|Número de publicación||WO2007031713 A1|
|Tipo de publicación||Solicitud|
|Número de solicitud||PCT/GB2006/003305|
|Fecha de publicación||22 Mar 2007|
|Fecha de presentación||7 Sep 2006|
|Fecha de prioridad||14 Sep 2005|
|Número de publicación||PCT/2006/3305, PCT/GB/2006/003305, PCT/GB/2006/03305, PCT/GB/6/003305, PCT/GB/6/03305, PCT/GB2006/003305, PCT/GB2006/03305, PCT/GB2006003305, PCT/GB200603305, PCT/GB6/003305, PCT/GB6/03305, PCT/GB6003305, PCT/GB603305, WO 2007/031713 A1, WO 2007031713 A1, WO 2007031713A1, WO-A1-2007031713, WO2007/031713A1, WO2007031713 A1, WO2007031713A1|
|Inventores||Yazhong Chen, Yuzhong Wang, Hengyong Xu|
|Solicitante||Bp P.L.C., Dalian Institute Of Chemical Physics|
|Exportar cita||BiBTeX, EndNote, RefMan|
|Citas de patentes (7), Otras citas (1), Citada por (9), Clasificaciones (35), Eventos legales (4)|
|Enlaces externos: Patentscope, Espacenet|
PROCESS FOR HYDROGEN PRODUCTION
This invention relates to the production of hydrogen, more specifically to the conversion of hydrocarbons into hydrogen in a reactor comprising a palladium membrane. There is increasing concern that man-made emissions of greenhouse gases due to the burning of fossil fuels may be contributing to increasing global temperatures through a "greenhouse effect". This is prompting international activity in finding and developing sustainable and clean sources of energy.
Hydrogen is increasingly being studied as an energy source, as it produces a large amount of energy per unit mass, and emits only water when combusted. In addition, hydrogen can be used in small-scale energy devices, such as proton membrane fuel cells, which combine the benefits of energy efficiency and clean emissions with flexibility.
Activity has also been directed towards extracting hydrogen from hydrocarbon fossil fuels, through processes such as steam reforming, and capturing the resulting CO2 for sequestration in order to make the fuel effectively greenhouse gas neutral. The hydrogen produced by such processes can ultimately be used as an energy source for a number of applications, for example as a fuel for a fuel cell-powered vehicle, or as a combustible fuel in a power station.
Existing processes for hydrogen production include, for example, the steam reforming or partial oxidation of natural gas. The cost of hydrogen produced by such means at large scales is typically low. However, for smaller scale applications, such as for a hydrogen vehicle filling station or an electricity distribution sub-station, the fixed costs of equivalent small-scale steam reforming units can result in the hydrogen produced being prohibitively expensive. The steam reforming of natural gas is a reaction which is limited by thermodynamic equilibrium, and in order to obtain relatively high natural gas conversions, the reaction temperature is typically in excess of 800°C, which is expensive in terms of energy usage, and places heavy demands on the reactor equipment, often necessitating the use of expensive reactor materials. A great deal of research has been done on the steam reforming reaction of natural gas in a palladium membrane reactor, such that the hydrogen produced permeates through the selectively permeable palladium membrane, which shifts the chemical equilibrium towards the producing further hydrogen. This allows lower temperature operation, and can result in improved hydrogen yields. Thus, for example, Shigeyuki Uemiya et al (Steam Reforming of Methane in a Hydrogen-Permeable Membrane Reactor, Appl. Catal. 67, 223, 1991), report the use of a Ni-based catalyst in a palladium membrane reactor at the relatively low temperature of 623-773 K to realise the steam reforming reaction of methane. Also, US 5,525,322 describes the recovery of hydrogen from water and hydrocarbons and an isotope method for the same, using a palladium membrane and nickel-based catalyst. By separating out the hydrogen produced, the following equilibria can be shifted to the right; CO + H2O → CO2 + H2, CH4 + H2O → CO + 3H2, and CH4 → C + 2H2, thus realising high conversions and hydrogen yields at relatively low starting material space velocity. However, under conditions of a relatively high space velocity, there is a marked fall in the hydrogen recovery rate, severely limiting the hydrogen-producing capacity of the device.
