Title of Invention

"A PROCESS FOR THE CONTINUOUS PRODUCTION OF CARBON MONOXIDE-FREE HYDROGEN FROM METHANE OR METHANE-RICH HYDROCARBONS"

Abstract

A process for the continuous production of carbon monoxide-free hydrogen from methane or methane-rich hydrocarbons This invention relates to a process for the continuous production of carbon monoxide-free hydrogen from methane or methane-rich hydrocarbons using group VIII metal(s) containing solid catalyst in two parallel reactors. This invention particularly relates to a process for the continuous production of carbon monoxide-free hydrogen from methane or methane-rich hydrocarbons using group VIII metal(s) containing solid catalyst in two parallel rectors operated in a cyclic manner fro two different reactions, one the decomposition of methane or methane-rich hydrocarbons to hydrogen and carbon, which is deposited on the catalyst, and second the gasification of the carbon deposited on the catalyst by steam in the presence or absence of oxygen, carried out simultaneously in the two reactors, and collecting the product stream of each of the two reactions separately.

Full Text

FIELD OF THE INVENTION 1 This invention relates to a process for the continuous production of carbon monoxide-free hydrogen from methane or methane-rich hydrocarbons using group VIII metal(s) containing solid catalyst in two parallel reactors. ,This invention particularly relates to a process for the continuous production of carbon monoxide-free hydrogen from methane or methane-rich hydrocarbons using group VIII metal(s) containing solid catalyst in two parallel reactors operated in a cyclic manner for two different reactions, one the decomposition of methane or methane-rich hydrocarbons to hydrogen and carbon, which is deposited on the catalyst, and second the gasification of the carbon deposited on the catalyst by steam in the presence or absence of oxygen, carried out simultaneously in the two reactors, and collecting the product stream of each of the two reactions separately. The process of this invention could be used in the petroleum refining and chemical industries for the production of carbon monoxide-free hydrogen required for proton exchange membrane fuel cells and also for the various hydro-treating processes in the petroleum refining and hydrogenation processes in chemical industries. The demand for hydrogen has been increasing day-by-day for the hydro treating processes in petroleum industries and also for hydrogen fuel cells, both stationery and non-stationery fuel cells. Since hydrogen is a non-polluting fuel, its use as a fuel, particularly for fuel cells has been increasing very fast. However, the well-established proton exchanged membrane fuel cells require carbon monoxide-free hydrogen as a fuel to avoid deactivation of the noble metal catalyst in the fuel cells. The main natural sources of hydrogen are hydrocarbons and water. Among the hydrocarbons methane has the highest hydrogen to carbon ratio and hence it is the most preferred choice among the hydrocarbons for hydrogen production. The conventional processes for the production of hydrogen are based on steam reforming of hydrocarbons, such as naphtha and methane or natural gas and auto thermal reforming of hydrocarbons, particularly heavier hydrocarbons. The hydrogen production processes have been recently reviewed by Fierro and coworkers [reference: Pena, M. A., Gomez, J. P. and Fierro, J. L. G.; Applied Catalysis A: General; volume 144, page 7-57, year 1996]. Both the hydrocarbon steam reforming and auto thermal reforming processes are operated at high temperatures, above about 900°C and the product stream of these processes contains appreciable amounts of carbon monoxide along with hydrogen. The removal of carbon monoxide at low concentrations from hydrogen is very costly. Hence, the hydrocarbon steam reforming and auto thermal reforming processes are not economical for the production of carbon monoxide-free hydrogen. Hence, there is a practical need to develop a process for the production of hydrogen form methane as it has highest hydrogen to carbon ratio among the hydrocarbons, at temperatures lower than that used in the conventional hydrocarbon steam reforming and auto thermal reforming processes. Production of Carbon Monoxide-Free Hydrogen form Methane A few processes are known also for the production of carbon monoxide-free hydrogen from methane. Recently, Kikuchi [reference: Kikuchi, E., Hydrogen-permselective membrane reactors, CATTECH, March 1997, page 67-74, Baltzer Science Publishers] has described a process based on steam reforming of methane in membrane reactor to