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Publication numberUS3828474 A
Publication typeGrant
Publication dateAug 13, 1974
Filing dateFeb 1, 1973
Priority dateFeb 1, 1973
Also published asCA1004466A, CA1004466A1, DE2342085A1
Publication numberUS 3828474 A, US 3828474A, US-A-3828474, US3828474 A, US3828474A
InventorsO Quartulli
Original AssigneePullman Inc
Export CitationBiBTeX, EndNote, RefMan
External Links: USPTO, USPTO Assignment, Espacenet
Process for producing high strength reducing gas
US 3828474 A
Abstract
This invention provides a process for producing a high strength reducing gas suitable for reducing metallic ores such as iron ore. The process is a multi-step process using a C3 to C15 hydrocarbon such as liquid naphtha as the starting material.
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ilnited States Patent [191 Quartulli [451 Aug. 13, 1974 PROCESS FOR PRODUCING HIGH STRENGTH REDUCING GAS [75] Inventor: Orlando J. Quartulli, London,

England [73] Assignee: Pullman Incorporated, Chicago, 111. [22] Filed: Feb. 1, 1973 [21] Appl. No.: 328,805

[52] US. Cl. 48/214, 48/196 R, 48/213, 252/373 [51] Int. Cl C01b 2/14 [58] Field of Search 48/197 R, 196 R, 214, 213, 48/211; 252/373 [56] ReferencesCited UNITED STATES PATENTS 3,071,453 l/l963 James 48/196 3,119,667 1/1964 McMahon 23/212 3,264,066 8/1966 Quartulli et al.. 23/212 3,388,074 6/1968 Reitmeier 252/373 3,441,395 4/1969 Dent 48/214 3,475,160 10/1969 l-lcinzelmann et a1. 75/26 3,642,460 2/1972 Thompson 48/214 f/YJ/V J a j n (a. n l

4 s! iii/.1 l l WZWID fli /W774 FOREIGN PATENTS OR APPLICATIONS 1,195,428 11/1967 Great Britain 48/214 820,257 9/1959 Great Britain 48/214 Primary Examiner-S. Leon Bashore Assistant Examiner-Peter F. Kratz [5 7] ABSTRACT This invention provides a process for producing a high strength reducing gas suitable for reducing metallic ores such as iron ore. The process is a multi-step process using a C to C hydrocarbon such as liquid naphtha as the starting material.

The first step of the process comprises gasifying the hydrocarbon by passing a preheated mixture of the hydrocarbon and steam through a bed of a reforming catalyst to produce a gas consisting essentially of methane, hydrogen, carbon oxides and steam. Carbon dioxide is then removed from this gas mixture and the resulting gas is'steam reformed in the presence of a reforming catalyst to produce reducing gas comprising hydrogen and carbon monoxide.

13 Claims, 3 Drawing Figures a FVZZ "/9 1 I1 1 :1" :l re I: I l l 1 W071 .l idea-war =0 P'DM6/A 6 7/ 645 MQQNM SHEET 3 BF 3 PATENTEU mm 3 2974 PROCESS FOR PRODUCING HIGH STRENGTH REDUCING GAS BACKGROUND OF THE INVENTION This invention is concerned with a process for producing a reducing gas suitable for reducing metallic ores such as iron ore.

Previous reforming processes for the production of high strength reducing gas for iron ore reduction utilize natural gas which is steam reformed in a single facility at various pressures and steam flows. With a natural gas feed it is possible to achieve H CO concentrations in the reformed gas of the order of 88 percent and greater by operation at relatively low steam input in conjunction with the use of a supported nickel catalyst.