Higher hydrocarbons, for example liquid fuels such as liquefied petroleum gas (LPG), gasoline, diesel, kerosene and the like, have a high energy density and are easily transported and stored, and can potentially be used in low-scale production of hydrogen. Usually, the steam reforming of such hydrocarbons or the catalytic partial oxidation needs to be carried out at temperatures in excess of 800°C to obtain acceptable hydrocarbon conversions and hydrogen yields. Furthermore, a common product when higher hydrocarbons are used is methane, which needs to be separated and recycled in order to improve the hydrogen yields. Therefore, there remains a need for a process to convert higher hydrocarbons into hydrogen at lower temperatures, and with reduced production of methane.
According to the present invention there is provided a process for producing hydrogen from a hydrocarbon in a reactor having a first and a second zone separated by a selective hydrogen-permeable membrane, which process comprises contacting the hydrocarbon with water in the first zone of the reactor, in the presence of a catalyst and under conditions that allow a reaction to occur to produce hydrogen, which hydrogen permeates through the selective hydrogen-permeable membrane to the second zone of.the reactor, characterised in that the gas hourly space velocity of hydrocarbon over the catalyst, when multiplied by the number of carbon atoms in the hydrocarbon, is maintained at a value of greater than 2000 h"1. In the present invention a hydrocarbon is reacted with steam in the first zone of the reactor, optionally in the presence of oxygen. The reactor has a selective hydrogen- permeable membrane that separates the first zone from a second zone. Hydrogen produced in the reaction permeates through "the membrane, reducing the hydrogen partial pressure in the first zone and hence shifting the equilibrium therein to allow further production of hydrogen. By such means, reduced reaction temperatures are necessary, typically below temperatures of 8000C often employed for steam reforming reactions.
The molar ratio of water (or steam) to carbon, wherein "carbon" relates to the total number of carbon atoms provided by the hydrocarbon, is selected so that sufficient water is present to minimise catalyst coking, while simultaneously minimising any excess to maintain energy efficiency. The molar steam to carbon ratio is typically in the range of from 50 : 1 to 1 : 50, such as in the range of from 10 : 1 to 1 : 10. Preferably, the steam to "carbon" molar ratio is in the range of 1.8 : 1 to 4 : 1. As an example, if decane was used as the hydrocarbon, and a molar ratio of steam to decane of 20 : 1 was used, this would equate to a steam to carbon molar ratio of 2 : 1.
In one embodiment of the invention, oxygen is also fed to the first zone. This allows a combined steam reforming and partial oxidation reaction to occur therein. Heat released by the exothermic partial oxidation reaction can reduce the amount of heat that need to be otherwise supplied to the reactor to compensate for that absorbed during the endothermic steam reforming reaction, and hence improves the energy efficiency of the process. Both partial oxidation and steam reforming reactions result in the formation of hydrogen and oxides of carbon (COx) from the hydrocarbon.
Typically, the temperature in the first zone of the reactor is below 8000C, such as in the range from 300 to 6000C. Pressures in the first zone of the reactor are typically maintained so as to maximise hydrogen yields, which are improved at lower pressures, while maintaining a sufficient hydrogen pressure so that a sufficient gradient exists across the selective hydrogen-permeable membrane. The pressure is typically up to 10 MPa, preferably 0.5 MPa or more, such as in the range of from 1 to 10 MPa.
By maintaining a high gas hourly space velocity (GHSV) of hydrocarbon, when multiplied by the number of carbon atoms in the hydrocarbon, at a value of greater than 2000 h"1, the conversion of oxides of carbon into methane in the presence of hydrogen is reduced, and better hydrogen yields are maintained. Henceforth, the GHSV of the hydrocarbon, when multiplied by the number of carbon atoms, will be shortened to "C- GHSV" for convenience. The GHSV is calculated by taking the volume of hydrocarbon fed to the reactor per hour, correcting the volume to that of a gas at standard temperature and pressure, i.e. at O0C and 1 aim pressure, and dividing by the volume of the catalyst in the reactor (or the relevant zone of the reactor). To obtain the C-GHSV, this valμe is then simply multiplied by the number of carbon atoms in the hydrocarbon. Preferably, the C- GHSV of hydrocarbon is 5 000 h'1 or more, more preferably is 10 000 h"1 or more, and yet more preferably is 15 000 h"1 or more.
