CA2354927A1 - Fuel process and apparatus and control system - Google Patents

Abstract

A fuel processor apparatus comprising a catalytic tubular reactor is heated using an infrared radiant burner to provide the endothermic heat of reaction needed to reform a mixture of hydrocarbon and steam for the production of a hydrogen-rich gas stream. The hydrogen-rich gas stream is further purified using a sequence of catalytic steps and is fed to a fuel cell whereupon a portion of the hydrogen contained in said gas stream is consumed for the production of electricity by electrochemical reaction with oxygen. An unused portion of the purified hydrogen-rich gas stream exits the fuel cell stack and is combusted in said infrared radiant burner. A
fuel cell control system rapidly responds to a variable fuel cell electric demand by adjusting the feed of hydrocarbon to the catalytic tubular reactor to maintain the surface temperature of the infrared radiant burner within defined limits.

Description

FUEL PROCESS AND APPARATUS AND CONTROL SYSTEM
FIELD OF THE INVENTION
The present invention relates to an endothermic catalytic reaction apparatus for producing hydrogen-containing gases from hydrocarbon feedstock.
BACKGROUND OF THE INVENTION
Proton exchange membrane (PEM) fuel cells have emerged as a viable option for the production of disbursed electrical power, typically in the range of 2-25 KW
and are particularly suited for use in residential and small commercial applications. PEM fuel cells generate electricity by the electrochemical reaction between hydrogen and oxygen. While oxygen is readily available from ambient air, hydrogen must be produced from commercially available fuels, such as natural gas or propane, using methods such as steam reforming, a process that involves a high temperature catalytic reaction between a hydrocarbon and steam to form a hydrogen-rich product gas that contains significant quantities of carbon monoxide. It is desirable that the PEM fuel cell deliver electricity upon demand. This requires that the fuel processor be capable of delivering variable quantities of hydrogen to the PEM fuel cell stack in accordance with the electric load requirement.
Because PEM fuel cells have a low tolerance to carbon monoxide, typically less than 10 ppm, additional processing steps are required to prepare a hydrogen-rich gas stream that is suitable for use in a PEM fuel cell. These additional processing steps typically include application of one or more stages of water-gas shift reaction that ultimately reduce the carbon monoxide concentration to about 10,000 ppm, and a preferential reaction step that further reduces the carbon monoxide concentration to less than 10 ppm. The physical embodiment of the process equipment that achieves the combination of reaction steps needed to convert the hydrocarbon feed to a useful hydrogen product is commonly referred to as a fuel processor. Table 1 summarizes the reaction steps of a fuel processor designed to produce a hydrogen-rich gas stream suitable for use in a PEM fuel cell.
Table 1. PEM fuel processor reaction steps 1. CH4 + H20 = CO + 3 Hz Steam reforming 2. CO + H20 = COZ + Hz Water-gas shift 3. CO + %zOz = COz Preferential oxidation Various methods to generate hydrogen from hydrocarbon fuels for industrial purposes using the combination of steam reforming and water-gas shift reaction steps are well known in the prior art. However, apparatus of compact design and rapid load-following response are desirable for use in a PEM fuel cell application.
In commercial steam reformers for large-scale production of hydrogen from hydrocarbon I 5 feeds, endothermic heat is commonly supplied by the combustion of carbonaceous fuel and oxidant in a diffusion or turbulent flame burner that radiates to the refractory walls of a combustion chamber, thereby heating them to incandescence, and providing a radiant source for heat transfer to a tubular reaction chamber. Uniform radiation to the surfaces of the tubular reaction chamber is essential since excessive local overheating of the tube surface can result in mechanical failure. In large-scale commercial steam reformers, maldistribution of heat within the furnace chamber is minimized by providing large spacing between the individual reactor tubes, the furnace walls, and the burner flames. However, for small-scale catalytic reaction apparatus that is uniquely compact, such as for the production of hydrogen for small fuel cell applications, special design features must be used to prevent tube overheating.
U.S. Patent No. 4,692,306 to Minet and Warren (the '306 patent) describe a compact reformer comprising an annular reaction chamber concentrically disposed around an internal burner chamber containing a vertically disposed cylindrical radiant burner that uniformly radiates in the radial direction. The inventors allege to achieve a uniquely uniform radiation pattern to a concentrically disposed annular reaction chamber that surrounds the radiant burner, thereby avoiding the problerr~s with flame impingement and local overheating of tube surfaces that are associated with the use of diffusion or turbulent flame burners in compact reformer apparatus.
However, there are practical limitations regarding the use of an annular reaction chamber geometry for very small-scale reformers having hydrogen production rates of less than about 500 SCFH. It is well known that the heat transfer coefficient of gaseous reactions contained within an annular reaction chamber is directly related to the velocity of the gaseous reactants within the annular space. In order to limit the reaction chamber wall temperature, the velocity of gaseous 1 S reactants within the annular space must be sufficiently high to absorb the radiant heat flux that impinges on the reaction chamber tube walls. For very small-scale reformers, this requires that the width of the annular reaction chamber space be small. It common practice in the art to limit the maximum diameter of the catalyst particles packed within an annular space to less than 20 percent of the width of the annular space in order to ensure that the catalyst is evenly distributed within the reaction chamber and to prevent gas channeling along the walls of the reaction chamber. However, for annulus having a small width dimension, this requires use of catalyst particles of particularly small diameters thereby resulting in a high pressure drop through the catalyst bed.
The compact endothermic catalytic reaction apparatus embodied in the present invention achieves the objects of compact design, while avoiding the problems of flame impingement, excessive reaction chamber wall temperatures, and excessive reaction chamber pressure drop by application of a tubular reaction chamber that is heated by radiant burner having an active radiant surface that is defined by a geometric arc in order to maximize the flux of radiant energy to the cold-plane surface of the tubular endothermic reaction chamber. The tubular endothermic reaction chamber is operated using a combination of catalyst particle sizes and reactant mass velocities to control the reactor pressure drop and the maximum reaction chamber tube wall temperature within certain desirable limits; and the radiant burner is operated at specific ranges of combustion intensity and excess air to control surface temperature of the radiant burner within certain desirable limits.
The present invention minimizes the radiant exchange between burner surfaces and adiabatic refractory surfaces that leads to overheating of the radiant burner and thus extends the practical range of tubular endothermic reaction chamber geometry that can be used in combination with radiant burners for converting hydrocarbon feedstock to useful industrial gases.
The benefits ofusing an infrared radiant burner to accomplish a compact design are described by the patent. However, these teachings do not describe the use of a control system that rapidly responds to load variations based on a measure of the radiant burner surface temperature.
U.S. Patent Nos.6,051,192 and 4,098,960 describe the use of a tubular steam reformer that is heated by a flame burner and a load-following control system that is based on fuel flow rate. US