It is well known in the art that to-obtain a high strength reducing gas without any cooling of the reformed gas or without removal of CO it is necessary to operate the reformer at a low ratio of steam per atom of carbon in the feed, otherwise known as the steamcarbon ratio. For instance, in the case of methane or natural gas, the reaction proceeds according to the following equation:

CH4 H2O c 3H,

(1) The stoichimetric requirement of steam per carbon atom for the above equation is 1.0. However, other reactions can occur within the reformer such as the carbon monoxide disproportionation reaction and the cracking reaction:

CH C+2H 3) Reactions (2) and (3) are associated with deposition of carbon on the catalyst which in turn causes deactivation of catalyst and, in-extreme cases, plugging of the catalyst bed. To avoid carbon deposition it is necessary to introduce large amounts of steam if a low activity reforming catalyst such as an unpromoted nickel catalyst is used, or to increase the steam-carbon ratio slightly above the stoichiometric level if ahigh activity nickel catalyst, such as an alkali metal promoted catalyst is used. In the former case steam-carbon ratios of from 2 to 3 to l are required depending on operating conditions. In the latter case, it is possible to operate the reformer at a relatively low steam-carbon ratio. For example, if it were assumed that reaction (1) is carried out at a steam-carbon ratio of 1.10 to 1.0.at conditions of 1840F and 25 psig., the concentration of H CO in the reducing gas would be of the order of 95 percent and better. A typical composition of the reducing gas at these conditions assuming a methane feed gas would be:

It will be noted that at conditions of 25 psig. and 1840F the effluent from the reformer consists largely of H C0, small amounts of CO and H 0, and residual methane, the amount of which is dependent upon the amount of nickel catalyst provided in the reformer.

If it is desired to operate the reformer in the same manner with higher molecular weight hydrocarbon feeds such as liquid naphtha, it has not been possible to achieve the high degree of H C0 obtainable with natural gas feeds. Because of the higher carbon number of the naphtha feed, it was necessary to introduce greater amounts of steam into the process in order to suppress carbon deposition as represented by equations (2) and (3). For example, with anon-promoted reforming catalyst, steam-carbon ratios of from 3 to 10 to 1.0 would be required depending on feedstock and operating pressure. However, even with a promoted reforming catalyst it is also necessary to use relatively larger amounts of steam albeit less than that required Qgc onventional nickel catalysts.

Despite operation at relatively low steam-carbon ratio in conjunction with a promoted nickel reforming catalyst, the content of H COin the reformed gas is only about 80 percent, which is generally considered to be unsuitable for a high efficiency iron ore reduction operation. The basic reason for the low H CO content of the reducing gas is the diluent effect of the additional reforming steam and associated C0 produced in the reforming operation. While removal of steam from .the reformer effluent would yield a high strength reducing gas, such a step would be uneconomic because it would be necessary to cool the reformer effluent for water condensation and removal after which it would have to be reheated prior to introduction into the iron ore reductionfacility. Cooling and reheating requires costly heat exchange equipment and associated large high temperature piping which would make the process unattractive. In addition, the water condensation step increases pressure drop and, thus, requires that the reformer operation be carried out at much higher pressure necessitating an increase in either the operating temperature or the steam-carbon ratio, or both, in order to meet the residual methane requirement.

The object of the present invention is to provide a multi-stage process wherein hydrocarbons having an average of from 3 to 15 carbon atoms can be converted into a high strength reducing gas suitable for reducing iron ore.

a process for producing a reducing gas from a hydrocarbon feedstock having an average of from 3 to 15 carbon atoms, which comprises the steps of:

a. gasifying the hydrocarbon by passing a preheated mixture of the hydrocarbon and steam through a bed of a reforming catalyst in a gasification reactor to produce an efi'luent gas comprising methane, hydrogen, carbon oxides and steam;

b. removing carbon dioxide from said effluent gasr and then According to the present invention, there is provided c. reforming the resulting gas in the presence of steam and a reforming catalyst to produce a gas containing generally from 85 to- 99 percent by volume hydrogen and carbon monoxide.

BRIEF DESCRIPTION or THE DRAWINGS FIGS. 1 to 3 are flow diagrams showing alternative embodiments of apparatus for carrying out the process of this invention.