The process of the present invention is particularly suitable for producing hydrogen from hydrocarbons having two or more carbon atoms, such as those found in or derived from liquefied hydrocarbons such as natural gas liquids (NGL) or liquefied petroleum gases (LPG), or the hydrocarbon can comprise one or more liquid hydrocarbons with in the range of, for example, from 6 to 30 carbon atoms, such as those present in naphtha, middle distillate, kerosene or atmospheric gas oil fractions of a refinery crude distillation unit, or products derived from the fractions such as gasoline, diesel, aviation fuel, heating oil and the like. In the prior art, hydrocarbons with two or more carbon atoms have typically been removed from a natural gas feed before being fed to a membrane-containing reactor, as the the higher hydrocarbons tend to cause fouling or coking of the reforming catalyst. Therefore, by maintaining the carbon space velocity (C-GHS V) at values of greater than 2000 h"1, such problems can be reduced, while still maintaining high conversions. In a preferred embodiment, the hydrocarbon is a hydrocarbon having in the range of from 4 to 16 carbon atoms, or a mixture of hydrocarbons having in the range of from 4 to 16 carbon atoms.
The reactor has two zones which are separated by a selective hydrogen-permeable membrane. Hydrogen is produced from the hydrocarbon in the first zone, and the hydrogen permeates through the membrane into the second zone of the reactor. By constantly removing hydrogen selectively from the first zone of the reactor, the extent of methanation is reduced, and the yield of hydrogen is improved.
Hydrogen permeates the membrane in response to a concentration gradient, such that it flows from a region of high hydrogen concentration or partial pressure to a region of low hydrogen concentration of partial pressure. Thus, to remove hydrogen from the first zone of the reactor where it is generated requires a lower partial pressure to be maintained in the second zone of the reactor. This is achieved by maintaining a reduced total pressure in the second zone of the reactor and/or by diluting any hydrogen that permeates, for example with a gas that is inert to hydrogen under conditions in the second zone such as one or more of nitrogen, steam or argon. The diluent gas can also be used to flush hydrogen from the second zone of the reactor. Typically, the membrane is a palladium membrane, optionally in the presence of other elements such as Ag or Cu. Preferably, the hydrogen permeation rate through the membrane is greater than 50 m3 m"2 h'1 bar"1. A membrane having a H2/N2 molar separation coefficient of greater than 10 000 is preferred.
The catalyst of the present invention is capable of catalysing the conversion of the hydrocarbon to oxides of carbon (COx) and hydrogen under the low temperature conditions in the first zone of the reactor. Typically, the catalyst will comprise one or more of the metals nickel, ruthenium, platinum, palladium, rhodium, rhenium and iridium.
The catalyst can be supported on an inert support, typically a refractory oxide. The refractory oxide is suitably selected from one or more of magnesia, alumina, silica and zirconia. The catalyst may also comprise additional components that can promote activity, or improve selectivity, such as alkaline earth elements and lanthanide elements.
In a preferred embodiment, nickel is used due to its low cost, and is supported on an alumina support, preferably with a nickel loading in the range of from 25 to 70% by weight, based on NiO, and is preferably in the range of from 30 to 50% by weight, based on NiO. In a further embodiment, the catalyst additionally comprises magnesium and lanthanum as promoters. The catalyst can be prepared by deposition techniques, where the catalyst and/or any promoters are deposited on the oxide support, or using co-precipitation techniques wherein a precipitate comprising the various constituents of the catalyst is obtained from a solution comprising soluble compounds of the various constituents, typically achieved by adding a base such as an alkali metal hydroxide, an alkaline earth metal hydroxide or'ammonium hydroxide.