Patent No. 4,943,493 describes a control system using an oxygen sensor. None of these systems have the rapid response time that can be obtained by using the burner surface temperature.
PEM fuel cells require supply of a hydrogen-rich gas stream containing very low concentrations of carbon monoxide, typically less than 10 ppm. The production of highly-purified hydrogen at the industrial scale is generally accomplished using pressure swing adsorption (PSA) technology. However, the need for compact design for PEM fuel cell applications generally precludes the use of PSA technology for hydrogen purification. Membrane purifiers have been developed to produce a highly-purified hydrogen product (U.S. Patents 5,612,Oi2; 5,639,431; and 5,861,137), however these membrane purifiers require high operating pressures.
The fuels commonly available for most residential and commercial applications are generally not supplied at the high operating pressure required for use with membrane separators.
Therefore, fuel processors with membrane separators generally require a means for compressing the fuel.
Preferential oxidation, sometimes referred to as preferential oxidation, has been developed as an economical method to reduce the concentration of carbon monoxide in the feedgas sent to PEM
fuel cells. Because the preferential oxidation reaction is highly exothermic, methods are needed to control the temperature within the preferential oxidation reactor. For instance, U.S. Patent No.5,658,681 describes a cooling method for a preferential oxidizer reactor.
SUMMARY OF THE INVENTION
The present invention relates to the processing and apparatus, the sequence of combining said apparatus, and the method of controlling said apparatus, that is used by a fuel processor in order to achieve the objectives of producing a hydrogen rich product gas having less than 10 ppm carbon monoxide using a compact design that offers rapid load-following response, high thermal efficiency, and low pressure operation. The fuel processor is the embodiment of a sequence of chemical reactions and separation processes that incorporates the following apparatus:
(a) a desulfurizer to remove the sulfur compounds that are contained hydrocarbon fuels, such as natural gas and liquefied petroleum gas, (b) an endothermic catalytic reaction apparatus comprising a burner, a combustion chamber, a steam reformer reactor and including a method of heating and control of said apparatus, (c) a steam generator that is heat integrated with several different reactors that are part of the fuel processor, (d) a unique mechanical design to accomplish successive stages of the water-gas shift reaction within a single mechanical vessel, (e) a preferential oxidation reactor of uniquely economic design capable of decreasing the carbon monoxide concentration to less than 10 ppm, (f) a method of control of said fuel processor to produce the correct net amount of hydrogen necessary to follow the electrical load that is required by an integrated polymer electrolyte membrane fuel cell stack, that uniquely fits the requirement of a residential electricity generator, and (g) equipment to separate and recycle water that is part of the process gas that exits from the water gas shift reactors.
The endothermic catalytic reaction apparatus includes a hairpin tubular reaction chamber vertically disposed within a combustion chamber and packed with a catalyst for the conversion of hydrocarbon to industrial gases by reaction with steam; said hairpin tubular reaction chamber having an upper portion that is surrounded by a convection chamber to enhance the transfer of heat from combustion products to the reaction chamber. A radiant burner is vertically disposed within a combustion chamber having a gas permeable zone that promotes the flameless combustion of fuel and oxidant in order to heat the metal fiber surface to incandescence and radiate energy to the reaction chamber. The radiant burner having a controlled geometric design so that the angle of radiation is predominantly incident upon the tubular reaction chamber surface.
Preferably, the hairpin tubular reaction chamber comprising a tube having outer diameters ranging from 3/8" to 4", preferably from 3/4" to 2.5" and the mass velocities in the chamber range from 200 lb/ft'-/h to 3000 lb/ft2/h, preferably from 400 lb/ft2/h to 1500 lb/ft2/h.
The average catalyst particle diameters ranging from 1/16" to 3/4", preferably from'/4" to %Z"
and pressure drops ranging from 1 psi to 8 psi.
The inlet operating pressure of the reaction chamber ranges from 0 psig to 20 psig and preferably from 2 psig to 10 psig. The exit temperature ranges from 1000 °F to 1700 °F and preferably from 1200 °F to 1400 °F, and the maximum tube wall temperatures that range from 1200 °F to 2000 °F and preferably from 1400 °F to 1600 °F.
Average heat fluxes range from 2,000 Btu/ft2/h to 20,000 Btu/ft2/h and preferably from 5,000 Btu/ftz/h to 10,000 Btu/ftz/h.
The hairpin tubular reaction chamber preferably has capacity to generate hydrogen plus carbon monoxide product in volumetric quantities ranging from 20 SCFH to 1000 SCFH and preferably from 25 SCFH to 600 SCFH, and is particularly well suited for generating hydrogen plus carbon monoxide in the range of 25 SCFH to 150 SCFH.
The radiant burner is preferably made of a supported porous ceramic material having extended life at operating temperatures up to 2100 °F.

The radiant burner may be made of a supported metal fiber material typically fabricated from an alloy containing principally iron, chromium, and aluminum and smaller quantities of yttrium, silicon, and manganese having extended life at operating temperatures up to 2000 °F.
The radiant burner having a controlled design such that the angle of radiation ranges from 45 to 180 degrees. Surface temperatures of the radiant burner range from 1700 °F to 2000 °F and preferably 1700 °F to 2000 °F.
The radiant burner limits the internal temperatures of the premixed fuel and oxidant to less than 600 °F and preferably less than 500 °F in order to prevent flashback, which is defined as the pre-combustion of the fuel and oxidant mixture within the burner feed pipe.
The radiant burner is insulated to prevent the transfer of heat to the premixed fuel gases from the non-permeable side of the radiant burner.
Combustion intensity in the radiant burner typically ranges from 100,000 Btu/ftz/h to 500,000 Btu/ft2lh and preferably from 150,000 Btu/ft2/h to 350,000 Btu/ft2lh.
Combustion intensity is defined as the higher heating value of the fuel combusted divided by the permeable radiant burner surface area.
The radiant burner that operates at an excess air ratio typically ranging from 10% to 200%
and preferably from 30% to 100%, wherein the excess air ratio is defined as percent combustion air in excess of the stoichiometric amount required for complete combustion of the burner fuel.
In a preferred embodiment, the endothermic catalytic reaction apparatus includes a a helical tubular reaction chamber disposed within a combustion chamber and packed with a catalyst for the conversion of hydrocarbon to industrial gases by reaction with steam, said helical tubular reaction chamber having an upper portion that is surrounded by a convection chamber to enhance the transfer of heat from combustion products to the reaction chamber and an exit section to convey reaction products to the exit means. The radiant burner is vertically disposed within a combustion chamber having a gas permeable zone that promotes the flameless combustion of fuel and oxidant in order to heat the metal fiber surface to incandescence and radiate energy to the reaction chamber; said radiant burner designed to radiate uniformly in the radial direction.
In this embodiment, the helical tubular reaction chamber includes a tube having outer diameters ranging from 3/8" to 4" and preferably from 3/4" to 2.5". The helical tubular reaction chamber having an outer coil diameter ranging from 2" to 24" and preferably from 5" to 18". Mass I O velocities ranging from 200 lb/ft2/h to 2000 lb/ftz/h and preferably from 400 lblft2/h to 1500 lb/ft2/h.
The average catalyst particle diameters range from 1/16" to 3/4"and preferably'/4" to %2" and pressure drops in the helical tubular section range from 1 psi to 8 psi. Inlet operating pressures ranging from 0 psig to 20 psig and preferably from 2 psig to 10 psig. exit temperature ranging from 1000 °F to 1700 °F and preferably from 1200 °F to 1400 °F. Maximum tube wall temperatures ranging from 1200 °F to 1700 °F and preferably from 1400 °F to 1600 °F.
The helical tubular reaction chamber having average heat fluxes ranging from 2000 Btu/ftZ/h to 20,000 Btulftz/h and preferably from 5,000 Btu/ft2/h to 10,000 Btu/ft2/h.
The helical tubular reaction chamber having capacity to generate hydrogen plus carbon monoxide product in volumetric quantities ranging from 20 SCFH to 1000 SCFH and preferably from 25 SCFH to 600 SCFH, and particularly well suited for generating hydrogen plus carbon monoxide in the range of 100 SCFH
to 300 SCFH.