DETAILED DESCRIPTION OF THE INVENTION Several reactions occur simultaneously in the gasification stage which is carried out as an adiabatic process. These include:

CnHm +n H O =n CO+ (n +m/2) H co+3H,=cH,+H,o

CO H2O CO2 H2 Other reactions including hydrocracking are known to occur in the gasification stage. Reactions (4) to (6) are the basic reactions occuring in the system. Reaction 4, conversion or gasification of the hydrocarbon to CO and hydrogen is endothermic. Reaction 5, covering methanation of CO, and reaction 6, the CO shift conversion, are exothermic. In the gasification reaction, steam is consumed in the process in the production of CO and hydrogen. However, as this reaction occurs, the highly exothermic reactions 5 and 6 proceed at a rapid rate with the net result that there is a rise in the temperature across the reactor. The gas so produced from the overall exothermic conversion consists essentially of methane, carbon dioxide, hydrogen and steam, with a small amount of carbon monoxide. The composition of the product gas from this gasification step is determined by the temperature, pressure, and steam ratio chosen for the equilibrium which involves a number of reactions notable of which are the aforementioned steam CH equation (1 and CO shift conversion reaction 6).

The formation of methane is favored by low temperature and high operating pressure as represented by the equilibrium expression defined by equation (I).

where CO, H CH and H denote molal flows of each component; ZN the total molal flow of wet gas; and P, the pressure at the outlet of the reactor in atmospheres absolute.

The concentration of CO and CO is influenced by the quantity of steam introduced into the system and the operating temperature which are represented by the equilibrium expression as defined by equation (6).

By removing the excess carbon in the form of CO from the effluent gas from the gasification stage it is possible to simulate a feed having essentially the same characteristics as a methane feed in terms of carbonhydrogen ratio and theoretical hydrogen which is easily reformable in a subsequent reformer operation. Removal of CO from the steam permits operation at essentially the same conditions as a natural gas reformer and thus, enables production of a reducing gas having high concentrations of H and CO.

The starting material for the process of the invention is a hydrocarbon fraction having an average of from 3 to 15 carbonatoms. Suitable hydrocarbons include liquid naphtha fractions, fractions rich in saturated hydrocarbons such as pentanes, hexanes and heptanes and similar light hydrocarbons including liquified propane gas and butane. Preferably the fractions are those having an end point of up to 350 420F.

The gasification step of the process comprises preheating a mixture of the hydrocarbon feedstock and steam, preferably in a steam/hydrocarbon weight ratio of at least 1.521, and passing the gaseous mixture through a bed of reforming catalyst to produce a methane rich gas. The steam/hydrocarbon ratio increases with the molecular weight of the hydrocarbon so that, for example, a suitable ratio for liquified propane gas is 1.611, for light naphtha is 1.8:1 and for heavy naphtha is 2.0:1. A suitable catalyst comprises nickel, cobalt, iron or mixtures thereof on a support, such as an unpromoted nickel catalyst or a supported nickel catalyst promoted with alkali metal such as sodium or potassium. Suitable supports include alumina, magnesia, zirconia and mixtures thereof. Such catalysts are disclosed in U.S. Pat. Nos. 3,567,411; 3,433,609 and 3,201,214.

The steam-hydrocarbon mixture is preferably introduced into the catalyst bed at a temperature of at least 660F. The reaction is exothermic and the catalyst bed is generally maintained by the heat of reaction at a temperature of from 750 to 1,020F. The reaction is preferably carried out at an elevated pressure, suitably of from psig. to 625 psig.

The gasification step may be carried out in a catalytic rich gas (CRG) reactor as developed by the Bristish Gas Council. Details of such a reactor are described in |.G.E. Journal, June 1969, pages 375 to 396. Additional processes for the production of methane rich gases are disclosed in Hydrocarbon Processing, April 1971, pages 97, 98, 99, 102, and 110.

