METHOD FOR THE GASIFICATION OF HYDROCARBON FEEDSTOCKS

Cross Reference to Related Applications This application claims priority to U.S. Provisional Patent Application No.

60/746,748 filed May 8, 2006, which is incorporated herein by reference in its entirety as if set forth in full.

Background of the Invention 1. Field of the Invention

The present invention is directed to a method for the gasification of hydrocarbon feedstocks to produce a syngas that is useful for the production of hydrogen, hydrogen- containing materials, electricity or other energy products. The method advantageously produces such high-value products while reducing the formation of carbon dioxide (CO ₂ ) as compared to other gasification methods.

2. Description of Related Art

Gasification is a well-known process that converts hydrocarbon materials such as coal, petroleum coke, biomass or similar feedstocks into a syngas comprising carbon monoxide (CO) and hydrogen (H ₂ ). During gasification, pyrolysis of the hydrocarbon material releases fuel-bound H ₂ , oxygen, nitrogen and sulfur leaving residual solid carbon char. Some of the carbon char is gasified to CO and H ₂ with steam (H ₂ O), and some of the carbon is gasified to CO with oxygen.

Table 1 is an illustrative list of major hydrogen-containing materials, naturally- occurring materials and man-made materials, arranged substantially according to their H ₂ content.

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Table 1

One goal of the energy industry is to use methods such as gasification to convert relatively low-value hydrocarbon materials to clean, high-value liquid or gaseous hydrocarbons that can be effectively utilized. Each of the high-value products listed in Table 1 can be synthesized from syngas (H ₂ and CO), except ammonia where N ₂ replaces the CO. However, for synthesis to occur the ratio of H ₂ to CO in the syngas (H ₂ :CO) must approximate the volumetric ratio of the H ₂ and CO in the balanced chemical equation. For example, to make methane (CH ₄ ), the volumetric ratio of H ₂ :CO must be about three, because the synthesis equation requires that three moles of H ₂ be available for each mole of CO:

3H ₂ + CO -» CH ₄ + H ₂ O (1 )

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In conventional gasification, some of the carbon is gasified to CO and H ₂ with steam (H ₂ O) in a highly endothermic reaction, and some carbon is gasified to CO with O ₂ in a balancing exothermic reaction:

C + H ₂ O -» H ₂ + CO (2) C + V ₂ O ₂ -> CO (3)

Conventional gasification produces some H ₂ from the energy that is released as C is oxidized to CO. However, the H ₂ produced is only 50 vol.% of the total gas released from the splitting of water (Equation 2) and that H ₂ is diluted to less than 25 vol.% when considering the additional CO that is produced by the exothermic oxidation of C with O ₂ that is required for balancing the heat (Equation 3). As a result, the syngas derived from solid hydrocarbons by conventional gasification has a H ₂ :CO ratio that is virtually always less than one, which is too low to be used for manufacturing the high- value hydrogen-containing materials that are needed for commerce.

In order to increase the amount of H ₂ derived from conventional gasification, some of the CO must be used to reduce H ₂ O and form CO ₂ as a by-product via the water gas shift reaction, which is conducted in a separate reactor after gasification.

CO + H ₂ O -» H ₂ + CO ₂ (4)

Thus, when conventional gasification is utilized to produce a high H ₂ syngas or a substantially pure H ₂ gas stream, three-fourths of the H ₂ derives from the conversion of CO to H ₂ by the water gas shift reaction. This step produces large amounts of CO ₂ , which is commingled with the H ₂ that is subsequently separated, such as with a pressure swing adsorption (PSA) unit. The net driving force for producing H ₂ and associated co-products from solid hydrocarbons is the oxidation of carbon contained within the hydrocarbon to CO ₂ . This oxidation of carbon is stepwise: after pyrolysis isolates H ₂ (and other gases such as O ₂ , N ₂ and sulfur) from the carbon, the carbon first is partially oxidized to CO and the CO is then oxidized to CO ₂ . About one-third of the heat produced by the complete oxidation of carbon to CO ₂ is released in the first oxidation step (C to CO), meaning that the oxidation of the CO to CO ₂ via the water gas

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shift produces about two-thirds of the total heat released in completely oxidizing carbon to CO ₂ .

Accordingly, while H ₂ production by coal gasification is an established commercial technology, it is only economically competitive with steam-methane reformation (SMR) for the production of H ₂ when natural gas is prohibitively expensive.

Most gasification of hydrocarbon materials such as coal is carried out in moving-bed gasifiers, fluidized bed gasifiers or entrained flow gasifiers. Among other factors, such coal gasification plants have a high capital cost and the gasification reactors generally have an availability of less than about 75 percent causing disruptions in the manufacture of syngas. This low availability is too disruptive to most follow-on processing, for example, the refining of crude oil or manufacture of ammonia.

Several methods for the gasification of hydrocarbon materials have been suggested that utilize a molten metal to facilitate the reaction. For example, Sumitomo Metal Industries has disclosed a method and apparatus for gasifying hydrocarbon materials utilizing a molten metal reactor. An example of this technology is disclosed in U.S. Patent No. 4,738,688 by Nakajima et al. As is disclosed in this patent, hydrocarbon material is gasified by blowing the hydrocarbon material onto the top surface of a molten metal bath with a gasifying agent such as oxygen. In this top- blowing process, the hydrocarbon material is decomposed at high temperature points that form above the molten metal. It is disclosed that the resulting gas is rich in CO and H ₂ and in the proportion of CO ₂ is rather small. The Sumitomo Metal Industries Technology is also disclosed, for example, in U.S. Patent No. 4,389,246 by Okamura et al., which discloses that a stirring gas can be injected into the bottom of the molten metal reactor to increase the efficiency of the process. Another method using a molten metal reactor was developed by Molten Metal

Technology, as is illustrated in U.S. Patent No. 5,395,405 by Nagel et al. In this method, organic waste is gasified by injecting the waste through the top, bottom or sides of the reactor. Gas can also be injected through the bottom of the reactor to create a fountain of molten metal droplets above the surface of the molten metal. U.S. Patent No. 5,358,697 by Nagel discloses that the molten metal can include two molten

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metal phases, where the second molten metal phase is immiscible in the first molten metal phase. The use of two metal phases enhances the oxidation of atomic carbon and forms CO ₂ , which is discharged to the atmosphere after scrubbing. U.S. Patent No. 5,537,940 by Nagel et al. discloses a sequential process wherein organics are injected into a molten metal, such that H ₂ is formed and removed while carbon dissolves into the molten metal. Thereafter, oxygen is injected into the metal to oxidize the carbon and remove carbon oxides. It is disclosed that the formation of CO is favored when the metal is iron.