produce hydrogen free of carbon monoxide. By applying Pd/ceramic composite membrane to steam reforming of methane over a commercial supported nickel catalyst, methane conversion upto 100% can be accomplished in Pd-membrane reactor at temperatures as low as 500°C to produce carbon monoxide-free hydrogen. In this process, the hydrogen produced in the steam reforming of methane is continuously removed from the reaction system by the selective permeation of hydrogen through the Pd-membrane. However this process has not yet been commercialized and it has following drawbacks/limitations: 1) Because of the use of a number of Pd-membrane tubes, the capital cost of this process is very high. 2) There is a possibility that the Pd-membrane becomes deactivated by deposition of carbonaceous matter. 3) There is also a problem of membrane stability and/or a possibility of membrane failure due to formation of pinholes in the membrane. A Japanese patent [Jpn. Kokai Tokkyo Koho JP 09234372 A2, September 09, 1997] disclosed a process for the manufacture of hydrogen by thermal decomposition of hydrocarbons at 200°C-1000°C using a catalyst containing nickel, alkali or alkaline earth compounds. A Russian patent [Russ. RU 2071932 Cl January 20, 1997] disclosed the production of hydrogen and carbon by thermal decomposition of methane on nickel catalyst. A recent Japanese patent [Jpn. Kokai Tokkyo Koho JP 11228102 A2, August 24, 1999] disclosed reactors for thermal decomposition of methane to form carbon and hydrogen. Hydrogen production by catalytic cracking of methane or natural gas and other hydrocarbons, at below 900°C using nickel-based catalyst is disclosed in a few publications [reference: Zhang, T and Amiridis, M. D., Applied Catalysis A: General, volume 167, page 161-172, year 1998; Muradov, N. Z., Energy Fuels, volume 12, page 41-48, year 1998; Kuvshinov, G. G. et. al. Hydrogen Energy Progress XI Proceedings of World Hydrogen Energy Conference, 11th, Volume 1, page 655-660, Edited by Veziroglu, T., year 1996; Muradov, N. Z., Proceedings of US DOE Hydrogen Program Review, volume 1, page 513-535, year 1996]. In the above prior art processes based on catalytic cracking or thermo-catalytic decomposition of methane or other hydrocarbons, the hydrogen produced is free from carbon monoxide and carbon dioxide but the catalyst deactivation is fast due to the carbon formed on the catalyst and this is accompanied with a fast increase in the pressure drop across the catalyst bed, making the process unpractical for the hydrogen production. Recently, Choudhary and Goodman reported a process for the production of carbon monoxide-free hydrogen involving stepwise methane steam reforming [reference: Choudhary, T. V. and Goodman, D. W., Catalysis Letter, volume 59, page 93-94, year 1999]. In this process, methane pulse and water pulses are passed over a pre-reduced nickel-based catalyst at 375°C, alternatively. When methane pulse is passed over the catalyst, the methane from the pulse is decomposed to hydrogen and carbon, leaving the carbon deposited on the catalyst according to the reaction: CH4 ->C + 2H2t (1) When the water pulse is passed over the catalyst with the carbon deposited on it, the carbon on the catalyst reacts with steam to form CO2, hydrogen and methane according to the reaction: C + 2H2O -»CO2 + 2H2 (2) In this process although the carbon monoxide-free hydrogen is produced by catalytic cracking of methane and the carbon deposited on the catalyst is removed by the cyclic operation of the methane and water pulses in the same reactor, the process is not operated in the steady state and the hydrogen produced is not continuous one. Hence, it is not practical and also not economical to produce carbon monoxide-free hydrogen on large scale by this transient process involving cyclic operation of methane and water pulses. Very recently, Choudhary et al have reported a possibility of the continuous production of hydrogen at 500°C by carrying out the above two reactions, Reactions 1 and 2, simultaneously in two parallel catalytic reactors in cyclic manner by switching a methane containing feed, 18.2 mol% CH4 in NI, and a steam containing feed, 20.5% steam in Na, between the two reactors at predecided intervals of time and combining the product streams of the two reactors (reference: V.R. Choudhary, S. Banerjee and A. M .Rajput, Journal of Catalysis, vol. 198 page 136 and year 2001). However, both the reactions, Reactions 1 and 2 are thermodynamically favored at higher temperatures. The methane decomposition, Reaction 1, is also favored at lower pressure or lower concentration of methane. Our