The methane-rich gas resulting from the gasification step is treated for removal of carbon dioxide. This is suitably effected by a regenerative circulating solvent system with preferred solvents being 20 percent monoethanolamine or promoted or unpromoted potassium carbonate solution.

The gas stream after removal of carbon dioxide is then subjected to catalytic steam reforming in a reformer furnace. The gas stream mixed with additional steam is heated to the inlet temperature of the furnace and then introduced into a furnace such as tubular refonner which contains a reforming catalyst. Suitable catalysts are those already described with reference to the gasit'ication step but it is particularly preferred to use an alkali metal promoted nickel catalyst in the upper inlet zone of the furnace and an unpromoted nickel catalyst in the lower outlet zone. Processes for steam reforming to produce high strength reducing gas are described in detail in Belgium Patent No. 765,920 issued Oct. 19, 1971 and in a paper entitled HIGH GRADE REDUCING GAS FOR METALLURGICAL APPLICATIONS by Finneran et al, 30th. Iron Making Conference, A.I.M.E. April 20, 1971, Pittsburgh, Pa.

The preferred temperature conditions in the reformer are an inlet temperature of from 600 to 1,200F, preferably from 900 to 1,000F, and an outlet temperature of from 1,650 to 2,100F, preferably from l,750 to 1,850F. The pressure in the reformer is suitably from atmospheric to 250 psig., preferably from to 150 psig. The preferred steam to carbon ratio is in the range of from about 1.1 to about 1.4 mols. of steam per atom of carbon in the feed.

The gas leaving the reformer at the outlet temperature thereof is suitably delivereddirectly to an iron ore reduction plant. The reducing gas generally contains at least 88 percent of carbon monoxide and hydrogen.

Among the advantages of the process of this invention is that the reforming catalyst in the reformer furnace is maintained in a high state of activity due to the presence of hydrogen in the feed. High hydrocarbon conversion should therefore be maintained throughout the life of the catalyst. Furthermore, the danger of carbon deposition in the reformer is minimized by virtue of the relatively high hydrogen content of the inlet gas and the size of the reformer furnace is somewhat smaller than that required for natural gas reformers of equivalent capacity due to the lower reformer duty resulting from the presence of hydrogen in the inlet gas.

Referring to the accompanying drawings,

FIG. 1 is a flow diagram of apparatus for carrying out the process of this invention disclosing a fixed bed gasification reactor such as a Gas Council CRG reactor for the primary gasification stage. Purified (i.e., desulfurized) naphtha is delivered through conduit 1 to a heat exchanger 2 located in the effluent circuit of a CRG reactor 3. The partially heated stream passing through conduit 4 is mixed with steam supplied through conduit 5, which steam can be produced within the process. The combined steam-naphtha mixture flows through conduit 6 to a preheater furnace 7 where it is heated to the inlet conditions of the CRG reactor 3. If the naphtha feed is supplied as a vapor it is not necessary to provide the heat exchanger 2, in which case the combined steam-naphtha mixture flows directly to the preheater furnace 7. In the CRG reactor 3 the naphtha is completely converted to a methane rich gas comprising a mixture of hydrogen, methane and carbon oxides in the presence of a steam-reforming catalyst. The gas mixture is normally in equilibrium with the unreacted steam at the outlet temperature of the reactor. Examples of commercially available catalysts particularly suited for the CRG gasification stage include British Gas Council Type A and B CRG nickel catalysts.

The gas mixture from CRG reactor 3 then flows through conduit 9 to the preheat exchange 2 giving up a portion of the heat to the incoming naphtha feed,

after which it flows to heat recovery equipment 10 which may comprise a waste heat boiler, a reboiler for supplying heat to the CO stripper of the CO removal system, or any-suitable system for recovery of waste heat. Following vapor-liquid separation in separator 11, the disengaged methane rich gas flows to a C0 absorber 12 where it is contacted countercurrently with a lean regenerative solvent introduced through inlet 13 for removal of the bulk of the CO in the stream. CO rich solvent leaves the absorber 12 at the bottom outlet 14 and flows to a solvent regeneration system (not shown). Alternate prior art CO removal processes are disclosed in Hydrocarbon Processing, April 1971, pages 96, 101, 103, 104, and 117.