The invention will now be illustrated in the following examples, and with reference to Figure 1 which schematically illustrates the steam reforming of a hydrocarbon in a reactor comprising a selective hydrogen-permeable palladium membrane. Figure 1 shows a plan view of a steam reforming process in accordance with the present invention. A hydrocarbon (such as iso-octane) and steam reactants 1 are fed into the first zone 2 of reactor 3, which reactor comprises a closed end ceramic alumina tube 4 supporting a palladium membrane 5 that separates the first zone from a second zone 6 on the permeate side of the palladium membrane. The reactants pass over a steam reforming catalyst 7 in the first zone of the reactor, where the iso-octane and steam react to form COx and hydrogen. Hydrogen permeates through the palladium membrane, while any unreacted starting materials, the COx, and any methane formed through methanation leave the first zone of the reactor through outlet 8. Hydrogen is removed from the second zone of the reactor through outlet 9. In the following examples, the methane selectivity and hydrogen selectivity are defined as follows:
Methane selectivity = 100 x [Product methane flow rate x 2 / (product methane flow rate x 2 + product hydrogen flow rate)]
Hydrogen selectivity = 100 x [Product hydrogen flow rate / (product methane flow rate x 2 + product hydrogen flow rate)]
Example 1: A catalyst was prepared by adding (MLi)2CO3 (as a precipitating agent) to an aqueous solution of Ni(NO3)2.6H2O and A1(NO3)3.9H2O. A pH of 8.0 - 8.5 was maintained, and the precipitate was aged in the supernatant solution for 2 h at room temperature. The precipitate was separated, washed with water, and dried at 12O0C before being calcined at 600°C for 4 h. The NiO content of the catalyst was 40 wt%, and the Al2O3 content was 60 wt%.
5.5 g of the catalyst was diluted with quartz sand and placed in the first zone of the palladium membrane reactor with a bed layer height of 100 mm. The hydrogen permeation rate of the membrane was 52 m3 m"2 bar"1 h"1, and the H2/N2 separation coefficient was over 10,000. A Daqing 6 # solvent oil, comprising a mixture of C5 to C7 hydrocarbons, was then fed to the catalyst-containing zone of the reactor at a rate of 0.34 ml/min of liquid, and liquid water was also fed at a rate of 0.80 ml/min. The reaction was conducted at 5500C and 1.1 MPa, and an Ar flushing gas was injected into the second (hydrogen permeate) zone of the reactor. The molar feed rate of argon fed to the second reactor in unit time was 1.8 times that of the molar rate of feeding carbon atoms to the first reactor zone.
The results showed that the product methane selectivity was 2.8%, and the hydrogen selectivity was 97.2%. The purity of the hydrogen having passed through the membrane (not accounting for the Ar flushing gas) was greater than 99.9%.
A catalyst was prepared by adding ammonium hydroxide (as a precipitating agent) to an aqueous solution of Ni(NO3)2.6H2O and A1(NO3)3.9H2O and La(NO3)3.6H2O at room temperature. A pH of 8.0 was maintained, and the precipitate was aged in the supernatant solution for 1 h at room temperature. The precipitate was separated, washed with water, and dried at 120°C before being calcined at 600°C for 4 h. The catalyst obtained had a composition of 50 wt% NiO, 42 wt% Al2O3 and 8 wt% La2O3. Catalyst performance was measured in a conventional fixed bed reactor with isp- octane as the starting material at a pressure of 0.8 MPa5 and at temperatures of 450°C, 500°C and 550°C respectively. A steam to carbon ratio of 2.73 was used in the reactions. The results are shown in Table 1.
Iso-octane conversion was almost 100%, even under the relatively mild steam reforming conditions used, and at high C-GHSV values of 30,000 h"1. The product (on a dry basis, i.e. without accounting for water) approached the values expected for thermodynamic equilibrium. Methane production at higher temperatures was greater than at lower temperatures.