Another technology using a molten metal reactor, referred to as the HyMeIt Technology, has been disclosed by Maione et al. For example, U.S. Patent No. 6,110,239 by Maione et al. discloses a process in which a high purity, high pressure H ₂ - rich gas stream and a high purity, high pressure CO-rich gas stream are produced separately and continuously using a molten metal gasifier containing at least 2 zones, to avoid the need to separate the gases in downstream equipment. The method can include introducing a hydrocarbon feed into a molten metal bath beneath the molten metal surface in a feed zone operating at a pressure above 5 atmospheres and decomposing the hydrocarbon feed into H ₂ , which leaves the feed zone as a H ₂ -rich gas, and into carbon, which dissolves in the molten metal. The carbon concentration in the metal is controlled to be at or below the limit of solubility of carbon in the molten metal. A portion of the molten metal is transferred from the feed zone to another molten metal oxidation zone operating at a pressure above 5 atmospheres into which an O ₂ - containing material is introduced beneath the molten metal surface to react with a portion of the carbon to form a CO-rich gas. In this manner, the carbon concentration in the molten metal is controlled so it does not reach the concentration at which the equilibrium oxygen concentration would exceed its solubility limit in the molten metal.

Other methods of carbon gasification using a molten metal bath include that disclosed U.S. Patent No. 4,496,369 by Torneman, which discloses a method and apparatus for the gasification of carbon by the injection of carbon, O ₂ gas and iron oxides beneath the surface of a molten iron metal bath.

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Steam reduction is another known method for the manufacture of H ₂ gas. The steam reduction method utilizes the oxidation of a metal (Me) to strip oxygen from steam, thereby forming hydrogen gas. This reaction is illustrated by Equation 5. xMe + yH ₂ O → Me _x Oy + yH ₂ (5) To complete the cycle in a two-step steam reduction process, the metal oxide must be reduced back to the metal using a reductant such as carbon or CO. For example, CO has an oxygen affinity that is similar to the oxygen affinity of H ₂ and they are equal at about 812 ⁰ C. At temperatures above about 812°C, CO has a greater affinity for oxygen than does H ₂ , and the CO or carbon will reduce the oxide of Equation 1 back to the metal as indicated by Equations 6 and 7.

Me _x O _y + yCO → xMe + yCO ₂ (6)

Me _x Oy + yC → xMe + yCO (7)

Generally stated, the function of the metal/metal oxide couple is to transfer oxygen from the steam to the reducing gas (CO) without allowing the H ₂ O/H ₂ of the hydrogen production step to contact the CO/CO ₂ of the metal oxide reduction step. Neither the metal nor the metal oxide is consumed by the overall process.

Oxygen partial pressure (pθ ₂ ) relates to the facility with which the metal may be oxidized (e.g., by steam) and the oxide may be reduced (e.g., by CO). A related mathematical expression is pH ₂ O/pH ₂ , which is proportional to the oxygen partial pressure. Also, an equivalent and inversely related quantity is the hydrogen fraction, expressed as:

— £*— ₍₈₎

(PH ₂ + pH ₂ O)

Certain metals react strongly with water, releasing hydrogen. The oxygen partial pressure in equilibrium with these metals and their oxides together is extremely low. Once the oxides are formed, they cannot be effectively reduced back to the metal.

Conversely, there is another group of metals that produce insignificant quantities of hydrogen when reacted with water. The oxygen partial pressure in equilibrium with

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these metals and their oxides together is quite high. The oxides, therefore, can be easily reduced by CO or carbon.

Between the two foregoing groups of metals are other metals characterized by an oxygen affinity that is roughly the same as the oxygen affinity of H ₂ . Included in this intermediate group of metals are, for example: germanium (Ge), iron (Fe), zinc (Zn), tungsten (W), molybdenum (Mo), indium (In), tin (Sn), cobalt (Co) and antimony (Sb). These are elements that readily produce H ₂ from H ₂ O wherein the resulting oxide can be reduced by carbon and/or CO. That is, these metals have an oxygen affinity such that their equilibrium pH ₂ O/pH ₂ is low enough to be practical for the production of hydrogen, yet the metal oxide is readily reduced by carbon at normal pyrometallurgical temperatures (e.g., about 1200 ⁰ C). These metals are referred to herein as reactive metals, meaning that the metal can be oxidized by steam and the metal oxide can be effectively reduced by carbon or CO.

Iron is a useful reactive metal, and the steam reduction/iron oxidation process was the primary industrial method for manufacturing hydrogen during the 19th and early 20th centuries. At elevated temperatures, iron strips oxygen from water, leaving pure hydrogen.

Fe + H ₂ O → FeO + H ₂ (9)

Excess water is required to maximize H ₂ production from a given amount of iron. After the H ₂ is produced, excess water is condensed leaving an uncontaminated hydrogen gas steam.