The radiant burner is a supported porous ceramic material having extended life at operating temperatures up to 2100 °F. Preferably the supported metal fiber material is fabricated from an alloy containing principally iron, chromium, and aluminum and smaller quantities of yttrium, silicon, and manganese having extended life at operating temperatures up to 2000 °F.
The radiant burner has S a controlled design such that the radiant energy emanates from the burner in a substantially uniform radial pattern. Surface temperatures range between 1500 °F and 2000 °F and preferably 1700 °F and 2000 °F. The radiant burner limits the internal temperatures of the premixed fuel and oxidant to less than 600 °F and preferably less than 500 °F in order to prevent flashback, wherein flashback is defined as the pre-combustion of the fuel and oxidant mixture within the burner feed pipe. The radiant burner operates at a combustion intensity typically ranging from 100,000 Btu/ft'-/h and 500,000 Btu/ft2lh from 150,000 Btu/ftZ/h to 350,000 Btu/ftZlhr, wherein the combustion intensity is defined as above.
The radiant burner that operates at an excess air ratio typically ranging from 10% to 200%
and preferably from 30% to 100%, wherein the excess air ratio is defined as above.
1 S In another embodiment, the invention relates to a fuel processor having a desulfurizer vessel, an endothermic catalytic reaction apparatus comprising a burner, a combustion chamber, a steam reformer reactor, and including a method of heating and control of said apparatus, a steam generator that is heat integrated with several different reactors within said fuel processor, a single vessel for successive stages of the water-gas shift reaction vessel that is heat integrated with the steam generator, and a compact preferential oxidation reactor capable of decreasing the carbon monoxide concentration to less than 10 ppm. A control system is linked to the entire fuel processor and permits it to produce the quantity of hydrogen required by an integrated polymer electrode membrane fuel cell stack when said fuel cell stack is following changes in electrical load. A system for the separation and recycle of water is also preferably provided.
An endothermic tubular reaction apparatus is also provided. The reaction apparatus includes a hairpin tubular reaction chamber vertically disposed within a combustion chamber and packed with a catalyst for the conversion of hydrocarbon to industrial gases by reaction with steam; said hairpin tubular reaction chamber having an upper portion that is surrounded by a convection chamber to enhance the transfer of heat from combustion products to the reaction chamber.
Aa radiant burner vertically disposed within a combustion chamber having a gas permeable zone that promotes the flameless combustion of fuel and oxidant in order to heat the metal fiber surface to incandescence and radiate energy to the reaction chamber; said radiant burner having a controlled geometric design so that the angle of radiation is predominantly incident upon the tubular reaction chamber surface.
The steam generator is heated by a combination of three sources in the form of process streams that provide recovered energy from the steam reformer and water gas shift reaction chambers 1 S plus a fourth energy source in the form of electricity that is transformed into heat by an immersion heater inside the steam generator. An electrical immersion heater is provided and has a capacity from 0.1 kW to 2 kW preferably 0.5 kW to 1.5 kW. One of the three process gas' streams used to provide energy is preferably the flue gas stream from the combustion chamber of the steam reformer.
In other embodiments, one of the three process gas streams used to provide energy is exit gas from the bed of high temperature water-gas shift catalyst that exits from the middle of the water gas shift reaction vessel flows through the steam generator to transfer heat and decrease its temperature, and then re-enters the water-gas shift reaction vessel at its mid-point to flow down over the low temperature water-gas shift catalyst. In yet other embodiments, one of the three process gas streams is the exit gas from the bed of low temperature water-gas shift catalyst that flows out of the bottom of the water-gas shift reaction vessel.
The water-gas shift reaction vessel that is heat integrated with the steam generator where successive stages of the water-gas shift reaction are performed at different temperatures each in separate chambers within a single reaction vessel. In preferred embodments, the first stage of the water-gas shift reaction is performed in a chamber of the water-gas shift reaction vessel that contains a commercial high temperature shift catalyst composed of chromium, iron, and copper oxides; at a temperature from 600 °F to 900°F and preferably from 600°F to 800°F; at a gas hourly space velocity from 200 h'' to 10,000 h'' and preferably from 1000 h'' to 8000 h'';
and at steam to carbon ratios expressed as moles of water per atom of carbon from 1.8 to 6 and preferably from 2 to 4 The process gas exiting from the first chamber of the water-gas shift reaction vessel may be directed to flow through a coil that passes through a steam generator, thereby decreasing its temperature to that desired for the subsequent stage of the water-gas reaction and transfernng the corresponding amount of energy to produce steam in the steam generator.
The subsequent stage of the water-gas shift reaction is preferably performed in a chamber of the water-gas shift reaction vessel that contains a commercial low temperature shift catalyst composed of copper, zinc, and aluminum oxides at a temperature from 350°F to 660°F and preferably 350°F to 550°F, at a gas hourly space velocity from 200 h'' to 10,000 h'' and preferably from 1000 h'' to 8000 h''; and at steam to carbon ratios, expressed as moles of water per atom of carbon, from 1.8 to 6 and preferably from 2 to 4.
The preferential oxidation reactor is preferably capable of decreasing the concentration of carbon monoxide to less than 10 ppm both at steady-state operation and during transient operating conditions between different steady-states. The preferential oxidation reactor may be provided as a series of tubular reactors in parallel.
The preferential oxidation reactor may have heat transfer fins on the 50 percent of its length closest to the feed inlet and may include a fan to enhance the removal of the exothermic heat of reaction by the mechanism of forced convection resulting from air passing over the reactor.
Typically the preferential oxidation reactor is operated at temperatures from 80°F to 200°F
preferably 100°F to 180°F. The preferential oxidation reactor is operated at gas hourly space velocities from 1000 h'' to 20,000 h'' preferably 2000 h'' to 10,000 h'' and at excess air stoichiometries of 80 % to 400 percent preferably 100 percent to 300 percent.
The excess air stoichiometry is defined as the ratio of amount of oxygen fed to the reactor divided by the amount 1 S of oxygen required to convert all of the carbon monoxide.
The desulfurizer is a vessel filled with an adsorbent consisting of an activated carbon and an activated carbon that has been promoted with copper.
The control system is preferably based on the temperature of the gas phase burner face temperature in the combustion chamber of said endothermic catalytic reaction apparatus.
An increase in the burner face temperature above its set point causes the control system to decrease the amount of hydrocarbon that is fed to the catalytic steam reformer reactor that is part of said endothermic catalytic reaction apparatus, and conversely a decrease in said temperature below its set point causes an increase in the amount of hydrocarbon fed to the catalytic steam reformer reactor. The set point for the burner face temperature is decreased by the control system if the temperature of the reformer exit temperature exceeds a predetermined value and conversely is increased by control system if the temperature of the reformer exit temperature becomes less than the predetermined value.
The system for the separation and recycle of water typically includes a vertically mounted separator vessel that allows the dried gas to flo~.v out of the top of the separator to the preferential oxidation reactor and the liquid phase water to flow out of the bottom to a pump suction that sends it to the steam generator.
The feedstock is any hydrocarbon or mixture ofhydrocarbons including natural gas, liquefied petroleum gas, gasoline, diesel fuel. The hydrocarbon is desulfurized in a desulfurizer, combined with steam, heated in a feed-effluent heat exchanger, heat further by exchange with flue gas from the combustion chamber, passed over a fixed bed of catalyst inside a combustion chamber to be converted to hydrogen and carbon monoxide, cooled in a feed-effluent heat exchanger, passed over a fixed bed of high temperature water-gas shift catalyst to convert some of the carbon monoxide to hydrogen, by flowing through a coil that passes through a steam generator, passed over a fixed bed of low temperature water-gas shift catalyst to convert more of the carbon monoxide to hydrogen, passed through another coil inside a steam generator where it transferred heat to generate steam, passed through an air cooler to decrease its temperature to cause water to condense in a separator, mixed with air and passed through a fixed bed of preferential oxidation catalyst where the remaining carbon monoxide is converted so that the hydrogen stream contains less than 10 ppm of carbon monoxide and passed into a fuel cell stack where most of the hydrogen is consumed in the generation of electricity.
The hydrogen not consumed in the fuel cell stack flows to the combustion chamber where it is burned. Some of the hydrocarbon is also burned directly in the combustion chamber. By passing the flue gas from the combustion chamber through a coil in the steam generator energy is used to generate steam.
BRIEF DESCRIPTION OF DRAWINGS
Fig. 1 is a process flow scheme of the fuel processing system according to the present invention.
Fig. 2 is a preferred embodiment of the feed control scheme of the invention.
Fig. 3 is a preferred embodiment of the catalytic reaction apparatus of the invention.
Fig. 4 is a geometric representation of a preferred embodiment of the catalytic reaction apparatus of the invention.
Fig. 5 is an alternative preferred embodiment of the steam reformer of the invention.
Fig. 6 is a preferred embodiment of the reaction vessel for the successive stages of the water-gas shift reaction, including heat exchange for steam generation.
Fig. 7 is a preferred embodiment of the vessel used to produce steam for the steam reforming process, including heat exchange with process streams from other reactors within the fuel processor.
Fig. 8 is a preferred embodiment of a preferential oxidation reactor, include heat exchange means for temperature control.

DETAILED DESCRIPTION OF THE INVENTION
The process and apparatus of the present invention achieves the object of generating a hydrogen-rich gas stream from a hydrocarbon feed having quality suitable for use in the generation of electricity in a PEM fuel cell stack.
Referring to Fig. 1, a hydrocarbon feed 1 such as natural gas or propane, is fed to a desulfiu-izer 2, preferably a copper impregnated activated carbon adsorbent wherein odorants, such as mercaptans, contained within the hydrocarbon feed are adsorbed to produce an essentially sulfur-free feed gas. A suitable adsorbent is commercially available as G-32J from Sud-Chemie Corporation. Desulfurization of the hydrocarbon feed prevents contamination of downstream catalytic processes necessary for the production of said hydrogen-rich gas stream.
The desulfurized feed 3 is mixed with steam 4 and the feed/steam mixture 5 is preheated by heat exchange against hot reformate 6 that exits from the steam reformer 7.
The steam reformer includes a tubular reactor 8 containing catalyst 9. The catalyst preferably contains an active nickel-oxide component that is well known in the art to promote the reaction between hydrocarbon and steam to produce a hydrogen-rich gas stream, or reformate, in a process commonly referred to as steam reforming. A suitable catalyst is commercially available as G-91 from Sud-Chemie Corporation.
A high degree of conversion of the hydrocarbon feed is achieved by heating the reformate to a temperature in the range of 1100°F to 1300°F at the exit 10 of tubular reactor 8. The temperature of the reformate at the exit of the steam reformer is preferably selected so that the methane concentration in the dry gas is less than about 3% by volume.