preliminary studies showed that both the methane conversion in Reaction 1 and degree of carbon gasification in Reaction 2 are decreased sharply with increasing the methane concentration and for decreasing the temperature. Hence, for using undiluted or less diluted methane as a feed so that the very high cost of separation of the diluent can be reduced and also for obtaining high conversion of methane, Reactions 1 and 2 need to be carried out at higher temperature, above about 600°C. However at such a high temperature a significant amount of carbon monoxide is formed in Reaction 2 and therefore, carbon monoxide-free the above cyclic process cannot obtain hydrogen. Because of the above mentioned drawbacks and limitations of all the prior art processes, there is a great need for developing a process for the continuous production of carbon monoxide-free hydrogen by catalytic decomposition of methane or natural gas at a temperature below about 900°C, while avoiding the carbon build up on the catalyst by its time-to-time removal by some means. The main object of this invention is to provide a process for the continuous production of carbon monoxide-free hydrogen from methane or methane-rich hydrocarbons at a temperature above about 600°C but below about 900°C. Another object is to provide a process for the continuous production of carbon monoxide-free hydrogen from methane or methane-rich hydrocarbons involving catalytic decomposition of methane to hydrogen and carbon, which is deposited on the catalyst used, and removal of the carbon by its gasification by steam with or without oxygen, while avoiding carbon build up on the catalyst and thereby avoiding the catalyst deactivation and increase in a pressure drop across the catalyst bed during the process. The present invention was accomplished by providing a process for the continuous production of carbon monoxide-free hydrogen by operating the process in two parallel reactors, both containing a solid catalyst comprising group VIII metal(s) and having different feeds so that the methane decomposition reaction and the carbon gasification reaction, involving the formation of CO, CC>2 and Fb are carried out simultaneously in cyclic manner in the two parallel reactors by regularly switching between them at an interval of time two different feed streams - one comprising methane and second comprising steam with or without oxygen, and thereby a continuous production of carbon monoxide-free hydrogen is effected from the methane decomposition reaction, the products of which are collected using the product stream switching valve, while collecting separately the products of the carbon reforming or gasification relation and flushing them from the reactor by pure hydrogen or by the products of methane decomposition, before switching simultaneously both the feed and product switching valve, without the catalyst deactivation and development of high pressure drop across the catalyst bed in both the reactors. Accordingly, this invention provides a process for the continuous production of carbon monoxide-free hydrogen from methane or methane-rich hydrocarbons, using a solid catalyst in two parallel catalytic reactors, the said process comprising the steps of: A process for the continuous production of carbon monoxide-free hydrogen from methane or methane-rich hydrocarbons, using a solid catalyst in two parallel catalytic reactors, said process comprising the steps of: i) treating the catalyst packed in two similar parallel reactors with a reducing agent, at a gas hourly space velocity in the range of 500 to 20,000 cm3, g-1. h-1 at a temperature in the range of 600 to 800°C and at a pressure of 1.0 atm for a period of at least 0.5 hour ; ii) contacting the first gaseous feed comprising methane or natural gas, called Feed A, at a gas hourly space velocity in the range of 250 cm3 .g-1. h-1 to 50,000 cm3, g-1. h-1 with the solid catalyst of first reactor, termed as Reactor A, at a temperature in the range of 600 to 1000°C and at a pressure of one atmosphere; iii) contacting simultaneously a second gaseous feed comprising steam with or without oxygen, hereafter termed Feed B1, at a gas hourly space velocity in the range of 250 cm3, g-1. h-1 to 50,000 cm3, g-1. h-1 with the solid catalyst in the second reactor, called Reactor B, at a temperature in the range of 600 to 1000°C, at a pressure of one atmosphere, iv) flushing the reactor through which feed B) was passed with the product gas A from feed A i.e. H2 termed as feed B2 , after a time ranging from 0.1 to 100 minutes so as to purge out all the product from stream B from feed Bl from this reactor before