The overhead from the CO absorber, consisting mainly of methane, hydrogen and small amounts of carbon monoxide, then flows through conduit 15 to the convection section of a tubular reformer furnace 16 where it is preheated to an intermediate temperature after which it is joined with steam at 17 and further preheated to the inlet temperature of the reformer catalyst tubes. The combined steam-gas mixture is essentially equivalent to a normal natural gas-steam mixture in terms of carbon-hydrogen ratio, theoretical hydrogen, and steam-carbon ratio and thus is easily reformable in the radiant section of the reformer furnace which consists of parallel connected reformer tubes manifolded for equal distribution of inlet gas and externally heated by fuel burners 19. The reformer tubes contain a supported nickel catalyst. The preferred arrangement of catalyst consists of an alkali metal promoted nickel reforming catalyst containing 0.1 to 10 wt. percent alkali as the metal in the upper zone of the reformer tubes and a high activity nickel catalyst in the lower zone. The ratio of alkali metal promoted catalyst to high activity catalyst varies with the requirements of the process. As a rule, the volumetric ratio of alkali metal promoted catalyst to the total catalyst volume is from 0.2 to about 0.8 to 1.0, the actual ratio employed depending on the inlet temperature of the reformer, the nature of the feed, and the temperature profile throughout the reforming tube. Examples of commercially available catalysts particularly suited for the second stage reformer include:

Upper zone: Imperial Chemical Industries, ICI 46-1 Lower zone: Imperial Chemical Industries, ICI 57-1 The gas leaving the reformer furnace flows into an outlet header 20 at essentially reformer catalyst outlet temperature and is delivered to an iron ore reduction facility (not shown). The facility may be a direct reduction unit or a blast furnace. The content of reductants in the reformed gas is of the order of 88 percent or better, preferably 88 to 99 percent C0 H Typical conditions and compositions for the process represented by the FIG. 1 flowsheet are given in Table l for a 365F end point naphtha feed. In this case, the incoming naphtha feed emanated from a desulfurization unit (not shown) in which a recycle gas containing hydrogen, was used for removal of sulfur from naphtha by catalytic hydro-desulfurization.

TABLE 1 COMPOSITION OF PROCESS STREAMS (vow/t PRO- CESS FEED TO CRG EFFLUENT FROM FEED TO co OVERHEAD FROM FEED TO EFFLUENT FROM REFORMER LOCA- REACTOR CRG REACTOR ABSORBER co ABSORBER FURNACE REFORMER FURN. 'l'lON cow- NI-ZN'I wET DRY wE'r DRY wE'r DRY WET DRY WET DRY WET DRY NAPH- 7.62 75.00

THA

TABLE ilcisfltifihed COMPOSITION OF PROCESS STREAMS (VOL/7r) PRO- CESS FEED TO CRG EFFLUENT FROM FEED TO CO OVERHEAD FROM FEED TO EFFLUENT FROM REFORMER LOCA- REACTOR CRG REACTOR ABSORBER CO ABSORBER FURNACE REFORMER FURN. TION COM P WET DRY WET DRY WET DRY WET DRY WET DRY WET DRY NENT C11 0.36 3.51 32.37 58.85 57.56 58.85 74.15 75.99 40.94 75.99 i 0.32 0.33 H 2.17 21.29 9.45 17.17 16.79 17.17 21.63 22.18 11.94 22.18 74.03 76.14 CO 0.01 0.10 0.70 1.26 1.24 1.26 1.59 1.63 0.88 1.63 22.41 23.05 C 0.01 0.10 12.50 22.72 22.22 22.72 0.20 0.20 0.11 0.20 0.47 0.48 H O 89.83 44.98 2.19 2.43 46.13 2.77