An example of this method is disclosed in U.S. Patent No. 6,663,681 by Kindig et al. In this method, steam is contacted with a molten metal mixture including a first reactive metal such as iron dissolved in a diluent metal such as tin. The reactive metal is oxidized to metal oxide, forming a hydrogen gas; thereafter, the metal oxide can be reduced back to the metal for further production of hydrogen without substantial movement of the metal or metal oxide to a second reactor.

An extension of this work is reported in U.S. Patent No. 6,685,754 by Kindig et al. This patent discloses a method for the production of a hydrogen-containing gas

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composition, such as a synthesis gas including H ₂ and CO. It is disclosed that the molar ratio of H ₂ :CO in the synthesis gas can be well-controlled to yield a ratio that is adequate for the synthesis of useful products such as methane or methanol. In this method, a molten metal is provided and steam is contacted with the molten metal to react the first portion of the steam with the metal to form hydrogen gas and a metal oxide. A hydrocarbon material is also contacted with the melt in the presence of the steam to react the hydrocarbon material with a second portion of the steam to form CO. A gas stream is extracted from the reactor, where the gas stream can have a molar H ₂ :CO ratio of at least about 1 :1. After a period of time, the steam contacting can be terminated and the metal oxide can be contacted with a reductant to reduce the metal oxide back to the molten metal.

Summary of the Invention

The present invention is directed to a highly stable, highly efficient process for the gasification of a wide range of hydrocarbon feedstocks. As used herein, a hydrocarbon feedstock is any material that comprises carbon and hydrogen, even where the hydrogen is present in relatively low amounts, such as in pet coke. The process can advantageously minimize the production of carbon dioxide (CO ₂ ) for the production of a given chemical product and net energy export. It is an advantage of the present invention that the solid or liquid hydrocarbons used as the feedstock can be low-value, contaminated hydrocarbons. The method can also have a lower capital cost than conventional gasification.

For the production of H ₂ , both conventional gasification (discussed above) and the gasification method of the present invention utilize the affinity of carbon to capture O ₂ as the driving force to dislodge O ₂ from water and form H ₂ . However, the gasification method of the present invention utilizes a unique chemical pathway as compared to conventional gasification. As is noted above, conventional gasification produces the majority of H ₂ by stripping the oxygen from H ₂ O with the conversion of CO to CO ₂ via the water-gas shift reaction (Equation 4). In contrast, the method of the present invention produces a majority of H ₂ by stripping the oxygen from H ₂ O with an

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intermediary, metallic iron, derived by reducing FeO _x with the energy released during the transition of C to CO, and produces only a minor portion of the overall H ₂ by stripping oxygen from H ₂ O with the conversion of CO to CO ₂ .

Upon heating a given quantity of a solid hydrocarbon feedstock, the fuel-bound H ₂ is released and can be recovered by both conventional gasification and the gasification method of the present invention. The remaining solid carbon is gasified to

CO (Equation 2), but in the present invention, this occurs by a different endothermic reaction:

C + FeO ^■ » Fe + CO (10) Both gasification processes can burn additional carbon in O ₂ to produce the balancing heat that is required for the reaction (Equation 3).

The gasification method of the present invention, however, must regenerate the FeO in order to continue gasifying the incoming carbon. According to the present invention, regeneration is achieved by oxidizing the iron with steam, such as is illustrated by Equation 9.

Thus, virtually all of the H ₂ that is produced according to the present invention (excluding fuel bound hydrogen) derives from this source, namely the oxidation of C to CO, and the production of CO ₂ is therefore reduced as compared to conventional gasification. According to the gasification method of the present invention, the endothermic reaction that gasifies the solid carbon with a metal oxide (Equation 10) requires heat. This heat is derived from the oxidation of additional carbon, but primarily from the heat released by the oxidation of the carbon to CO. Some CO ₂ is generated, but its concentration is significantly reduced as compared to conventional gasification because the reactor off gas during this step is maintained in a highly reducing state such that substantial quantities of CO do not combust to CO ₂ within the reactor.

To regenerate the FeO and allow for continued gasification of the carbon, the iron must be oxidized. According to the present invention, steam can advantageously

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be used to oxidize the iron and simultaneously produce H ₂ (Equation 9), in a separate gas stream.

Therefore, the method of the present invention can: (1) generate H ₂ from the transition of C to CO; and (2) generate a large amount of excess CO that is separate from the H ₂ . The excess CO can be used to generate electricity, essentially harnessing the CO to CO ₂ transition to produce power with high efficiency, such as in a combined cycle generator. Optionally, additional H ₂ can be manufactured by shifting the CO with H ₂ O in a water gas shift reaction and subsequently separating the co-produced H ₂ and CO ₂ . As a result, the CO ₂ emitted per unit of H ₂ or unit of energy produced according to the gasification method of the present invention is significantly less than that for conventional gasification, and the small amount of CO ₂ that is produced can be concentrated to greater than 90% for ease of sequestration.

The gasification method of the present invention can provide two separate syngas streams - one with a very high H ₂ :CO ratio that is referred to herein as a high H ₂ syngas stream, and a second with a lower H ₂ :CO ratio than the high H ₂ syngas stream, which is referred to herein as a high CO syngas stream. The high CO syngas stream is comparable to a conventional gasification syngas stream in that the H ₂ :CO ratio is less than one. However, the raw high CO syngas stream from the present invention advantageously has a higher COiCO ₂ ratio than the raw syngas stream produced by conventional gasification for at least the following reason.

For conventional gasification a minimum ratio of CO:CO ₂ must be established to extract H ₂ from H ₂ O. In contrast, gasification according to the present invention involves oxidation of the metal (iron) as represented by Equation 5 and reduction of the just-formed oxide by carbon as illustrated by Equations 6 and 7. Addition of Equations 5 and 6 eliminates the metal and metal oxide and leaves just carbon to reduce the water, the same reaction as for conventional gasification. Therefore, based upon a superficial comparison, the same CO:CO ₂ ratio should be required in both cases.