Feed A enters it, v) simultaneously switching (swapping) the two feeds, Feed A and Feed Bl using a feed stream switch over valve, the two product streams, product gas A obtained from Feed A and product stream B obtained from feed Bl using product switch over valve between the two parallel reactors, vi) collecting separately the two different gaseous products one consisting of carbon monoxide free H2 and unconverted methane obtained from Feed A and second consisting of CO,CO2,CH4 and H2 with or without 02 obtained from Feeds B1 and B2 after the removal of water by condensation, second consisting of CO, CO2 CH4 and H2, with or without O2 , obtained from Feeds B1 and B2 after the removal of water by condensation, vii) separating the carbon monoxide-free hydrogen and unconverted methane from Product Gas A and viii) repeating the steps (ii) to (vii) in a cycle. In an embodiment of the present invention the two parallel reactors are preferably fixed-bed reactors. In an another embodiment the solid catalyst used in step (i) is a group VIII metal(s) selected from the group consisting of Fe, Co, Ni, Ru, Rh, Pd, Pt, Ir, Os and a mixture thereof. In yet another embodiment the reducing agent used in step (i) is pure H2 or mixture containing at least 5 mol % Ha. In yet another embodiment in step (ii), the concentration of methane in Feed A is above 80 mol %. In yet another embodiment in step (iii), the oxygen to steam ratio used in Feed BI ranges from 0.0 to 2.0. In yet another embodiment in step (iii), the concentration of steam in the Feed BI ranges from about 50mol% to about 500mol%. In yet another embodiment in step (iii), the concentration of oxygen in feed BI ranges from 0 mol% to about 100mol%. In yet another embodiment the solid catalyst used is preferably a group VIII metal(s) selected from nickel with or without cobalt deposited on various micro or meso porous metal oxides selected from the group consisting of alumina, silica-alumina, silica, zerconia, yettria, ceria, magnesia and zeolites or zeolite like materials. In yet another embodiment the preferred temperature in each of the two reactors, Reactor A and Reactor B is ranging from 600°C to 800°C. In yet another embodiment the preferred gas hourly space velocity of feed A used in step (ii) ranges from 500 to 25000 cmY'h'1. In yet another embodiment the preferred gas hourly space velocity of feed BI in step (iii), ranges from 500 to 25000 crnYV. In yet another embodiment the preferred the oxygen to steam ratio used in Feed BI ranges from zero to 1.0. In yet another embodiment the preferred interval of time of the feed stream and product stream switch over ranges from 1 minute to 30 minutes. In yet another embodiment the preferred concentration of steam in the feed BI ranges from about 50mole% to about 100mole%. In still another embodiment the preferred concentration of oxygen in feed BI ranges from zero mol% to about 50mol%. The process of this invention will be more fully understood by reference to the attached drawing to which reference is made in the examples. REFERENCE TO THE DRAWING FIG. 1 shows a schematic flow sheet of the process of this invention. The description of labels 1-9 in the drawing are as follows: 1= Reactor A; 2 = Reactor B: 3 = catalyst fixed bed or fluid bed; 4 = Feed A (which comprises methane or methane-rich hydrocarbons); 5 = Feed BI (which comprises steam with or without oxygen gas) or Feed B2 (which is pure Ha or product of Feed A); 6 = Products Gas A (CO-free Ha with unconverted methane, obtained from Feed A); 7 = Products Gas B (CO, COa, CH4 and Ha, with or without Oi, obtained from Feeds BI and B2); 8 = Feed stream switch over valve and 9 = product stream switch over valve. The methane decomposition and carbon gasification reactions occur in Reactor A and Reactor B, respectively. After switching the two valves simultaneously, the methane decomposition and carbon gasification reactions occur in Reactor B and Reactor A, respectively. The gaseous products of the two reactors are collected separately in the cyclic operation, as shown in . In the process of this invention, the two parallel reactors may be two parallel fluid bed reactors or two parallel fixed-bed reactors; the preferred two parallel reactors are fixed bed reactors. The two different feed streams, Feed A and Feed Bl and also the two different product streams, Product Gas A (products of Feed A) and Product Gas B (products of Feeds Bl and B2), can be switched between the two reactors simultaneously using the two feed and product switch over valves operated manually or automatically, as shown in Figure 1. The group VIII metals are