DH 1( )0.00 100.00 100.00 100.00 100.00 10000 100.00 100.00 100.00 100.00 100.00 100.00 TOTAL TEMP. "F 840 975 990 150 170 150 170 950 1000 1850 PRESS psig. 475 500 450 475 200 300 200 -300 75 100 FIG. 2 shows a modified version of the FIG. 1 process flowsheet in which recycle hydrogen from the CRG reactor outlet is delivered to a naphtha hydro-desulfurization system comprising a preheater and a catalytic facility in which a combination of cobalt-molybdate and zinc oxide catalysts or nickel-molybdate and zinc oxide catalysts is provided. These are considered conventional catalysts for removal of sulfur compounds from naphtha and other similar hydrocarbon feeds.

In using the apparatus of FIG. 2, the naphtha feed has a sulfur compound content at an intermediate level (say 50 ppm) and the sulfur is removed by treatment with nickel-molybdate and zinc oxide catalysts or similar systems. For much higher sulfur contents, it is necessary to provide additional desulfurization facilities in the design.

Referring to FIG. 2, liquid naphtha is delivered through conduit 100 and joined at inlet 101 with a recycle gas stream in an amount equivalent to the requirements of the desulfurization facility. The recycle gas stream supplies the hydrogen required to convert organic sulfur to H 8 in the downstream desulfurization step. The combined naphtha-recycle gas mixture passes through conduit 102 to a convection bank coil shown generally by reference numeral 103, after which the preheated mixture flows through conduit 104 to a heat exchanger 105 for further preheating by recovery of heat from the CRG reactor effluent circuit. The mixture flows through conduit 106 to desulfurizer heater 107 for superheating to the required inlet conditions of a desulfurizer 109. The desulfurizer 109 is provided with a bed of nickel-molybdate catalyst which serves to catalyze the destructive hydrogenation of organic sulfur compounds to H 8, which in turn is absorbed by contact with a layer of zinc oxide absorbant provided in the lower portion of the desulfurizer. I

The desulfurized naphtha-recycle gas mixture leaves the desulfurizer through outlet 110 and is joined at inlet 111 with steam produced within the process. The combined mixture flows through conduit 112 to a heater 113 where it is heated to the inlet conditions of a fixed bed CRG gasification reactor 115. In the CRG reactor 1 15, the naphtha is converted to a mixture of hydrogen, methane, carbon dioxide and small amounts of carbon monoxide. The CRG effluent containing the converted gases and unreacted steam, flows through conduit 116 to the feed preheat exchanger 105, after which it is fed to heat recovery equipment 117 which may include a number of items such as a steam generator, a boiler feedwater heater and/or a reboiler for supplying heat to a C0 stripper (not shown) provided in the CO removal system. Other items may be included in the CRG effluent circuit for removing heat. The type of heat recovery system has no bearing, however, on the basic design of this invention.

Following disengagement of water from converted gases in a separator 118, the converted gases flow to a C0 absorber 119 and are contacted countercurrently with a lean regenerative solvent introduced through inlet 120 for removal of the bulk of the CO contained in the stream. CO rich solvent at the bottom of the CO absorber flows through outlet 121 to a solvent regeneration system (not shown) and is subsequently returned, after CO stripping, to the CO absorber.

The overhead from the CO absorber, consisting mainly of methane and hydrogen plus small amounts of carbon monoxide, then flows through conduit 123 to tubular reformer furnace 122 and is processed in the same manner as that described with reference to FIG. 1. Before delivery to the reformer furnace, a slipstream is taken through conduit 101 from the absorber overheat, which serves to supply the recycle hydrogen required for the desulfurization step. The slipstream gas is compressed in a compressor 124 prior to delivery to the desulfurization system.