The endothermic heat of reaction within the tubular reactor is supplied by combusting fuel and air in infrared radiant burner 11. A primary portion of the endothermic heat of reaction is transferred by radiation from burner face 12 to sections of the tubular reactor that are located within radiant chamber 13. A secondary portion of the endothermic heat of reaction is transferred by S convection from the flue gases contained within the annular space defined between outer flue gas conduit 14 and inner flue gas conduit 15. The flue gases exiting the steam reformer are sent to flue gas waste heat recovery coil 16 that are in communication with water 17 contained in a steam generator 18 to recover heat to generate steam that is used in the process.
The hot reformate 6 exiting the steam reformer is cooled to a temperature of about 6S0 °F to 700°F by heat exchange against the feed/steam mixture 5 in a feed/effluent exchanger 19. The cooled gas mixture enters an integrated reaction vessel that includes a high temperature shift reactor 20, intercooler coil 21 and low temperature shift reactor 22. High temperature shift reactor ZO
contains a commercially available catalyst, which preferably includes the active elements of iron and chromium, that promotes a reaction between steam and carbon monoxide, commonly known as the 1 S water-gas shift, to produce hydrogen and carbon dioxide.
The high temperature shift effluent gases 23 pass into the intercooler coil that is in communication with water contained in the steam generator wherein the effluent gases are cooled to a temperature range of about 430°F-460°F before entering the inlet 24 to low temperature shift reactor. A portion of the high temperature shift reactor shell 25 is also in communication with water contained in the steam generator, which serves to remove a portion of the exothermic heat ofreaction occurring within the high temperature shift reactor so that the temperature of the gases at the exit 23 is typically in the range of 600°F - 650°F.
Low temperature shift reactor 22 contains a catalyst which preferably includes the active elements of copper and zinc, that further promotes water-gas shift reaction at lower temperatures so as to produce a hydrogen-rich gas stream at the exit containing about 1%
carbon monoxide by volume. The low temperature shift effluent gases 26 enter after cooler coil 27 wherein the gases are further cooled.
After exiting after cooler coil 27, the hydrogen-rich gas stream is cooled, preferably to about 90°F in air-cooled heat exchanger 28 to condense water contained within the gas stream. The condensed water is separated from dry gas stream 29 in separator vessel 30 where it is combined with additional water 31 and returned to the steam generator via boiler feedwater pump 32, after being preheated by heat exchange against flue gases in boiler feedwater preheater 33.
The dry gas stream exiting the separator vessel is reheated to the temperature range of typically between 150 °F to 170 °F by heat exchange against flue gases in reheater 34. The reheated dry gas stream is then mixed with a metered quantity of air 35 before being sent to preferential oxidizer reactor 36. Preferential oxidizer 36 contains catalyst 37 that is packed within the tubes 38 of finned tube air-cooled heat exchanger 39. Catalyst 37 contains platinum and other catalytic promoters to preferentially combine oxygen with carbon monoxide to produce carbon dioxide even in the presence of a high concentration of hydrogen. Catalyst 37 preferably has an optimum operating temperature in the range of 120°F to 220°F. Since the oxidation reaction is highly exothermic the heat of reaction is removed by passing air over the surface of the catalyst-packed finned tubes. The airflow pattern is maintained to keep all points within the catalyst bed within the optimum operating temperature range. The concentration of carbon monoxide in anode feed gas 40 at the exit of the preferential oxidizer is typically less than 10 ppm.
The anode feed gas 40 exiting the preferential oxidizer is of a quality suitable for feed to a PEM fuel cell stack 41 for the production of electricity 42 by electrochemical reaction between oxygen and a portion of the hydrogen contained in the anode feed gas. The unused portion of the anode feed gas, referred to as anode offgas 43, is transported to infrared radiant burner 11 of steam reformer 7 wherein it is combusted with air to provide the endothermic heat of reaction for the production of reformate within the tubular reactor of the steam reformer.
The feed control system of the present invention provides rapid response to independently varying electric load demand variations at the fuel cell stack. Refernng to Fig. 2, anode feed gas 40 is sent to fuel cell stack 41 wherein a portion of the hydrogen contained in the anode feed gas is consumed by electrochemical reaction to produce electricity 42. Anode off gas 43 flows to an infrared radiant burner 11 wherein said gas is premixed with air 44 and is ignited on burner face 12.
1 S Burner face 12 includes a porous material consisting of ceramic or metal fibers. The porous surface has a low thermal inertia so that the temperature of the burner face responds extremely rapidly to variations in the quantity of anode off gas that is combusted within the infrared radiant burner.
For instance, the burner face temperature can be heated from room temperature to 1600°F
in a period of just a few seconds from the start of fuel flow to the burner.
Therefore, the temperature of said burner face provides a rapid measure of the quantity of anode off gas that exits the fuel cell stack.

The quantity of hydrogen contained in the anode feed gas that is consumed in the fuel cell stack is proportional to the quantity of electrical power 42 that is produced in the stack. It is desirable that the fuel cell stack be capable of delivering a varying quantity of electricity as dictated by the system electric load demand. Therefore, it is desirable that the fuel processor automatically responds to the variable electric load demand by producing a proportionately varying quantity of hydrogen that is sent to the fuel cell stack.
To accomplish this objective, temperature emitter 50 shown in Fig. 2 including a thermocouple or infrared thermometer to measure the temperature of the burner face. Burner face temperature controller 51 maintains a desired setpoint temperature at the burner face by varying the flow rate of hydrocarbon feed at flow control valve 52. The burner face setpoint temperature is typically maintained in the range of 1300°F to 1900°F. For example, ifthe electric load increases, the quantity of hydrogen consumed in the fuel cell stack increases and the quantity of anode off gas sent to the infrared radiant burner decreases. This results in a rapid reduction in the burner face temperature. The control system of the present invention will respond by increasing the flow rate of hydrocarbon feed as needed to increase the hydrogen production rate and thus the flow of anode offgas that is sent to the infrared burner in order to maintain the desired burner face setpoint temperature.
It is desirable to maintain a relatively constant temperature at exit 10 of tubular reactor 8.
Since the absorbed duty in the tubular reactor varies in proportion to the quantity of hydrocarbon feed, it is desirable to vary the quantity of anode offgas that is combusted in the infrared burner in order to match said absorbed duty in the tubular reactor. Because the temperature of the burner face can be proportional to the quantity of anode offgas that is combusted within the infrared radiant burner, modulation of the burner face set point temperature provides a means to control the absorbed duty in the tubular reactor.
Temperature emitter 53 monitors the exit temperature of the tubular reactor.
Tubular reactor exit temperature controller 54 resets the setpoint temperature of burner face temperature controller 51. For instance, if the tubular reactor exit temperature is above the desired setpoint temperature, the tubular reactor exit temperature controller will lower the setpoint temperature of the burner face temperature controller. This results in a reduction in the flow of anode offgas to the infrared burner by modulation of the hydrocarbon feed flow control valve.
The control system of the present invention provides a means for rapidly modulating the production of hydrogen fed to a fuel cell stack in response to rapid changes in the electric load demand. The present invention also provides means for trim control to maintain a nearly constant temperature at the exit of the tubular reactor over wide load variations.
Example of Control System Set Point Change Time (min) Tubular Reactor Burner Face Exit Temperature Temperature (min) (°F) (°F) T + 7 1230 1750 T + 8 1229 1670 Table 1 shows that the steam reformer tubular reactor exit temperature increased slowly from 1200°F
to 1230°F over a 7 minute period. When it reached 1230°F, the set point for the burner face temperature was automatically decreased from 1750°F to 1670°F by the control system. The table shows that as a consequence of the change in the burner face temperature, the tubular reactor exit temperature began to decrease. It is also apparent in this example that the change in the gas phase temperature at the burner face is much more rapid than the change in the temperature of the solid phase steel of which the tubular reactor exit is composed. The table demonstrates that the gas phase at the burner face has a much smaller thermal inertia than any solid phase temperature, such as the tubular reactor exit temperature or the temperature of a refractory in the combustion chamber.
Therefore the control system of the present invention, that is based on the burner face temperature is much more sensitive to fluctuations in temperatures caused by changes in the amount of anode off gas returned to the burner, that are caused by changes in demand for electrical power. Hence the control system of the present invention is more capable of responding to transients in electrical power demand than control systems based on measurement of solid phase temperatures, such as the temperature of a solid refractory material.
The endothermic catalytic reaction apparatus of the invention allows for the production of industrial gases from a hydrocarbon feedstock that is simultaneously compact, thermally efficient, has improved life expectancy and low pressure drop, and is particularly well suited for the small scale generation of useful gases for fuel cell applications in the range of 1 kW to 10 kW.
A compact burner chamber employing a radiant burner assembly is configured to more uniformly distribute radiant energy along the axial length of a tubular reaction chamber by orienting the arc of radiation predominantly in the direction of the cold-plane surface of the tubular endothermic reaction chamber.