However, to create the metallic iron, an intermediary on the path to creating H ₂ according to the present invention, FeO must yield its oxygen to carbon. According to

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the present invention, the FeO is in ionic solution with other oxides comprising the slag (e.g., SiO ₂ , CaO, AI ₂ O ₃ and MgO), and because the cations of the other oxides in the slag also exert a binding force on the oxygen, additional energy is required to extricate the oxygen from the mixture of FeO _x and other oxides. The additional energy required to extricate the oxygen from the solution of mixed oxides arises from the partial combustion of additional carbon, which advantageously increases the ratio of CO:CO ₂ .

Syngas with a high ratio of CO:CO ₂ , the driving force created to remove oxygen from FeO _x , contains more energy than syngas with a low CO:CO ₂ ratio, the driving force created to remove oxygen from H ₂ O. Therefore, more useful work can be obtained from a given amount of high CO:CO ₂ syngas per unit of CO ₂ produced than from a syngas with a lower CO:CO ₂ ratio, a significant advantage of the present invention.

For synthesizing valuable hydrocarbon gases such as those listed in Table 1 above, a precise blend of the high H ₂ and high CO syngas streams can be calculated such that the blended syngas stream contains the H ₂ :CO ratio required for the high- value hydrocarbon or other H ₂ -containing product being synthesized - for example, 3H ₂ :CO for methane, 3H ₂ :N ₂ for ammonia, and 7H ₂ :3CO for propane. For liquid fuels synthesized by the Fischer-Tropsch (F-T) process or similar methods, the H ₂ :CO ratio must be at least about 1.0, and typically is 1.5 or greater, depending upon the catalyst and process employed. Ratios of 1.0 and greater can be readily achieved according to one embodiment of the present invention by combining the proper proportions of the two syngas streams. When synthesizing gases or liquids, there will virtually always be some residual high CO syngas; this remaining syngas is available to produce electricity, such as by the highly efficient IGCC circuit.

The method of the present invention can also co-produce electricity, nitrogen, sulfur, slag and steam with the two syngas streams. The value of these salable co- products can substantially or completely off-set the cost of H ₂ production.

More specifically, the method of the present invention can include co-creating a high H ₂ syngas stream, a high CO syngas stream, N ₂ , electricity and useful by-products such as sulfur, a pozzolanic slag (a cementitious material) and steam. The high H ₂ syngas stream, after heat recovery and purification, can be comprised either of pure H ₂ ,

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or pure H ₂ with traces of CO (that is, H ₂ content is much greater than the CO content). The high CO syngas stream, after heat recovery and purification, can be comprised of CO and H ₂ in a molecular ratio reflecting the C: H ratio in the hydrocarbon that is being reacted (that is, the CO content is greater than the H ₂ content). After co-creating the high H ₂ syngas stream, the high CO syngas stream and electricity, they can be merged in various ways that utilize substantially all of the energy contained therein for subsequent conversion into H ₂ or H ₂ -containing commodities and/or additional electricity and/or steam. By way of example, the commodities can be selected to include: pure hydrogen; pure hydrogen and electricity and/or steam; ammonia and electricity and/or steam; methane and electricity and/or steam; liquid fuels such as gasoline, diesel and jet fuel, and electricity and/or steam; or solely electricity and/or steam. Steam is useful as a source of process heat required by many industries.

According to one aspect of the present invention, the high H ₂ syngas stream, the high CO syngas stream, N ₂ , electricity and other by-products can be produced by processing a hydrocarbon feedstock, water and air through the following equipment:

(1) two or more molten metal reactors, each producing a hot crude syngas stream;

(2) two or more gas-purifying trains that are designed to recover a purified syngas and heat from the hot crude syngas streams, while rejecting particulate material, water-soluble halogens and sulfurous compounds, and optionally rejecting or recovering H ₂ , to form refined syngas streams;

(3) a steam generator (boiler);

(4) an air separation plant to produce substantially pure O ₂ from air;

(5) equipment for sulfide roasting, such as a fluidized bed; (6) equipment for processing sulfur, such as a Claus plant;

(7) equipment for generating electricity and/or steam;

(8) heat recovery equipment, such as where steam is the medium for heat recovery; and

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(9) gas-compression and other general support equipment.

According to one aspect of the invention, the energy that is available in the syngas streams can be used in conjunction with one or more of the following chemical conversion processes: 1. catalyzed gas-synthesis loops operating at relatively low temperatures and high pressure;

2. the Fischer-Tropsch process, including modifications thereof;

3. electrical generation, such as by gas-turbine combined cycle;

4. the electrolysis of water; and/or 5. the water gas shift reaction for producing additional pure H ₂ , followed by separation of CO ₂ .

These chemical conversion processes can be used to produce, for example:

1. Pure hydrogen, such as by combining co-product hydrogen with hydrogen produced either by the electrolysis of water or by water gas shift of some of the CO otherwise dedicated to electricity production;

2. Pure hydrogen and electricity, where CO is rejected from the high H ₂ syngas stream;

3. Pure hydrogen and electricity, where CO is rejected from both the high H ₂ and high CO syngas streams; 4. Ammonia and electricity, where ammonia can be produced in an ammonia synthesis loop, such as from the high H ₂ syngas stream after purification to pure hydrogen (or conversion of the small amount of CO to methane) and N ₂ ;

5. Methane, electricity and steam, where methane can be synthesized in a methanation loop, such as by combining the high H ₂ syngas stream with a portion of the high CO syngas stream - copius amounts of heat are released by the methanation reaction, and this heat can be converted into

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useful steam, which, among other applications, may be used to dry any moisture associated with the incoming hydrocarbon feedstock;

6. Liquid fuels and electricity, where the liquid fuels can be produced utilizing the Fischer-Tropsch process, such as by combining the high H ₂ syngas stream with a substantial portion of the high CO syngas stream; or

7. Electricity, such as by combining the co-product electricity with electricity generated from burning the purified high H ₂ and high CO syngas streams in a combined cycle gas-fired turbine.