5.5 g of the above-mentioned catalyst, with particle sizes of 420-630 μm, was then . placed in a reactor comprising a palladium membrane and diluted with quartz sand, the height of the combined catalyst and diluent bed being 100mm. The permeation rate of hydrogen across the palladium membrane was 58 m3 m"2 bar"1 h"1 and the H2/N2 separation coefficient was greater than 10 000. Liquid iso-octane was fed at a rate of 0.34 ml/min and liquid water was fed at a rate of 0.80 ml/min to the first (catalyst containing) zone of the reactor. The reaction was conducted at 550°C and 1.3 MPa. The molar feed rate ratio of Ar flushing gas fed to the second (hydrogen permeate) zone of the reactor to the carbon fed to the first zone of the reactor was 1.8. Selectivity to methane was 2.3%, and the hydrogen selectivity was 97.7%. The purity of the hydrogen having passed through the membrane (not accounting for the Ar flushing gas) was greater than 99.9%.
A catalyst was prepared by adding (NH4)2CO3 (as a precipitating agent) to an aqueous solution of Ni(NO3)2.6H2O and A1(NO3)3.9H2O and Mg(NO3)2.6H2O at room temperature. A pH of 8.0 - 8.5 was maintained, and the precipitate was aged in the supernatant solution for 2 h at room temperature. The precipitate was separated, washed with water, and dried at 120°C before being calcined at 700°C for 4 h. The catalyst obtained had a composition of 36 wt% NiO, 12 wt% MgO and 52 wt% Al2O3.
The prepared catalyst was pressed and crushed into particle sizes having a diameter of 420-630 μm. 5.5 g of the catalyst were diluted with quartz sand and placed in the first zone of the palladium membrane reactor to a bed layer height of 100 mm. The hydrogen permeation rate of the membrane was 65 m m" bar" h" , and the H2/N2 separation coefficient was greater than 10 000. With iso-octane as the hydrocarbon feed, the reaction was conducted with a C-GHSV of 2000 h'1, a reaction temperature 550°C, a H2O/C ratio 2.73 and a pressure 1.7 MPa. The molar feed rate ratio of Argon flush gas fed to the second reactor zone to carbon fed to the first reaction zone was 1.8. The product methane selectivity was 1.8%, and the hydrogen selectivity was
98.2%. The purity of the hydrogen having passed through the membrane (not accounting for the argon flushing gas) was greater than 99.9%.
Example 4: 5.5 g of the Ni/MgO-Al2O3 catalyst from Example 3 was used and placed in the first zone of a palladium membrane reactor. The permeation rate of the membrane was 68
1X 0 1 1 m m" bar" h" , and the H2/N2 separation coefficient was greater than 10 000. Iso-octane was fed to the reactor at a liquid feed rate of 0.17 ml/min, and a liquid water feed rate of 0.40 ml/min. The molar feed rate ratio of argon flushing gas fed to the second reactor zone compared to carbon fed to the first reactor zone was 1.8. The results are shown in table 2. Methane selectivity was 2.6%, and hydrogen selectivity was 97.4%. Comparative Example 5:
5.5 g of the Ni/MgO- Al2O3 catalyst from Example 3 was used and placed in a conventional stainless steel reactor (with no selective hydrogen-permeable membrane) with iso-octane as the starting material. The feed rate of the liquid iso-octane was 0.17 ml/min, and the liquid water feed rate was 0.40 ml/min. The reaction was conducted at 55O0C and 0.9 MPa. The dry base composition of the product formed is shown in Table 3.
Even though the iso-octane conversion was almost 100%, the product selectivity of methane was 62%, while the hydrogen selectivity was 38%.
Gas hourly space velocity of the hydrocarbon feed multiplied by the number of carbon atoms in the hydrocarbon.
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|Clasificación internacional||C01B3/38, C01B3/50|
|Clasificación cooperativa||Y02P20/52, Y02P30/30, B01J8/025, B01J8/0257, B01J23/755, C01B2203/048, C01B2203/0244, C01B2203/1094, C01B2203/1058, C01B2203/041, C01B2203/1247, C01B3/505, B01J37/03, B01J23/78, B01J19/2475, C01B2203/1082, C01B3/38, C01B2203/0233, B01J8/009, C01B2203/86, B01J8/0278, B01J23/83|
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