Fe, Co, Ni, Ru, Rh, Pd, Pt, Ir and Os. The main product of the process of this invention is carbon monoxide-free hydrogen, which is formed in the methane decomposition reaction. The side products are a mixture carbon dioxide, carbon monoxide and hydrogen, which are formed in the reforming or gasification by steam and/or oxygen of the carbon deposited on the catalyst. At a particular time, the main reactions occurring in the two parallel reactors are as follows: In the Reactor-A through which the feed comprising methane (Feed A) is passed, the catalytic decomposition of methane occurs producing two moles of hydrogen and one mole of carbon, which is deposited on the catalyst, per mole of methane reacted. At the same time, in the second reactor through which the feed comprising steam with or without oxygen, Feed Bl, is passed, the steam and/or oxygen react with the carbon deposited on the catalyst to produce carbon monoxide, carbon dioxide and hydrogen. These two different reactions occur in a cyclic manner when the two feeds and two product lines are simultaneously switched between the two reactors at an interval of time and the gaseous products of the two reactions are collected separately. In the process of this invention, the two parallel reactors, Reactor A and Reactor B, are preferably fixed-bed reactors. The preferred solid catalyst used in the process of this invention may be selected from nickel, with or without cobalt, deposited on various micro or meso porous metal oxides, such as alumina, silica-alumina, silica, zerconia, yettria, ceria, magnesia and the like or zeolites or zeolite-like materials, such as HY, Ce-Na-Y, HM, Hp, H-ZSM-5, MCM-41 and the like. The nickel and/or cobalt present in the catalyst are in their reduced form or in their zero oxidation state The solid catalyst comprising group VIII metal(s) used in the process of this invention can be prepared by the coprecipitation or impregnation catalyst preparation techniques known in the prior art. The role of step-i of the process of this invention is to reduce the transition metal oxide present at least on the surface of the catalyst, for example nickel oxide, cobalt oxide, iron oxide, etc., present in the catalyst are reduced to the corresponding metal. This step is important one; the reduction of group VIII metal oxide present on the catalyst surface to its metallic form is essential for the catalytic activity of the in the process of this invention. In step-ii of the process of this invention, the methane or methane-rich hydrocarbons, and steam with or without and oxygen are reactants, which are converted at least partly in the process. The role of steam is to react with the carbon, which is formed in the decomposition of methane on the reduced catalyst, producing carbon monoxide, carbon dioxide and hydrogen and thereby removing the carbon deposited on the catalyst. The role of the oxygen is to activate the carbon, which is otherwise difficult to gasify by steam alone. The oxygen is consumed at least partly by its reaction with the carbon to form CO and COj. Role of the solid catalyst is to catalyse the methane decomposition reaction and the carbon gasification by steam and/or oxygen. In the process of this invention, two different products obtained from the two different feeds, Feed A and Feed Bl are collected separately. The product stream obtained from Feed A contains only hydrogen and unconverted methane; it is free from carbon monoxide. The unconverted methane is separated from hydrogen by the pressure swing adsorption-separation processes known in the prior art, and thereby, carbon monoxide-free hydrogen is produced by the process of this invention. The product stream obtained from Feed Bl contains the carbon gasification products CO, CO2, Ha, methane and unconverted steam and/or oxygen, which may be separated by processes known in the prior art. In the process of this invention, the flushing of Feed Bl and the products of its reaction from the reactor by pure Ha or by the products of Feed A, before switching simultaneously the two switch over valves, as shown in Figure 1, is essential to avoid contamination of the products of Feed A by the product of Feed Blin the next cycle. The present invention is described with respect to the following examples illustrating the process of this invention for the production of carbon monoxide-free hydrogen from methane or methane-rich hydrocarbons and steam, with or without oxygen, over different solid catalysts comprising nickel, with or without cobalt, at different process conditions. These examples are provided for illustrative purposes only and are not