Elemental sulfur and pozzolanic slag can be produced as by-products of the process and are salable commodities of the method, further decreasing the net operating cost of the process.

Thus, according to one embodiment, a method for the production of a commodity from raw material reactants is provided. The method can include the steps of providing reactants to a reactor system, the reactants including at least H ₂ O, air and a hydrocarbon-bearing feedstock, reducing a portion of the H ₂ O by contacting the H ₂ O with a reactive metal to reduce the H ₂ O to H ₂ to recover a first syngas stream comprising the H _2, oxidizing at least a portion of carbon contained in the hydrocarbon- bearing material to carbon oxides by contacting the carbon with a metal oxide disposed in a slag layer to recover a second syngas stream comprising CO, and processing the first syngas stream and the second syngas stream to produce at least one energy commodity from the reactor system selected from the group consisting of H ₂ , electricity, gaseous hydrocarbon fuels , liquid hydrocarbon fuels and ammonia.

According to one aspect of this embodiment of the present invention, the second syngas stream comprises at least about 50 vol.% CO. According to another aspect, the reducing step comprises contacting the H ₂ O with a molten reactive metal to convert the reactive metal to a metal oxide. According to yet another aspect, the reactive metal comprises iron. According to another aspect, the reactive metal or the carbon is further contacted with O ₂ to generate heat.

The method can also comprise the further step of heating the H ₂ O to a temperature of at least about 200 ⁰ C and not greater than about 600°C before the

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reducing step. According to another aspect, the step of heating H ₂ O comprises utilizing heat recovered from at least one of the first syngas stream and the second syngas stream. According to another aspect, the hydrocarbon-bearing material is selected from the group consisting of pet coke, coal, municipal waste, rubber tires and biomass. According to still another aspect, the hydrocarbon material further comprises sulfur- bearing or chlorine-bearing compounds, and the method can further comprise the steps of recovering sulfur-containing compounds from the high CO syngas stream, oxidizing the sulfur compounds to form SO ₂ , contacting the SO ₂ with H ₂ S or H ₂ to reduce the SO ₂ , and extracting elemental sulfur from the contacting step. The method can also comprise the step of removing chlorine-containing compounds from the high CO syngas stream by contacting the high CO syngas stream with water to dissolve the chlorine- containing compounds in water and removing the chlorine-containing compounds by water purification.

According to one aspect, the method includes the steps of blending at least a portion of the first syngas stream and at least a portion of the second syngas stream to form a blended precursor syngas stream and producing a gaseous fuel, or producing a liquid fuel from the blended precursor syngas stream such as by a Fischer-Tropsch type synthesis. The blended precursor syngas stream can have a H ₂ :CO ratio of at least 1:1.

According to another aspect, the second syngas stream has a relatively low CO ₂ content, such as not greater than about 25 vol.% CO ₂ or not greater than about 20 vol.% CO ₂ .

According to another embodiment of the present invention, a method for refining a hydrocarbon feedstock comprising hydrocarbons represented as C _x H _y is provided. The method can include the steps of reducing H ₂ O with a molten metal to produce a high H ₂ syngas stream and a metal oxide compound, gasifying a hydrocarbon feedstock comprising C _x H _y by contacting the feedstock with the metal oxide compound to form a high CO syngas stream that is separate from the high H ₂ syngas stream, combining the high CO syngas stream and the high H ₂ syngas stream to form a blended syngas stream, and converting the blended syngas stream to a hydrocarbon-containing product, where x > y in the hydrocarbon feedstock and x < y in the hydrocarbon containing product.

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According to one aspect, the step of reducing H ₂ O comprises contacting H ₂ O with a reactive metal, such as iron, and in particular molten iron. The step of gasifying the solid hydrocarbon feedstock can include contacting at least a portion of the solid hydrocarbon feedstock with oxygen, and the step of gasifying the solid hydrocarbon material can include contacting a portion of the solid hydrocarbon material with molten iron oxide.

According to one aspect, the hydrocarbon feedstock comprises not greater than about 15 mol.% H ₂ , such as not greater than about 10 mol.% H ₂ . The hydrocarbon feedstock can be selected, for example, from the group consisting of tanker sludge, refinery bottoms, municipal waste, rubber tires, biomass, petroleum coke, animal waste and coal. According to one aspect, the hydrocarbon feedstock comprises coal having a sulfur content of at least about 2 wt.%. According to another aspect, the high CO syngas stream comprises at least about 50 vol.% CO.

According to another embodiment, a method is provided for the production of electricity, solid elemental sulfur, pozzolanic slag and at least one hydrogen-containing commodity from a sulfur-containing hydrocarbon fuel comprising carbon and not greater than about 10 mol.% hydrogen. The method can include the steps of providing reactants to a system of reactors, the reactants including at least air, water and the sulfur-containing hydrocarbon feedstock, reducing at least a portion of the water to form H _2, oxidizing at least a portion of the carbon to CO, oxidizing at least a portion of the carbon to CO ₂ and recovering a high H ₂ syngas stream and a high CO syngas stream from the system of reactors. The high H ₂ syngas stream and the high CO syngas stream can be selectively combined to form a precursor syngas stream, which can be reacted to form a H ₂ -containing commodity comprising greater than 10 mol.% H ₂ . According to one aspect, the method substantially precludes the emission of noxious compounds, such as those selected from the group of sulfurous compounds, nitrogen oxides, dioxins, furans and particulates. The carbon dioxide effluent from the method can comprise at least about 90 mol.% CO ₂ for ease of sequestration. Further, the hydrogen commodity can be substantially free of sulfur and sulfur-containing compounds, and can be selected from the group consisting of hydrogen, ammonia, methane, ethane, propane, gasoline, diesel and jet fuel. The hydrogen commodity can