In a preferred embodiment, the radiant burner assembly comprises a woven metal fiber attached to a support structure that permits the efflux of fuel and oxidant from the burner core to the outer surface of the metal fiber. The properties of the metal fiber stabilize the combustion in a shallow zone proximal to the outer surface of the metal fiber. The combustion reaction heats the metal fiber to incandescence and provides a source of radiant energy that is transferred to the reaction chamber. In another embodiment, the radiant burner assembly comprises a porous ceramic fiber burner that accomplishes the same object by serving as a radiant source of energy.
The metal fiber is typically made of an alloy containing principally iron;
chromium, and aluminum and smaller quantities of yttrium, silicon, and manganese having extended life at operating temperatures up to 2000°F. The metal fiber radiant burner offers advantages due to high durability against thermal and mechanical shock.
In a preferred embodiment, the tubular reaction chamber consists of a U-tube, sometimes referred to as a hair-pin tube, which is substantially filled with catalyst and which passes into and out of the combustion chamber. The radiant burner axis is vertically disposed within the combustion 1 S chamber and is oriented parallel to the axis of the U-tube reaction chamber. The active radiant surface of the cylindrical radiant burner assembly is defined by a geometric arc that bisects the cylindrical assembly so as to maximize the flux of radiant energy that is directed to the surface of the U-tube reaction chamber. This embodiment minimizes the radiant exchange between the radiant burner assembly and the adiabatic refractory surfaces that enclose the combustion chamber, thereby reducing the radiant burner surface temperature for a given combustion intensity and reducing the tendency for flashback of the premixed fuel and oxidant contained within the radiant burner core.