Other means may be provided for supplying recycle hydrogen for the desulfurization step such as a small separate reformer furnace which may be located in a slipstream circuit of the CRG reactor and serves to convert the methane in the CRG effluent to hydrogen which is then returned to the desulfurization facility. Similarly, other systems can be employed for desulfurization such as a cobalt molybdate catalyst in conjunction with zinc oxide treatment. For large amounts of sulfur in the feed, the desulfurization design may be modified to include use of an H 8 stripper which serves to remove H 8 by fractionation rather than by zinc oxide absorption. For this design, only nickelmolybdate or cobalt-molybdate catalyst is provided. The catalyst serves to destinctively hydrogenate organic sulfur to H 8, which in turn is removed in the H 8 stripper.

FIG. 3 shows a further modification of the FIG. 1 and 2 flowsheets wherein recycle hydrogen is obtained from the outlet of the reformer furnace. The flowsheet also shows a C0 slipstream taken from the overhead of the CO stripper which can be recycled for adjusting the CO/CO ratio in the recycle hydrogen gas in the interest of improving sulfur removal or for suppressing carbon deposition of sulfur treatment catalysts'due to CO disproportionation.

Referring to FIG. 3, liquid naphtha is delivered through conduit 200 and heated in a reformer furnace coil shown generally by reference numeral 201. The preheated stream is joined with a recycle stream flowing through conduit 203 in an amount equivalent to the hydrogen required for naphtha desulfurization. The combined mixture then flows through conduit 204 to heat exchanger 205 for further preheating to desulfurizer heater inlet conditions. The mixture is then superheated in a desulfurizer heater 207 to the required inlet conditions of the desulfurizer 209 which contains nickel-molybdate and zinc oxide catalysts. The desulfurized stream in outlet 210 is joined with steam introduced through conduit 211, the steam being produced within the process. The mixture flows through conduit 212 to a heater 213 in which it is heated to the inlet conditions of a CRG reactor 215 for conversion to hydrogen, methane, carbon dioxide and carbon monoxide. The CRG effluent containing converted gas and unreacted steam, flows through outlet 216 to the feed preheat exchanger 205 after which it is delivered to heat recovery equipment 217.

Following vapor-liquid separation in a separator 218, the converted gas is treated for CO removal in an absorber 219 contact with aregenerative solvent supplied from a C stripping facility 220. The CO stripping facility incorporates a C0 stripper and reboiling and overhead condensing equipment.

The overhead from the CO absorber, consisting mainly of methane and hydrogen plus small amounts of carbon monoxide, then flows through conduit 221 to a reformer furnace 230 for conversion to high strength reducing gas. A slipstream is taken through conduit 222 from the reducing gas product circuit and recycled to the desulfurization system along with a small amount of CO product taken from the CO stripping facility through conduit 223. Although not considered essential, the slipstream CO product may be used for adjustment of the CO/CO ratio of the feedstream flowing to desulfurization equipment. The combined slipstreams are compressed in a recycle compressor 225 and ultimately joined with the incoming naphtha feed.

A bypass circuit 226 may be provided for diverting a portion of the CRG reactor outlet gas directly to the reformer furnace in the event that a downward adjustment of the content of reducing gas in the reformer effluent is required.

SUMMARY OF ADVANTAGES OF THE INVENTION 1. The reducing gas produced by the process described herein may be passed directly to any type of direct reduction unit. Direct reduction of iron ore is discussed in the Journal of Metals, December 1958, pp. 804-809, and in Metals Progress, January 1960, pp. 1 l l] l5.

2. A high content reducing gas having concentrations of H, C0 of 88 percent and greater can be achieved with a wide range of naphtha and similar feeds. The concentration of H CO in the reducing gas with a naphtha feed, therefore, is the same as or higher than that obtained from a natural gas feed.

3. The reforming catalyst in the reformer furnace is maintained in a high state of activity due to the presence of hydrogen in the feed entering the catalyst tubes. High hydrocarbon conversion, therefore, is expected to be maintained throughout the like of the reforming catalyst.

4. The danger of carbon deposition in the reformer tubes is minimized by virtue of the relatively high hydrogen content of the reformer inlet gas.