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also comprise substantially pure hydrogen, or can comprise a material selected from the group consisting of ammonia, urea or other nitrogen-containing compounds. The hydrogen commodity can also comprise a material selected from the group consisting of methane, ethane, propane, butane or other gaseous hydrocarbons. Yet another embodiment of the present invention includes a method of producing substantially pure hydrogen from a solid hydrocarbon-containing feedstock by utilizing both the energy derived from the transition of carbon to CO and the energy derived from the transition of CO to CO ₂ in which at least about two-thirds of the hydrogen is produced by utilizing the energy derived from the transition of carbon to CO, leaving the energy available from the transition of CO to CO ₂ to produce electricity or additional hydrogen. According to one aspect, at least a portion of the energy derived from the transition of CO to CO ₂ is captured in a turbine to produce electricity.

Brief Description of the Drawings Fig. 1 schematically illustrates the reactants and the resultant products that can be produced according to an embodiment of the present invention.

Fig. 2 illustrates a binary phase diagram for a tin-iron metal mixture that is useful in accordance with an embodiment of the present invention.

Fig. 3 illustrates a desired temperature operating window that insures slag fluidity for a FeO/CaO/SiO ₂ slag system with a basicity (CaO:SiO ₂ ) of 0.8 that is useful according to the present invention.

Fig. 4 illustrates equilibrium for FeO and carbon in steelmaking reactors at 161O ⁰ C.

Fig. 5 illustrates a reactor that is useful for metal oxidation or metal oxide reduction reactions according to an embodiment of the present invention.

Fig. 6 illustrates a process flow for the continuous production of hydrogen or hydrogen commodities according to an embodiment of the present invention.

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Fig. 7 illustrates a detailed process flow that enhances the production of electrical power and reduces the production of H ₂ according to an embodiment of the present invention.

Fig. 8 illustrates a detailed process flow that enhances the production of H ₂ and reduces the production of electricity according to an embodiment of the present invention.

Description of the Invention

An overview of the method of the present invention is illustrated in Fig. 1. The method 100 includes providing reactants to a reactor system 108, where the reactants include at least air 102, water 104 and a hydrocarbon-bearing feedstock 106.

According to the present invention, the hydrocarbon feedstock 106 can advantageously include relatively low-value hydrocarbons, including those having less than about 10 mol.% H ₂ and which can also include impurities such as sulfur. More specifically, the hydrocarbon feedstock according to the present invention can include low-cost, high heating-value carbon sources such as petroleum coke, scrap tires and liquid petroleum residues; medium cost, high heating-value and low-ash, high rank coals that may contain high levels of sulfur; or low-cost, low-heating value materials such as low-rank (sub bituminous) coal, biomass and the organic portion of municipal waste products. Plastics contained in municipal waste can also be a useful feedstock. Particularly preferred are petroleum coke, other petroleum residues, scrap tires and high-rank low-ash coal. Petroleum coke (also referred to as pet coke) is a black solid that is obtained mainly by cracking and carbonizing residues from the distillation of petroleum oils, especially heavier petroleum oils (tars). It is an advantage of the present invention that the hydrocarbon feedstock can be a low-value hydrocarbon material such as high-sulfur pet coke or low-ash high-rank coal, as the sulfur can be readily controlled by the method of the present invention. Petroleum coke can advantageously have lower ash content than coal and therefore produce less slag.

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It is an advantage of the present invention that low pressure steam can be made available from the process and is suitable for drying incoming hydrocarbon feed 106 to remove moisture, if the feedstock is high in moisture content. A hydrocarbon feedstock having high moisture content can advantageously lead to the formation of additional H ₂ due to the reduction of the water in the reactor. For example, Powder River Basin (PRB) coal has a relatively high moisture content, such as from about 25 wt.% to about 30 wt.%. In some cases, however, the moisture content of the feed may be so high that the metal in the metal oxide reduction reactor is oxidized as fast (or faster) than the iron can be produced by the reduction reaction. Steam, otherwise used to generate electricity, can be utilized to dry the incoming feed to preclude or otherwise reduce the inadvertent oxidation of the metallic iron. If high-moisture hydrocarbon feedstocks are not dried, the gasification can occur in a single reactor, as is disclosed in co-pending and commonly-owned U.S. Patent application entitled, METHOD FOR THE GASIFICATION OF MOISTURE-CONTAINING HYDROCARBON FEEDSTOCKS filed on May 8, 2007. The air 102 provides a source of O ₂ and can optionally provide a source of N ₂ , which may be needed for the production of nitrogen-containing end- products such as ammonia. The method 100 includes providing for the withdrawal from the reactor system 108 of products that can include flue gas 116, electricity 112 (or electricity and steam), elemental sulfur 114, a high CO syngas stream 118, a high H ₂ syngas stream 110, waste water 120 and slag 122.

Generally, the reactor system 108 generically represented in Fig. 1 can include at least two, and preferably at least three or four, substantially identical molten metal reactors, each of which contains a molten metal or molten metal alloy, and molten slag. Substantially identical gas trains are in gaseous communication with each reactor and are adapted to receive and process the crude syngas streams produced by the reactors. The gas trains according to the present invention can be adapted to treat the crude syngas streams to recover the sensible heat, to clean and purify the syngas streams of impurities such as sulfurous compounds, particulates and mercury, and to preclude the formation of noxious compounds such as dioxins and furans, and form a refined syngas stream. The refined high CO syngas stream and/or the high H ₂ syngas stream can optionally be burned in a gas turbine with air at a temperature which virtually

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precludes formation of nitrogen oxides (NO _x ). Alternatively, the syngas stream(s) can be burned with oxygen at a low temperature thereby creating a sequesterable flue gas.