In this embodiment, the center to center spacing beriveen the U-tubes that form the reaction chamber, the center to center spacing between the radiant burner and the U-tube reaction chamber, and the radiation angle of the radiant burner are simultaneously controlled so that the direct radiant flux from the burner that bisects the projected surface of the reaction chamber tube wall is a minimum of 50% of the total radiation flux that emanates from the active radiant burner surface.
In another preferred embodiment, the tubular reaction chamber consists of a helical coil that is substantially filled with catalyst and has inlet and outlet means that pass into and out of the combustion chamber. The helical coil is wrapped at specific angles so that the coil free area is in the range of 50% to 75%, wherein the free area is defined by the ratio of the free area between helical tube conduits and the cylindrical surface that bisects the helical coil circle. The radiant burner axis is vertically disposed within the combustion chamber and the cylindrical radiant burner is located at the center of the helical coil. In this embodiment, the active radiant surface of the cylindrical radiant burner assembly is defined by a 360-degree arc.
In both preferred embodiments, the radiant burner is operated at a combustion intensity and an excess air ratio that is carefully controlled to limit the radiant burner surface temperature to less than 2000°F, and preferably in the range of 1500°F to 1800°F, in order to provide extended life for the radiant burner, and also to limit the temperature of the premixed fuel and oxidant within the core of the radiant burner to less than 600°F, and preferably less than 500°F, in order to avoid pre-ignition or flashback of the premixed fuel contained in the burner core.
In both preferred embodiments, the catalyst particle diameters and reactant mass velocities are carefully controlled to simultaneously limit the reactor pressure drop to less than 8 prig, and preferably in the range of 2 psig to 4 psig in order that hydrocarbon feeds available only at low pressure can be used without compression, and to limit the reaction chamber tube wall temperatures to less than 1600°F, and preferably in the range of 1300°F to 1550°F, in order to allow extended life of the tube using relatively inexpensive tube alloys.
The upper portion of the combustion chamber is configured to form a convective chamber to enhance heat transfer from the combustion products to the tubular reaction chamber.
Refernng to Fig. 3 catalytic reaction apparatus 60 includes several primary chambers including a combustion chamber 61, convection chamber 62, and a reaction chamber 63. The combustion chamber 61 is defined by the volume enclosed by refractory insulation 64. The reaction chamber 63 is defined by the volume enclosed by a tubular reactor conduit 65.
The tubular reactor conduit 65 is formed in a U-tube or hairpin configuration and can be removed from the combustion chamber through a top flange 66. The tubular reactor conduit 65 passes concentrically through a convection chamber 62 defined by the space enclosed between the convection conduit 67 and the tubular reactor conduit 65. The reaction chamber is packed with catalyst from the inlet means 68, where reactants enter, to the outlet means 69 where products exit.
An axially extending, vertically disposed radiant burner 70 is supported by a burner gas conduit 71 that conveys a mixture of fuel and oxidant from an inlet means 72 to the radiant burner.
In this embodiment, the radiant burner 70 comprises a gas permeable metal fiber zone 73 and a non-permeable zone 74. Fuel and oxidant pass through the permeable metal fiber zone 73 where they are ignited on the surface thereby combusting and releasing heat to form an incandescent zone that radiates energy outward in an arc 75. The arc angle is controlled by design so that the radiating pattern maximizes the flux of radiant energy to the surface of the tubular reactor conduit 65 while minimizing the flux of radiant energy to the internal wall 76 combustion chamber 61. Fuel and oxidant are initially ignited on the surface of the permeable metal fiber zone 73 using an ignitor 77.
Once ignited, the combustion reaction on the surface of the metal fiber zone 73 is self sustaining.
S Specifically, the radiant arc angle 75 is controlled so that the direct radiant flux from the burner that bisects the projected surface of the reaction chamber tube wall is a minimum of 50% of the total radiation flux that emanates from the active radiant burner surface.
As an illustration of this teaching, FIG. 4 depicts a geometric representation of the preferred embodiment of the present invention. The active radiant zone 73 emits radiation along a line of sight defined by a radiant arc 75 that impinges on the reaction chamber conduits 65 and the inner surface 76 of the combustion chamber. The emitted radiation is bisected by a hypothetical plane passing through the centerline of the U-tube reaction chamber. The projected area of the reaction chamber surfaces per unit tube length receiving direct radiation from the burner within the controlled radiant arc is given by a + a = 2a. The total radiation within the arc falling on the hypothetical plane is given by 1 S c + c + a + a + b = 2c + 2a + b. In the preferred embodiment of the present invention, the ratio of 2a divided by 2c + 2a + b is typically greater than 0.5, or 50%.
In the present invention, the radiant burner combustion intensity is controlled in the range of 150,000 Btu/ft2/h and 350,000 Btu/ftZ/h and the excess combustion air is controlled in the range of 30% to 100% to prevent overheating of the surface of the radiant burner and to prevent overheating of the premixed fuel and oxidant contained within the burner core.
In the present invention, the reactant mass velocity is controlled in the range of 400 lb/ftZ/h to 1500 lb/ftz/h in order to limit the reaction chamber tube wall temperature to the desired range of 1300 °F to 1550 °F.
Combustion products emanating from the permeable metal fiber zone 73 enter an inlet means 77 leading to the convection chamber 62 wherein the combustion products exchange heat with tubular reaction chamber 65. The combustion products exit the convection chamber at an outlet means 78.
DETAILED DESCRIPTION OF PREFERRED EMBODIMENT
Example 1 A compact endothermic catalytic reaction apparatus according to the preferred embodiment was constructed and tested. The reaction chamber consisted of a 1" schedule 40 pipe constructed of 310 stainless steel that was formed in a U-tube arrangement spaced on 3"
centers. The reaction chamber was packed with a commercial steam reforming catalyst that was crushed and screened to an average size of'/4" x 3/s".
1 S The radiant burner consisted of a 4" long by 1 %2" outer diameter cylindrical assembly that had an active radiant angle of 120 degrees. The burner assembly was placed in an insulated combustion chamber having dimensions of 6" internal diameter and 10" height.
The radiant burner assembly was spaced approximately 4" from the U-tube centerline. The convection chamber consisted of a 2" tube constructed of 304 stainless steel.
The radiant burner was fired using a mixture of propane and air at a total higher heating value firing rate of 12,000 Btu/h. The reactant mixture consisted of 1 lb/h of propane and approximately 3.5 lb/h of steam and was fed to the reaction chamber at a temperature of approximately 800 °F. The reactant mixture was heated in the reaction chamber to an exit temperature of 1150 °F. The measured tube wall temperature of the reaction chamber was 1300 °F, the radiant burner surface temperature was 1650 °F, and the combustion products exit temperature was 1100 °F. The estimated hydrogen plus carbon monoxide yield was 67 SCFH.
FIG. 5 depicts an alternative embodiment of the present invention. In this embodiment, the reaction chamber is defined by the volume enclosed by a tubular reactor conduit comprising an upper section 81 consisting of a vertically disposed tube that is connected to the inlet means 82, a lower section 83 consisting of a helical coil, and an exit section 84 consisting of a vertically disposed tube that is connected to an exit means 85. The upper section 81 of tubular reactor conduit passes concentrically through the convection chamber 86. The reaction chamber is packed with catalyst from the inlet means 82, where reactants enter, to the outlet zone 87 of the lower section 83.
An axially extending radiant burner 88 is vertically disposed along the central axis of the helical coil section 83 of the tubular reaction conduit. The radiant burner is supported by a burner 1 S gas conduit 89 that conveys a mixture of fuel and oxidant from and inlet means 90 to the radiant burner. In this embodiment, the radiant burner 88 comprises a gas permeable metal fiber zone that surrounds the entire circumference of the radiant burner. Fuel and oxidant pass through the permeable metal fiber zone where they are ignited on the surface thereby combusting and releasing heat to form an incandescent zone that radiates energy in a predominantly uniform radial direction.
In order to maximize the amount of hydrogen that is produced, the maximum amount of carbon monoxide in the steam reformer exit process gas should be converted by the water gas shift reaction. Although decreasing the temperature will shift the reaction equilibrium to increase the theoretical yield of hydrogen, it will also decrease the rate of the reaction.
The use of successive stages for the water-gas shift reaction is well known in the prior art. The initial stage is performed at a temperature that will provide a rapid reaction rate, even though a substantial amount of carbon monoxide remains unconverted, as a result of the reaction equilibrium limitation. Subsequently the temperature is decreased so that the reaction equilibrium will permit the conversion of additional carbon monoxide, although that conversion will occur at a slower rate.
In a preferred embodiment, the successive stages of the water-gas shift reaction are performed in a single reaction vessel having unique means of heat exchange.
The heat exchange with the steam boiler permits the generation of steam while simultaneously decreasing the temperature of the exit gas from the first stage of the water gas shift reaction, prior to its processing in a subsequent stage at a decreased temperature. An electrical heater is the heat exchange means used to enhance the rate of heating the low temperature shift reactor during start-up.
Referring to Fig. 6 the single vessel containing successive stages of the water-gas shift reaction, 102 includes a high temperature shift reactor 109 and a low temperature shift reactor 104 that are separated by a steel bulkhead 103 in the interior of the vessel. A
suitable high temperature water-gas shift catalyst, composed of chromium, iron, and copper oxides is commercially available as G-3 C from Sud-Chemie. A suitable low temperature water-gas shift catalyst, composed of copper, zinc, and aluminum oxides is available as C 18-7 from Sud-Chemie. The process feed gas to the first stage of water gas shift enters the vessel through entrance 101 and flows through the bed of high temperature water-gas shift catalyst 109 and exits through heat exchange coil 107. Fitting 108 is joined to a corresponding fitting on the steam boiler, thereby permitting the exit gas from the first stage of the water-gas shift reactor to be cooled from its reaction temperature having a range of 600°F to 900°F preferably 600°F to 800°F to the reaction temperature of the subsequent water-gas shift reaction stage 350°F to 660° preferably 350°F to 550°F. The process gas flows through the bed of low temperature shift catalyst 104 and subsequently out of the vessel through exit 105. The electrical heater 106 is used during start-up to increase the temperature of the low temperature shift catalyst rapidly.
Illustration of Preferred Embodiment A single vessel containing two reactors for successive stages of the water gas shift reaction and an intermediate heat exchange coil between the two stages was constructed and tested. The single vessel consisted of a 3 inch schedule 40 pipe having a length of 21 inches that formed the inside diameter for both stages of the water gas shift reaction. A solid steel bulkhead divided the upper reaction chamber that was used for the first stage of the water gas shift reactor from the lower one that was used for the second stage, so that each chamber had the same length. The upper reaction chamber was filled a commercial high temperature shift catalyst that was supported on a 304 stainless steel screen located 7.5 inches from the top of the vessel.
After leaving the high temperature shift catalyst bed, the process gas flows through a U-shaped coil of %2 inch 304 stainless steel schedule 40 pipe through a steam generator and back to the lower reaction chamber. The %i inch coil was contained inside a 3 inch schedule 40 pipe that was welded to the side of the vessel. The 3 inch pipe on the side of the vessel had a flange that was bolted to the steam generator. The lower reaction chamber was filled with a low temperature shift catalyst. The process gas flowed out of the low temperature shift catalyst bed and exited from the vessel at the bottom of the vessel.
In a preferred embodiment the steam boiler that is a part of the present invention enhances the thermal efficiency of the fuel processor by generating steam through the combined utilization of the energy contents of the flue gas from the combustion chamber of the steam reformer, the process gas at the exit of the first stage of the water-gas shift reactor, and the process gas at the exit of the final stage of the water gas shift reactor, plus the utilization of electrical energy for start-up at times when the boiler is cold.
Refernng to Fig. 7 the steam boiler 122 has an inlet for the supply of boiler feed water 126 that is converted into steam that exits through boiler outlet 121. Inside the boiler, steam can be generated from heat supplied by four different sources. Flue gas from the combustion chamber of the steam reformer flows through coil 125. Process gas from the exit of the first stage of the water gas shift reaction flows through the coil shown in Fig. 6 that enters the steam boiler through opening 128 in Fig. 7 that is connected to the water-gas shift vessel by fitting 127 being connected to the corresponding flange of the water-gas shift reaction vessel. Process gas from the exit of the final stage of the water-gas shift reaction flows through coil 123. Electrical energy is converted to heat by immersion heater 124 for rapid start-up when the boiler has cooled below its operating temp erature.
Illustration of Preferred Embodiment A steam generator corresponding to the preferred embodiment was constructed and tested. It consisted of a 3-inch schedule 40 pipe. A 150 pound blind flange was welded to the bottom of the pipe. A hemispherical cap was welded to the top. The boiler feedwater flowed in a %z inch tube side port at the bottom of the generator. The steam that was generated flowed out of the center of the cap at the top. A fitting on the side of the steam generator was used for the connection that permitted the coil from the high temperature water gas shift chamber to enter the steam generator. Entrance and exit ports were located on the blind flange at the bottom of the steam generator for the process gas from the low temperature water gas shift chamber and the flue gas from the combustion chamber of the reformer. A port for the electrical immersion heater was also located on the blind flange at the bottom of the reactor. A port for water blow-down was located on the side of the steam generator. Ports for instruments that measured temperature, pressure, and level were also located on the side of the steam generator.
Because carbon monoxide is a poison for the platinum catalysts used in fuel cell stacks, the carbon monoxide concentration at the exit of the last stage of the water-gas shift reaction must be decreased to less than 10 ppm. It is desirable to oxidize the carbon monoxide preferentially to carbon dioxide with none of the hydrogen being oxidized to water. As the temperature increases 1 S there is a tendency for the reaction to become progressively less preferential and more and more of the hydrogen to become oxidized. Because the amount of oxygen in the form of air that is added to the preferential oxidation reactor is limited, it is essential that not too much of it is used by the reaction with hydrogen. If too much of the oxygen reacts with hydrogen for the undesirable conversion to water there will be insufficient remaining for the desirable conversion of carbon ?0 monoxide to carbon dioxide.

Since both of these oxidation reactions are exothermic, there is a tendency for the temperature to increase as the reaction occurs. Therefore a means of removing the exothermic heat of reaction from the preferential oxidation reactor is required to prevent the temperature from increasing excessively and causing the reaction to become non-preferential.
The heat removal must be controlled carefully. If too much heat is removed, parts of the reactor can become cold and the desirable conversion of carbon monoxide to carbon dioxide will not occur at a sufficient rate to obtain the 10 ppm carbon monoxide requirement.
In a preferred embodiment, heat is removed from the first part of the preferential oxidation reactor and not from the second part. In the first part of the reactor the concentrations of oxygen and carbon monoxide are the greatest. Therefore most of the conversion occurs in the first part and therefore most of the heat of reaction is produced in the first part of the reactor. However, the concentration of carbon monoxide must be decreased by at least t<vo orders of magnitude. By not cooling the second part of the reactor an appropriate temperature can be maintained to ensure that the preferential reaction will continue.
Referring to Fig. 8 the preferential oxidation reactor is composed of an inlet header 142 to which are connected the inlets of several reaction tubes 148 filled with catalyst. The exits of the reaction tubes 148 are connected to an exit header 147. The catalyst contained~platinum on a support, having a composition that is known to practitioners of the art, US
Patent 5039646. On the inlet header 142 diametrically opposed to each reaction tube 148 there is a port 141 for loading and discharging catalyst. The upper portion of the reaction tube has metal fins 143 attached to it. Their function is to enhance the removal of the exothermic heat of the oxidation reaction. A fan 145 draws air through an inlet grid 149 and forces it out through an exit grid 146 of a chamber 144 that houses the reaction tubes.
Illustration of preferred embodiment A preferential oxidation reactor was constructed and tested. The inlet and outlet headers were constructed from'/4 inch stainless steel tubing. The tops of six %z inch stainless steel tubes wer a welded at a 90-degree angle to the inlet header. The bottoms of the tubes were welded to the outlet header. Each of the tubes was filled with a commercial preferential oxidation catalyst to a cata lyst bed depth of 12 inches. Cooling fins were attached to the exterior surface of the upper SO percent of the %2 inch reaction tubes. A fan having a maximum capacity of 530 cubic feet per hour was used to draw air over the cooling fins of the upper part of the preferential oxidation reaction tubes.