to be construed as limitations on the process of this invention. EXAMPLES Definitions of the terms used in the examples GHS V = Gas hourly space velocity, defined as the volume of a gaseous feed, measured at 0°C and 1.0 atmospheric pressure, passed over unit mass of catalyst per unit time. Percent methane conversion is defined as the mole% of methane present in the feed converted into products other than methane. It is estimated as follows. Methane conversion (%) = [(moles of methane in feed - moles of methane in products) + (moles of methane in feed)] x 100 Hydrogen productivity in the process is expressed as the amount of hydrogen in mmol produced per hour per gram of the catalyst used in the process. EXAMPLES 1 - 8 These examples illustrate the process of this invention for the continuous production of carbon monoxide-free hydrogen from methane, using Ni-ZrC>2 (Ni/Zr mole ratio = 1.0) catalyst at different process conditions. The Ni/ZrOa catalyst was obtained by reducing NiO/ZrO2 by hydrogen at 600°C for 6h. The NiO-ZrOi catalyst was prepared by coprecipitating mixed hydroxides of nickel and zirconium from a 650cm3 aqueous solution containing 58.78g Ni(NO3)2.6H2O and 46.76g ZrO(NO3)2 xHaO using an aqueous solution of NaOH at a pH of 9.0 at room temperature, filtering and thoroughly washing the precipitate with deionized water, drying the washed precipitate at 105°C for 18 h, pressing and crushing to particles of 0.3-0.4 mm size and calcining in air at 600°C for 2h. The surface area of the NiO/ZrO2 catalyst was 63 m2 g"1. The process of this invention using the Ni/ZrOa catalyst was carried out in two parallel stainless steel fixed bed reactors as shown schematically in Figure 1. Each reactor was packed with 0.4 g catalyst particles of size 0.3-0.6 mm and was kept in a tubular furnace. The two parallel reactors had two different feeds connected through a four way flow switch valve and also had two different product lines connected through a four way flow switch valve, as shown in Figure 1. Both the flow switch valves were operated simultaneously at an interval of time. The temperature in both the reactors was measured by a Chromel-Alumel thermocouple located at the center of the catalyst bed in both the reactors. Before carrying out the catalytic reaction, the catalyst in both the reactors was reduced by pretreating it with a H2-N2 mixture at the conditions given in Table-1. After the reduction of catalyst in both the reactors, Feed A comprising methane was passed over the reduced catalyst in Reactor A and simultaneously Feed B2 containing steam with or without oxygen was passed over the reduced catalyst in Reactor B, and before switching the two valves for the next cyclic operation, the Feed B1 was replaced by pure H2 and the reactor and feed and product lines are flushed by the F^ of volume at least one time the volume of the reactor and feed and product lines between the two switch valves. The cyclic process operation was accomplished by switching Feed A and Feed Bl between the two reactor regularly at an interval of time and replacing Feed Bl by pure H2 as described above before switching simultaneously the two switch valves, at the process conditions given in. The gaseous products obtained from the two different feeds were collected separately in two gas collectors after condensing and removing the water from them by condensation. The collected two different gaseous products were analyzed by gas chromatograph using a spherocarb column and thermal conductivity detector. For the gas chromatographic analysis of hydrogen in the products, high purity nitrogen was used as a carrier gas. Whereas, for the gas chromatographic analysis of the methane, carbon dioxide and carbon monoxide present in the products, helium was used as a carrier gas. The results obtained at the different process conditions are presented in Table-1. The results in Example 7 clearly show that when the cyclic process is operated at lower temperature, 501°C, than that, >600°C, used in the process of this invention, the conversion of methane and the productivity of CO-free H2 are very poor. EXAMPLES 9-11 These examples also illustrate the process of this invention for the production of carbon monoxide-free hydrogen from methane and steam with or without C»2, using following solid catalysts: Ni/Si-MCM-41, Ni/Ce-NaY and Co-Ni/MgO/SA5205. The Ni/Si-MCM-41 (12 wt% Ni) was prepared by impregnating 5.0g particles (0.3-0.4 mm of size) of high silica MCM-41 [prepared by the procedure given in the reference: Choudhary, V. R. and Sansare, S. D., Proc. Indian Acad. Sci. (Chem. Sci.), volume 109, number 4, page 