The present invention will now be described in greater detail, and with reference to Figs. 2-8. According to the present invention, some hydrogen gas (H ₂ ) is formed by contacting steam (H ₂ O) with a molten metal mixture that includes at least a first reactive metal such that the H ₂ O is reduced to H ₂ and the reactive metal is oxidized. A reactor operating in this mode is referred to herein as a metal oxidation reactor. The reactive metal can be at least partially dissolved in at least one diluent metal. The diluent metal may also be reactive with the steam, but is, by definition, less reactive with steam than the reactive metal. Thus, the oxygen from the steam preferentially reacts with the reactive metal to oxidize the reactive metal to its metal oxide and reduce a portion of the steam to form a high H ₂ syngas stream that will also include excess steam.

After substantial conversion of the reactive metal to a metal oxide by the steam, the steam injection is terminated and, a hydrocarbon feedstock and other reactants are injected into the reactor under conditions of intense mixing, such as by using submerged lances. Under these conditions, the metal oxide is reduced back to the reactive metal, which re-dissolves into the molten metal mixture. A reactor operating in this mode is referred to herein as a metal oxide reduction reactor. By switching the flows of steam and the hydrocarbon feedstock between two or more reactors, two separated product syngas streams (a high H ₂ syngas and a high CO syngas) can be produced substantially continuously.

To initiate the production of a high H ₂ syngas stream, steam is contacted with a molten metal mixture that includes at least a first reactive metal. The reactive metal preferably has an oxygen affinity that is similar to the oxygen affinity of H ₂ and reacts with the steam to form a metal oxide. For example, the reactive metal can be selected from germanium (Ge), iron (Fe), zinc (Zn), tungsten (W), molybdenum (Mo), indium (In), tin (Sn), cobalt (Co) and antimony (Sb). The molten metal mixture can include one or more reactive metals. The reactive metal preferably should: (1) be soluble in the diluent metal(s); (2) have a very low vapor pressure at the oxidation/reduction temperature(s); and (3) produce one or more oxides when reacted with steam that also

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has a very low vapor pressure at the oxidation/reduction temperature(s). A particularly preferred reactive metal according to the present invention is iron and according to one embodiment the reactive metal consists essentially of iron.

The reactive metal is preferably at least partially dissolved within a second metal, or mixture of metals. The metal into which the reactive metal is dissolved is referred to herein as the diluent metal. The diluent metal may also be reactive with steam, in which case it can be selected from the group of reactive metals disclosed hereinabove, provided that the diluent metal is less reactive than the reactive metal. Alternatively, the diluent metal can be selected from the metals wherein the oxygen partial pressure (pθ ₂ ) in equilibrium with the metal and oxides together is relatively high. These include nickel (Ni), copper (Cu), ruthenium (Ru), rhodium (Rh), palladium (Pd), silver (Ag), cadmium (Cd), rhenium (Re), osmium (Os), iridium (Ir), platinum (Pt), gold (Au), mercury, (Hg), lead (Pb), bismuth (Bi), selenium (Se) and tellurium (Te). More than one diluent metal can be utilized in the molten metal mixture. The diluent metal should not be a metal wherein the oxygen partial pressure in equilibrium with the metal and metal oxide together is extremely low.

Preferably, the diluent metal should: (1) combine with the reactive metal to be liquid in the temperature range of 400 ⁰ C to 1400 ⁰ C; (2) have a very low vapor pressure over this temperature range; and (3) have the capacity to hold the reactive metal in solution. According to a preferred embodiment of the present invention, the diluent metal is tin and in one embodiment, the diluent metal consists essentially of tin. However, the molten metal mixture can also include additional diluent metals, particularly copper and nickel.

A particularly preferred molten metal mixture for steam reduction according to the present invention includes iron as the reactive metal and tin as the diluent metal. Iron has a high solubility in molten tin at elevated temperatures and the melting temperature of the mixture is substantially lower than the melting temperature of pure iron (1538°C).

Although tin is also reactive with steam, it is less reactive than iron. For convenience, the following discussion will refer to iron and tin as the reactive and diluent metals respectively, although the present invention is not limited thereto.

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Due to thermodynamics, steam reduction reactions to form H ₂ require an excess of steam above the stoichiometric requirement. The total steam requirement (the mass ratio of steam required to H ₂ produced) for iron is much less than for tin at all temperatures and iron will preferentially oxidize in the molten metal mixture. One significant advantage of utilizing a reactive metal dissolved in a diluent metal is that the residence time of the steam within the reactor is increased with respect to the mass of the reactive metal. That is, a given mass of iron will occupy a first volume as pure iron, but the same mass of iron will be distributed over about twice the volume if the iron is in a 50 weight percent mixture with a diluent metal such as tin. It is preferred that the metal mixture be maintained at a temperature above the solidus line AC of Fig. 2 (e.g., above about 1134°C), and more preferably above the liquidus line (I-II-III-IV) of Fig. 2. A metal-steam reaction temperature that is too high, however, adds significantly to the operating cost. For the completely molten iron/tin system illustrated in Fig. 2, the melt should be maintained at a temperature above the solidus temperature of about 1134°C, more preferably at a temperature of at least about 1200 ⁰ C, and even more preferably at a temperature of at least about 1300 ⁰ C. For the purpose of reasonable economics, the temperature should not be greater than about 1450°C and more preferably is not greater than about 1400 ⁰ C. A particularly preferred temperature range for the completely molten tin/iron mixture is from about 1300 ⁰ C to 1400°C. At 1300 ⁰ C, about 75 weight percent iron dissolves in tin with sufficient superheat and the mixture stays in the molten state as iron is oxidized. Also, the reaction between steam and liquid iron dissolved in tin to form pure hydrogen at 1300 ⁰ C is quite vigorous and the reaction kinetics are excellent. Furthermore, the thermodynamics for the steam/iron system even at 1200°C are relatively good, requiring an excess of only about 12.2 tons of steam to produce each ton of hydrogen (1.37 moles of steam per mole of hydrogen). The preferred operating temperature will also be influenced by slag conditions, as is discussed below.