Claims (58)

1. Endothermic catalytic reaction apparatus comprising:
(a) a hairpin tubular reaction chamber vertically disposed within a combustion chamber and packed with a catalyst for the conversion of hydrocarbon to industrial gases by reaction with steam; said hairpin tubular reaction chamber having an upper portion that is surrounded by a convection chamber to enhance the transfer of heat from combustion products to the reaction chamber.

(b) a radiant burner vertically disposed within a combustion chamber having a gas permeable zone that promotes the flameless combustion of fuel and oxidant in order to heat the metal fiber surface to incandescence and radiate energy to the reaction chamber;
said radiant burner having a controlled geometric design so that the angle of radiation is predominantly incident upon the tubular reaction chamber surface.

2. The apparatus of claim 1, a hairpin tubular reaction chamber comprising a tube having outer diameters ranging from 3/8" to 4", preferably from 3/4" to 2.5".

3. The apparatus of claim 1, a hairpin tubular reaction chamber having mass velocities ranging from 200 lb/ft2/h to 3000 lb/ft2/h, preferably from 400 lb/ft2/h to 1500 lb/ft2/h.

4. The apparatus of claim 1, a hairpin tubular reaction chamber having average catalyst particle diameters ranging from 1/16" to 3/4", preferably from 1/4" to 1/2" and pressure drops ranging from 1 psi to 8 psi.

5. The apparatus of claim 1, a hairpin tubular reaction chamber having an inlet operating pressures ranging from 0 psig to 20 psig and preferably from 2 psig to 10 psig.

6. The apparatus of claim 1, a hairpin tubular reaction chamber having an exit temperature ranging from 1000 ÀF to 1700 ÀF and preferably from 1200 ÀF to 1400 ÀF.

7. The apparatus of claim 1, a hairpin tubular reaction chamber having maximum tube wall temperatures that range from 1200 ÀF to 2000 ÀF and preferably from 1400 ÀF to 1600 ÀF.

8. The apparatus of claim 1, a hairpin tubular reaction chamber having average heat fluxes ranging from 2,000 Btu/ft2/h to 20,000 Btu/ft2/h and preferably from 5,000 Btu/ft2/h to 10,000 Btu/ft2/h.

9. The apparatus of claim 1, a hairpin tubular reaction chamber having capacity to generate hydrogen plus carbon monoxide product in volumetric quantities ranging from 20 SCFH to 1000 SCFH and preferably from 25 SCFH to 600 SCFH, and particularly well suited for generating hydrogen plus carbon monoxide in the range of 25 SCFH to 150 SCFH.

10. The apparatus of claim 1, a radiant burner comprising a supported porous ceramic material having extended life at operating temperatures up to 2100 ÀF.

11. The apparatus of claim 1, a radiant burner comprising a supported metal fiber material typically fabricated from an alloy containing principally iron, chromium, and aluminum and smaller quantities of yttrium, silicon, and manganese having extended life at operating temperatures up to 2000 ÀF.

12. The apparatus of claim 1, a radiant burner having a controlled design such that the angle of radiation ranges from 45 to 180 degrees.

13. The apparatus of claim 1, a radiant burner having surface temperatures ranging from 1700 ÀF to 2000 ÀF and preferably 1700 ÀF to 2000 ÀF.

14. The apparatus of claim 1, a radiant burner that limits the internal temperatures of the premixed fuel and oxidant to less than 600 ÀF and preferably less than 500 ÀF
in order to prevent flashback, wherein flashback is defined as the pre-combustion of the fuel and oxidant mixture within the burner feed pipe.

15. The apparatus of claim 1, a radiant burner that is insulated to prevent the transfer of heat to the premixed fuel gases from the non-permeable side of the radiant burner.

16. The apparatus of claim 1, a radiant burner that operates at a combustion intensity typically ranging from 100,000 Btu/ft2/h to 500,000 Btu/ft2/h and preferably from 150,000 Btu/ft2/h to 350,000 Btu/ft2/h, wherein the combustion intensity is defined as the higher heating value of the fuel combusted divided by the permeable radiant burner surface area.

17. The apparatus of claim 1, a radiant burner that operates at an excess air ratio typically ranging from 10% to 200% and preferably from 30% to 100%, wherein the excess air ratio is defined as percent combustion air in excess of the stoichiometric amount required for complete combustion of the burner fuel.

18. Endothermic catalytic reaction apparatus comprising:
(a) A helical tubular reaction chamber disposed within a combustion chamber and packed with a catalyst for the conversion of hydrocarbon to industrial gases by reaction with steam, said helical tubular reaction chamber having an upper portion that is surrounded by a convection chamber to enhance the transfer of heat from combustion products to the reaction chamber and an exit section to convey reaction products to the exit means.

(b) A radiant burner vertically disposed within a combustion chamber having a gas permeable zone that promotes the flameless combustion of fuel and oxidant in order to heat the metal fiber surface to incandescence and radiate energy to the reaction chamber;
said radiant burner designed to radiate uniformly in the radial direction.

19. The apparatus of claim 18, a helical tubular reaction chamber comprising a tube having outer diameters ranging from 3/8" to 4" and preferably from 3/4" to 2.5".

20. The apparatus of claim 18, a helical tubular reaction chamber having an outer coil diameter ranging from 2" to 24" and preferably from 5" to 18".

21. The apparatus of claim 18, a helical tubular reaction chamber having mass velocities ranging from 200 lb/ft2/h to 2000 lb/ft2/h and preferably from 400 lb/ft2/h to 1500 lb/ft2/h.

22. The apparatus of claim 18, a helical tubular reaction chamber having average catalyst particle diameters ranging from 1/16" to 3/4"and preferably 1/4" to 1/2" and pressure drops ranging from 1 psi to 8 psi.

23. The apparatus of claim 18, a helical tubular reaction chamber having an inlet operating pressures ranging from 0 psig to 20 psig and preferably from 2 psig to 10 psig.

24. The apparatus of claim 18, a helical tubular reaction chamber having an exit temperature ranging from 1000 ÀF to 1700 ÀF and preferably from 1200 ÀF to 1400 ÀF.

25. The apparatus of 18, a helical tubular reaction chamber having maximum tube wall temperatures ranging from 1200 ÀF to 1700 ÀF and preferably from 1400 ÀF to 1600 ÀF.

26. The apparatus of claim 18, a helical tubular reaction chamber having average heat fluxes ranging from 2000 Btu/ft2/h to 20,000 Btu/ft2/h and preferably from 5,000 Btu/ft2/h to 10,000 Btu/ft2/h.

27. The apparatus of claim 18, a helical tubular reaction chamber having capacity to generate hydrogen plus carbon monoxide product in volumetric quantities ranging from 20 SCFH to 1000 SCFH and preferably from 25 SCFH to 600 SCFH, and particularly well suited for generating hydrogen plus carbon monoxide in the range of 100 SCFH to 300 SCFH.

28. The apparatus of claim 18, a radiant burner comprising a supported porous ceramic material having extended life at operating temperatures up to 2100 ÀF.

29. The apparatus of claim 18, a radiant burner comprising a supported metal fiber material typically fabricated from an alloy containing principally iron, chromium, and aluminum and smaller quantities of yttrium, silicon, and manganese having extended life at operating temperatures up to 2000 ÀF.

30. The apparatus of claim 18, a radiant burner having a controlled design such that the radiant energy emanates from the burner in a substantially uniform radial pattern.

31. The apparatus of claim 18, a radiant burner having surface temperatures ranging between 1500 ÀF and 2000 ÀF and preferably 1700 ÀF and 2000 ÀF.