229-233, August 1997] with 3.4g Ni(NC»3)2.6H2O from its aqueous solution by the incipient wetness technique, followed by drying at 105°C for 12h and calcining at 500°C for 2h and by reducing the catalyst by hydrogen at 500°C for 4h. The NiO/Ce-NaY (10 wt% Ni, 72% Ce-exchanged NaY) was prepared by impregnating 5.0 g particles (0.3-0.4 mm of size) 72% Ce-exchanged NaY (prepared by the procedure given in the reference: Choudhary, V. R., Srinivasan, K. R. and Akolekar, D. B., Zeolites, volume 9, page 115-119, year 1989) with 2.75g Ni(NO3)2.6H2O from its aqueous solution by the incipient wetness technique, followed by drying at 105°C for 12h and calcining at 500°C for 2h and then reducing the catalyst by H2 at 600°C for 12h.. The NiO-CoO/MgO/SA5205 with Co/Ni mole ratio of 0.2 and NiO-CoO and MgO loadings on SA 5205 support, which is macroporous low surface area sintered silica-alumina catalyst carrier obtained from Norton Company (USA), of 14.5 wt.% and 7.2 wt.%, respectively was prepared by the procedure described by Choudhary et al (Ref. V.R. Choudhary et al AIChEJ Journal vol. 47 page 1632 year 2001). The catalyst was reduced by 20 % H2 in N2 at SOOT for 4h. The process of this invention over each of the above catalysts was carried out in the two parallel reactors and following the procedure same as that described in the earlier examples (Examples 1-10), at the catalyst pretreatment and catalytic process conditions given in Table-2. The results, the methane conversion and the CO-free hydrogen produced in the process of this invention over the above mentioned catalysts are included in. (Table Remove) Natural gas containing 94.7 mol % methane, 2.5 mol % ethane, 1.0 mol % C3-C4 hydrocarbon, 0.2 mol % CO2 and balance nitrogen Novel Features and Advantages of the Process of this Invention over the Prior Art Processes for the Production of Hydrogen 1. Unlike the prior art steam reforming and auto thermal reforming processes, carbon monoxide-free hydrogen can be directly produced by the process of this invention. 2. Unlike the prior art steam reforming and autothermal reforming processes, the conversion of methane and steam take place separately in two different reactors, two parallel reactors, each having a different feed, and the two different feeds: one comprising methane and second comprising steam with or without oxygen, are switched regularly between the two reactors at an interval of time so that when methane decomposition reaction occurs in one reactor producing hydrogen and carbon deposited on the catalyst, at the same time the gasification of the carbon by its reaction with steam and/or oxygen, producing carbon monoxide, carbon dioxide and hydrogen, takes place in the second reactor, and these two reactions in two separate reactors occur place in cyclic manner and the products of these two reactions are collected separately, so that there is no build up of carbon on the catalyst in both the parallel reactors and C()-free Ib is produced from the hydrocarbon decomposition reaction. 3. Unlike the prior art processes based on low temperature methane decomposition, the production of carbon monoxide-free hydrogen in the process of this invention is continuous without build up of carbon on the catalyst and consequently without the build up of a large pressure drop across the catalyst bed and also without the catalyst deactivation by carbon deposition or coking. 4. Since in the process of this invention, the hydrocarbon decomposition and carbon gasification reactions are carried out at higher temperature than that employed in the prior art processes based on the methane decomposition, the thermodynamic barrier on both the reactions in the present case is much lower and also the conversions of methane are much higher. We Claim: 1. A process for the continuous production of carbon monoxide-free hydrogen from methane or methane-rich hydrocarbons, using a solid catalyst in two parallel catalytic reactors, the said process comprising the steps of: i) treating the catalyst packed in two similar parallel reactors with a reducing agent, at a gas hourly space velocity in the range of 500 to 20,000 cm3, g-1. h-1 at a temperature in the range of 600 to 800°C and at a pressure of 1.0 atm for a period of at least 0.5 hour ; ii) contacting the first gaseous feed comprising methane or natural gas, called Feed A, at a gas hourly space velocity in the range of 250 cm .g g-1 . h-1 to 50,000 cm3, g-1. h-1 with the solid catalyst of first reactor, termed as Reactor A, at a temperature in the range of 600 to 1000°C and at a pressure of one atmosphere; iii) contacting simultaneously a second gaseous