of 10.5 MPa and a LHSV of 0.7 h" for hydrogenating pre-processing and

hydrocracking. The 80-145°C boiling point range fraction (first naphtha) was separated off with an apparatus conforming to the "Distillation Test Method Using 15-Theoretical Plate Rectification Column" in JIS-K2601, "Crude Oil Test Methods". The properties of the obtained first naphtha were paraffins: 98 vol%, naphthenes; 2 vol%, aromatics: 0 vol%, sulfur content: <1 ppm by weight, oxygen content: 45 ppm by weight. [0094] (Example 3)
A reaction tube (20 mm inner diameter), packed with 70 mL of a commercially available catalyst (specific surface area: 180 m2/g) comprising 3 wt% nickel as oxide (NiO) with respect to the catalyst weight and 18 wt% molybdenum as oxide (M0O3) with respect to the catalyst weight loaded onto an alumina carrier as the second hydrogenation catalyst, was installed in a fixed bed circulating reactor. As the naphtha fraction there was prepared a blend of straight-run naphtha (initial boiling point: 90°C, end point: 155°C, paraffins: 65 vol%, naphthenes: 25 vol%, aromatics: 10 vol%, sulfur content: 330 ppm by weight) and the first naphtha obtained in Example 2 in a volume ratio of 70:30. Next, straight-run light oil (3 wt% sulfur content) containing added dimethyl disulfide was used for pre-sulfurization of the catalyst for 4 hours under conditions with a catalyst layer mean temperature of 300°C, a hydrogen partial pressure of 3 MPa, a LHSV of 1 h'1 and a hydrogen/oil ratio 200 NL/L. After the pre-sulfurization, the naphtha fraction was passed through the reactor under conditions with a reaction temperature of 310°C, a pressure of 2.5 MPa, a LHSV of 5 h" and a hydrogen/oil ratio of 40 NL/L for a second hydrotreatment.

The sulfur content of the product oil (second naphtha) was 0.2 ppm by weight, and the oxygen content was less than 5 ppm by weight. [0095] (Example 4)
To 500 g of globular γ-alumina there was added 500 mL of distilled water. There were further added thereto 500 mL of a 0.012 mol/L platinic chloride aqueous solution and 500 mL of a 0.020 mol/L aqueous solution comprising a 0.1N hydrochloric acid solution added to stannic chloride. The water was then evaporated off with an evaporator. The obtained sample was dried at 120°C for 10 hours and then fired at 400°C for one hour to obtain a catalytic reforming catalyst. The specific surface area of the obtained catalyst was measured as 195 m /g by the nitrogen adsorption method. [0096] (Example 5)
A reaction tube (20 mm inner diameter) packed with the catalytic reforming catalyst (30 mL) obtained in Example 4 was installed in a fixed bed circulating reactor. The catalyst was then subjected to reduction pre-processing under conditions of 2 MPa hydrogen partial pressure, 530°C for reaction pre-processing. The second naphtha obtained in Example 3 was passed through the reactor while adjusting the moisture content for catalytic reforming at a reaction temperature of 530°C. The properties of the product are shown in Table 2. In Table 2, the hydrogen product yield is expressed as NL with respect to 1 L of the first stock oil. The benzene, toluene and xylene yields were calculated from values measured by the methods described in JIS-K2536-2, "Petroleum Products - Component Test Methods (Determining Total Components by Gas Chromatography)". The

research octane number is the research method octane number measured according to JIS-K2280, "Octane Number and Cetane Number Test Method". [0097] [Table 2]
(Table Removed)

[0098] (Comparative Example 1)
Product oil was obtained in the same manner as Example 3, except that the straight-run naphtha alone was passed through for the second hydrotreatment without using the first naphtha. The sulfur content of the obtained product oil was 0.2 ppm by weight, and the oxygen content was less than 0.1 ppm by weight. [0099] (Comparative Example 2)
A reaction tube (20 mm inner diameter) packed with the catalytic reforming catalyst (30 mL) obtained in Example 4 was installed in a fixed bed circulating reactor. The catalyst was then subjected to reduction pre-processing under conditions of 2 MPa hydrogen partial pressure, 530°C for reaction pre-processing. Next, the product oil obtained in Comparative Example 1 was passed through the reactor while adjusting the moisture content for catalytic reforming at a reaction temperature of 505°C. The properties of the product are shown in fable 2.
[0100] By adjusting the operating conditions when the first stock oil containing an animal or vegetable oil-derived fat or oil component is subjected to a hydrocracking step and catalytic reforming step as shown above, it is possible to maintain a reformulated gasoline yield in the obtained product which is basically equivalent to that achieved when using a stock oil containing no animal or vegetable oil-derived fat or oil component, while also increasing the hydrogen product yield. [0101] According to the invention it is possible to provide a process for producing hydrogen which is adequately effective for LCA-C02 reduction.

CLAIMS
1. A process for producing hydrogen comprising
a hydrocracking step in which a first stock oil or a second stock oil derived from the first stock oil is subjected to hydrocracking by contact with a hydrocracking catalyst in a hydrogen-containing atmosphere to obtain a first naphtha, and
a catalytic reforming step in which the first naphtha or a second naphtha derived from the first naphtha is subjected to catalytic reforming to obtain hydrogen,
wherein the first stock oil contains an animal or vegetable-derived fat or oil component, the animal or vegetable oil-derived fat or oil component including fractions with boiling points of 230°C or higher,
the hydrocracking catalyst comprises at least one metal selected from the group consisting of metals belonging to Group 6A and Group 8 of the Periodic Table and an acidic inorganic oxide, and
the first naphtha includes a fraction with boiling point of 100-120°C.
2. A process for producing hydrogen according to claim 1,
which further comprises a hydrotreatment step in which, prior to
the catalytic reforming step, the naphtha fraction containing the first naphtha is contacted with a second hydrogenation catalyst in a hydrogen-containing atmosphere, to obtain the second naphtha with a lower sulfur content and oxygen content than the naphtha fraction,
wherein the second hydrogenation catalyst comprises at least one metal selected from the group consisting of metals belonging to

Group 6 A and Group 8 of the Periodic Table, and
the second naphtha has a sulfur content of no greater than 1.0 ppm by weight and an oxygen content of no greater than 10 ppm by weight.
3. A process for producing hydrogen according to claim 1 or 2,
which further comprises a hydrogenating pre-processing step in
which, prior to the hydrocracking step, the first stock oil is contacted with the first hydrogenation catalyst in a hydrogen-containing atmosphere to obtain a second stock oil,
wherein the first hydrogenation catalyst comprises at least one metal selected from the group consisting of metals belonging to Group 6A and Group 8 of the Periodic Table.
4. A process for producing hydrogen according to any one of
claims 1 to 3, wherein the conditions for hydrocracking in the
hydrocracking step are a hydrogen pressure of 6-20 MPa, a LHSV of
0.2-1.5 h-1 and a hydrogen/oil ratio of 200-2000 NL/L.
5. A process for producing reformulated gasoline wherein
reformulated gasoline is obtained by a catalytic reforming step
according to any one of claims 1 to 4.
6. A process for producing aromatic hydrocarbons wherein a C6-
8 aromatic hydrocarbon is obtained by a catalytic reforming step
according to any one of claims 1 to 4.