32. The apparatus of 18, a radiant burner that limits the internal temperatures of the premixed fuel and oxidant to less than 600 ÀF and preferably less than 500 ÀF in order to prevent flashback, wherein flashback is defined as the pre-combustion of the fuel and oxidant mixture within the burner feed pipe.

33. The apparatus of claim 1, a radiant burner that operates at a combustion intensity typically ranging from 100,000 Btu/ft2/h and 500,000 Btu/ft2/h from 150,000 Btu/ft2/h to 350,000 Btu/ft2/hr, wherein the combustion intensity is defined as the higher heating value of the fuel combusted divided by the permeable radiant burner surface area.

34. The apparatus of claim 1, a radiant burner that operates at an excess air ratio typically ranging from 10% to 200% and preferably from 30% to 100%, wherein the excess air ratio is defined as percent combustion air in excess of the stoichiometric amount required for complete combustion of the burner fuel.

35. A fuel processor comprising a desulfurizer vessel, an endothermic catalytic reaction apparatus comprising a burner, a combustion chamber, a steam reformer reactor, and including a method of heating and control of said apparatus, a steam generator that is heat integrated with several different reactors within said fuel processor, a single vessel for successive stages of the water-gas shift reaction vessel that is heat integrated with the steam generator, a compact preferential oxidation reactor capable of decreasing the carbon monoxide concentration to less than 10 ppm, a control system for the entire fuel processor that permits it to produce the quantity of hydrogen required by an integrated polymer electrode membrane fuel cell stack when said fuel cell stack is following changes in electrical load, and a system for the separation and recycle of water;
(a) said endothermic tubular reaction apparatus comprising a hairpin tubular reaction chamber vertically disposed within a combustion chamber and packed with a catalyst for the conversion of hydrocarbon to industrial gases by reaction with steam; said hairpin tubular reaction chamber having an upper portion that is surrounded by a convection chamber to enhance the transfer of heat from combustion products to the reaction chamber.
(b) a radiant burner vertically disposed within a combustion chamber having a gas permeable zone that promotes the flameless combustion of fuel and oxidant in order to heat the metal fiber surface to incandescence and radiate energy to the reaction chamber;
said radiant burner having a controlled geometric design so that the angle of radiation is predominantly incident upon the tubular reaction chamber surface.

36. The fuel processor of claim 35 where said steam generator that is heated by a combination of three sources in the form of process streams that provide recovered energy from the steam reformer and water gas shift reaction chambers plus a fourth energy source in the form of electricity that is transformed into heat by an immersion heater inside the steam generator.

37. The fuel processor of claim 36 an electrical immersion heater having a capacity from 0.1 kW
to 2 kW preferably 0.5 kW to 1.5 kW.

38. The fuel processor of claim 36 where one of the three process gas streams used to provide energy is the flue gas stream from the combustion chamber of the steam reformer.

39. The fuel processor of claim 36 where one of the three process gas streams used to provide energy is exit gas from the bed of high temperature water-gas shift catalyst that exits from the middle of the water gas shift reaction vessel flows through the steam generator to transfer heat and decrease its temperature, and then re-enters the water-gas shift reaction vessel at its mid-point to flow down over the low temperature water-gas shift catalyst.

40. The fuel processor of claim 36 where one of the three process gas streams used to provide energy is the exit gas from the bed of low temperature water-gas shift catalyst that flows out of the bottom of the water-gas shift reaction vessel.

41. The fuel processor of claim 35 where said water-gas shift reaction vessel that is heat integrated with the steam generator where successive stages of the water-gas shift reaction are performed at different temperatures each in separate chambers within a single reaction vessel.

42. The fuel processor of claim 41 where the first stage of the water-gas shift reaction is performed in a chamber of the water-gas shift reaction vessel that contains a commercial high temperature shift catalyst composed of chromium, iron, and copper oxides;
at a temperature from 600ÀF to 900ÀF and preferably from 600ÀF to 800ÀF; at a gas hourly space velocity from 200 h-1 to 10,000 h-1 and preferably from 1000 h-1 to 8000 h-1;
and at steam to carbon ratios expressed as moles of water per atom of carbon from 1.8 to 6 and preferably from 2 to 4.

43. The fuel processor of claim 41 where the process gas exiting from the first chamber of the water-gas shift reaction vessel flows through a coil that passes through a steam generator, thereby decreasing its temperature to that desired for the subsequent stage of the water-gas reaction and transferring the corresponding amount of energy to produce steam in the steam generator.

44. The fuel processor of claim 41 where the subsequent stage of the water-gas shift reaction is performed in a chamber of the water-gas shift reaction vessel that contains a commercial low temperature shift catalyst composed of copper, zinc, and aluminum oxides at a temperature from 350ÀF to 660ÀF and preferably 350ÀF to 550ÀF, at a gas hourly space velocity from 200 h-1 to 10,000 h-1 and preferably from 1000 h-1 to 8000 h-1; and at steam to carbon ratios, expressed as moles of water per atom of carbon, from 1.8 to 6 and preferably from 2 to 4.

45. The fuel processor of claim 35 where said preferential oxidation reactor is capable of decreasing the concentration of carbon monoxide to less than 10 ppm both at steady-state operation and during transient operating conditions between different steady-states.

46. The fuel processor of claim 45 where said preferential oxidation reactor is a series of tubular reactors in parallel.

47. The fuel processor of claim 45 where said preferential oxidation reactor has heat transfer fins on the 50 percent of its length closest to the feed inlet.

48. The fuel processor of claim 45 where said preferential oxidation reactor includes a fan to enhance the removal of the exothermic heat of reaction by the mechanism of forced convection resulting from air passing over the reactor.

49. The fuel processor of claim 45 where said preferential oxidation reactor is operated at temperatures from 80ÀF to 200ÀF preferably 100ÀF to 180ÀF.

50. The fuel processor of claim 45 where said preferential oxidation reactor is operated at gas hourly space velocities from 1000 h-1 to 20,000 h-1 preferably 2000 h-1 to 10,000 h-1.

51. The fuel processor of claim 45 where said preferential oxidation reactor is operated at excess air stoichiometries of 80% to 400 percent preferably 100 percent to 300 percent, where excess air stoichiometry is defined as the ratio of amount of oxygen fed to the reactor divided by the amount of oxygen required to convert all of the carbon monoxide.

52. The fuel processor of claim 35 where the desulfurizer is a vessel filled with an adsorbent consisting of an activated carbon and an activated carbon that has been promoted with copper.

53. The fuel processor of claim 35 where said control system is based on the temperature of the gas phase burner face temperature in the combustion chamber of said endothermic catalytic reaction apparatus.

54. The fuel processor of claim 53 where an increase in the burner face temperature above its set point causes the control system to decrease the amount of hydrocarbon that is fed to the catalytic steam reformer reactor that is part of said endothermic catalytic reaction apparatus, and conversely a decrease in said temperature below its set point causes an increase in the amount of hydrocarbon fed to the catalytic steam reformer reactor.

55. The fuel processor of claim 53 where the set point for the burner face temperature is decreased by the control system if the temperature of the reformer exit temperature exceeds a predetermined value and conversely is increased by control system if the temperature of the reformer exit temperature becomes less than the predetermined value.

56. The fuel processor of claim 35 where said system for the separation and recycle of water consists of a vertically mounted separator vessel that allows the dried gas to flow out of the top of the separator to the preferential oxidation reactor and the liquid phase water to flow out of the bottom to a pump suction that sends it to the steam generator.

57. The fuel processor of claim 35 where the feedstock is any hydrocarbon or mixture of hydrocarbons including natural gas, liquefied petroleum gas, gasoline, diesel fuel.

58. The fuel processor of claim 35 where the hydrocarbon is desulfurized in a desulfurizer, combined with steam, heated in a feed-effluent heat exchanger, heat further by exchange with flue gas from the combustion chamber, passed over a fixed bed of catalyst inside a combustion chamber to be converted to hydrogen and carbon monoxide, cooled in a feed-effluent heat exchanger, passed over a fixed bed of high temperature water-gas shift catalyst to convert some of the carbon monoxide to hydrogen, by flowing through a coil that passes through a steam generator, passed over a fixed bed of low temperature water-gas shift catalyst to convert more of the carbon monoxide to hydrogen, passed through another coil inside a steam generator where it transferred heat to generate steam, passed through an air cooler to decrease its temperature to cause water to condense in a separator, mixed with air and passed through a fixed bed of preferential oxidation catalyst where the remaining carbon monoxide is converted so that the hydrogen stream contains less than 10 ppm of carbon monoxide and passed into a fuel cell stack where most of the hydrogen is consumed in the generation of electricity.