feed comprising steam with or without oxygen, hereafter termed Feed B1, at a gas hourly space velocity in the range of 250 cm3, g-1 . h-1 to 50,000 cm3, g-1 . h-1 with the solid catalyst in the second reactor, called Reactor B, at a temperature in the range of 600 to 1000°C, at a pressure of one atmosphere, iv) flushing the reactor through which feed B1 was passed with the product gas A from feed A i.e. H2 termed as feed B2 , after a time ranging from 0.1 to 100 minutes so as to purge out all the product from stream B from feed B1 from this reactor before Feed A enters it, v) simultaneously switching (swapping) the two feeds, Feed A and Feed B1 using a feed stream switch over valve, the two product streams, product gas A obtained from Feed A and product stream B obtained from feed BI using product switch over valve between the two parallel reactors, vi) collecting separately the two different gaseous products one consisting of carbon monoxide free H2 and unconverted methane obtained from Feed A and second consisting of CO,CO2,CH4 and H2 with or without O2 obtained from Feeds B1 and B2 after the removal of water by condensation, second consisting of CO, CO2 CH4 and H2, with or without O2 , obtained from Feeds Bl and B2 after the removal of water by condensation, vii) separating the carbon monoxide-free hydrogen and unconverted methane from Product Gas A and viii)repeating the steps (ii) to (vii) in a cycle. 2. A process as claimed in claim 1, wherein the two parallel reactors re preferably fixed-bed reactors. 3. A process as claims 1&2, wherein the solid catalyst used in step (i) is a group VIII metal (s) selected from the group consisting of Fe, Co, Ni, Ru, Rh, Pd, Pt, Ir, Os and a mixture thereof. 4. A process as claimed in claims 1-3, wherein the reducing agent used in step (i) is pure H2 pr H2/N2 mixture containing at least 5 mol% H2. 5. A process as claimed in claims 1-4, wherein in step (ii), the concentration of methane in Feed A is above 80 mol%. 6. A process as claimed in claims 1-5, wherein in step (iii), the oxygen to steam ratio used in Feed B1 ranges from 0 to 2.0. 7. A process as claimed in claims 1-6, wherein in step (iii), the concentration of steam in the Feed B1 ranges from 50 mol% to 500mol%. 8. A process as claimed in claims 1-7, wherein in step (iii), the concentration of oxygen in feed B1 ranges from 0 mol% to 100mol%. 9. A process as claimed in claims 1-8, wherein the solid catalyst used is preferably a group VIII metal(s) selected from nickel with or without cobalt deposited on various micro or meso porous metal oxides selected from the group consisting of alumina, silica-alumina, silica, zerconia, yettria, ceria, magnesia and zeolites or zeolite like materials. 10. A process as claimed in claims 1-9, wherein the preferred temperature in each of the two parallel reactors is ranging from 600°C to 800°C. 11. A process as claimed in claims 1-10, wherein the preferred gas hourly space velocity of feed A used in step (ii) ranges from 500 to 25000 cm3g-1h-1. 12. A process as claimed in claims 1-11, wherein the preferred gas hourly space velocity of feed B1 in step (iii), ranges from 500 to 25000 cm3g-1h-1. 13. A process as claimed in claims 1-12, wherein the preferred oxygen to steam ratio used in Feed B1 ranges from 0.01 to 1.0. 14. A process as claimed claims 1-13, wherein the preferred interval of time of the feed stream and product stream switch over ranges from 1 minute to 30 minutes. 15. A process as claimed Claims 1-14, wherein the preferred concentration of steam in the feed B1 ranges from 50 mole% to 100 mole%. 16. A process for the continuous production of carbon monoxide-free hydrogen from methane or methane-rich hydrocarbons, using a solid catalyst in two parallel catalytic reactors, substantially as herein described with reference to the examples and drawing accompanying this specification.

Documents:

1225-DEL-2002-Abstract-(28-11-2008).pdf

1225-DEL-2002-Claims-(28-11-2008).pdf

1225-del-2002-claims-17-12-2008.pdf

1225-del-2002-claims.pdf

1225-DEL-2002-Correspondence-Others-(17-12-2008).pdf

1225-DEL-2002-Correspondence-Others-(28-11-2008).pdf

1225-DEL-2002-Correspondence-Others-05-12-2008.pdf

1225-del-2002-correspondence-others.pdf

1225-del-2002-correspondence-po.pdf

1225-DEL-2002-Description (Complete)-(17-12-2008).pdf

1225-DEL-2002-Description (Complete)-(28-11-2008).pdf

1225-del-2002-description (complete).pdf

1225-del-2002-drawings.pdf

1225-del-2002-form-1.pdf

1225-del-2002-form-18.pdf

1225-del-2002-form-2.pdf

1225-DEL-2002-Form-3-(28-11-2008)..pdf

1225-del-2002-form-3.pdf

1225-DEL-2002-Petition-137-05-12-2008.pdf