5. The size of the reformer furnace is somewhat smaller than that required for natural gas reformers of equivalent capacity due to the lower reformer duty resulting from the presence of hydrogen in the inlet gas.

Having thus described our invention, we claim:

1. A process for producing a reducing gas from a hydrocarbon feedstock having an average of from three to 15 carbon atoms, which comprises the steps of:

a. gasifying the hydrocarbon by passing a preheated mixture of the hydrocarbon and steam through a bed of a reforming catalyst in a gasification reactor to produce an effluent gas comprising methane, hydrogen, carbon oxides and steam;

b. removing carbon dioxide from said effluent gas;

and then c. reforming the resulting purified gas in the presence of steam and a reforming catalyst in reforming reactor to produce a gas containing at least 88% of hydrogen and carbon monoxide.

2. A process according-to claim 1, wherein the hydrocarbon feedstock is selected from the group consisting of liquid naphtha liquified propane and liquified butane and mixtures thereof.

3. A process according to vclaim 1, wherein the mixture of hydrocarbon and steam is introduced into the gasification reactor at a temperature of a least 660F.

4. A process according to claim 1 wherein gasification is carried out under a pressure of from psig. to 650 psig.

5. A process according to claim 1, wherein carbon dioxide removal is effected by solvent absorption.

6. A process according to claim 5, wherein the solvent is selected from the group consisting of monoethanolamine, potassium promoted carbonate solution and unpromoted potassium carbonate solution.

7. A process according to claim 1, wherein the gas is heated to a temperature of from 600 to 1,200F before reforming in step (c).

8. A process according to claim 1 wherein the gas leaves the reformer in step (c) at a temperature of from 1,650 to 2,100F.

9. A process according to claim 1, wherein the reforming in step (c) is carried out under a pressure of from atmospheric pressure to 250 psig.

10. A process according to claim 1 wherein the reforming catalyst in the reforming reactor comprises an alkali metal promoted nickel catalyst in the inlet zone of the reformer and an unpromoted nickel catalyst in the outlet zone of the reformer.

11. A process according to claim 5 wherein a hydrogen-containing stream from said solvent absorption step is employed in a hydrodesulfurization of said hydrocarbon feedstock prior to the gasification step.

12. A process according to claim 5 wherein a C0 containing stream from the solvent absorption step is recycled with a hydrogen-containing stream and employed in a hydrodesulfurization of said hydrocarbon feedstock prior to the gasification step.

(c) in a tubular reactor in the presence of about 1.1 to about 1.4 Mols of steam per atom of carbon in said purified effluent stream and in the presence of a catalyst comprising nickel on a support material at an inlet temperature in the range of from about 600F to about l,200F;

. recovering from step (d) a high strength reducing gas containing at least percent hydrogen plus carbon monoxide and less than 10 percent water and carbon dioxide;

passing said reducing gas at an exit temperature in the range of from about 1,650F to about 1 ,200F

to a unit for reducing metallic ores.

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Classifications
U.S. Classification48/214.00A, 48/213, 48/127.7, 252/373
International ClassificationC01B3/48, C10G11/20, C10G49/00, C21B13/00, C01B3/38
Cooperative ClassificationC10G49/007, C01B3/38
European ClassificationC01B3/38, C10G49/00H
Legal Events
DateCodeEventDescription
Mar 31, 1988ASAssignment
Owner name: M. W. KELLOGG, THE, THREE GREENWAY PLAZA, HOUSTON,
Free format text: ASSIGNMENT OF ASSIGNORS INTEREST;ASSIGNOR:M.W. KELLOGG COMPANY, THE;REEL/FRAME:004846/0930
Effective date: 19880111
Free format text: ASSIGNMENT OF ASSIGNORS INTEREST;ASSIGNOR:M.W. KELLOGG COMPANY, THE;REEL/FRAME:4846/930
Owner name: M. W. KELLOGG, THE,TEXAS