It is preferred that that the molten metal mixture initially include at least about 3 weight percent iron, more preferably at least about 10 weight percent iron, even more preferably at least about 20 weight percent iron and even more preferably at least about

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40 weight percent iron. Further, the amount of iron in the molten metal mixture should preferably not exceed about 85 weight percent and more preferably should not exceed about 75 weight percent. The balance of the metal mixture in a preferred embodiment consists essentially of tin. Accordingly, the amount of tin in the system is preferably not greater than about 97 weight percent, more preferably is not greater than about 90 weight percent , even more preferably is not greater than about 80 weight percent, and even more preferably is not greater than 60 weight percent. The molten metal mixture preferably includes at least about 15 weight percent tin and more preferably at least about 25 weight percent tin. Due to the relatively high price of tin compared to iron, and within the limits previously given, mixtures low in tin and high in iron are generally preferred.

A method for reacting steam in a tin/iron molten metal mixture to form hydrogen is disclosed in commonly-owned U.S. Patent Nos. 6,663,681 and 6,685,754, both by Kindig et al. Each of these U.S. Patents is incorporated herein by reference in its entirety.

Thus, steam is contacted with the molten metal mixture to generate H ₂ and to oxidize the reactive metal to a metal oxide. The steam is contacted with the molten metal mixture in a manner that promotes good mixing and contact with the molten metal mixture. For example, the steam preferably can be contacted with the molten metal mixture by injection through submerged lances at velocities in the lance approaching the velocity of sound, or through a porous (ceramic) diffuser disposed at the bottom of a reactor. Both methods shear the steam creating small bubbles (large surface area), which facilitates the rate and completeness of the steam-iron reaction. Preferred reactor systems in this regard are discussed below. The metal oxidation reactor temperature can be controlled to maintain a substantially constant temperature by controlling the incoming steam temperature and quantity, and/or by adding O ₂ to the reactor to burn a portion of the iron.

The metal oxidation reactor can be maintained at an elevated pressure if necessary for adequate residence time in the reactor, as is discussed above. For example, it may be desirable to maintain an elevated pressure, such as at least about 3

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atmospheres (44.1 psi). A slightly elevated pressure can be beneficial for minimizing the size of the first H ₂ compression stage, and for commerce, the pressure of the H ₂ will have to be increased by compression to between 200 and 450 psi, depending upon the user. Although, significantly increased pressure in the metal oxidation reactor adds to capital cost, it also reduces or eliminates the H ₂ compression requirement and associated cost. Therefore the pressure in the metal oxidation reactor may be about 50 psi and more preferably about 225 psi and most preferably is 400 psi.

In one preferred embodiment, H ₂ production in the metal oxidation reactor (i.e., high H ₂ syngas production) preferably begins when the concentration of iron in the alloy is about 65 weight percent (about 35 weight percent tin) and continues until the iron concentration drops to about 40 weight percent (about 60 percent tin). Over the same time period, the concentration of the iron oxide in the molten slag above the metal increases from about 30 weight percent FeO (the balance being mixed oxides of silicon, aluminum, calcium, magnesium, etc.) to a preferred maximum concentration of about 65 weight percent FeO, which values are dictated by the temperature of the slag-freeze line. The freeze line for the slag rises steeply at both high and low concentrations of iron oxide, the points where injection of steam is either initiated or terminated. This is illustrated in Fig. 3. The slag composition illustrated in Fig. 3 has a basicity (CaO:SiO ₂ ratio) of 0.8. A slag layer provides a number of advantages, including preventing the iron from exiting the reactor. The temperature in all reactors should be sufficient to maintain the slag layer that forms over the metal mixture in the molten state over a range of compositions, as illustrated in Fig. 3. As the reactive metal is oxidized, a decrease will occur in the concentration of the reactive metal in the metal alloy and the metal alloy should remain molten as the reactive metal is oxidized. Similar to the range of compositions for the molten alloy discussed previously with respect to Fig. 2, there is a range of preferred slag compositions required to ensure adequate slag fluidity and reactivity. Fig. 3 illustrates the preferred operating temperature window (A-B-C-D) superimposed on a graph above the slag freeze line for a FeO _x /CaO/SiO ₂ slag as a function of temperature and FeO _x content. The reactor operating temperature must be above the slag freeze line where the slag is molten. Slag properties can be adjusted,

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for example fluxes such as SiO ₂ , CaO, MgO, Na ₂ O and K ₂ O can be added to the reactor to adjust the properties of the slag. Moreover, sulfur and other anions may be incorporated in the slag to secure satisfactory slag chemistry.

The product gas that emanates from the metal oxidation reactor is a high H ₂ syngas. Inadvertently, a small amount of CO typically is generated in the metal oxidation reactor from the reaction between liquid carbon (dissolved in the melt) and the steam that is injected into the molten metal to make hydrogen. The presence of CO in the high H ₂ syngas may be undesirable, so it is an advantage of the present invention that the amount of carbon that can be dissolved in molten iron tends toward zero as the iron oxide content of the slag increases, particularly above 10 weight percent and more particularly above 30 weight percent FeO in the slag.

For example, Fig. 4 illustrates the composition of dissolved carbon with respect to iron oxide in the slag for the equilibrium condition and the conditions encountered in two steel-making operations (BOF and OBM) at 1610 ⁰ C. Fig. 4 is derived from "Fundamentals of Iron and Steelmaking", the AISE Steel Foundation (1999), and illustrates that at an FeO content of about 35%, the lowest preferred FeO content in the reactor according to one embodiment of the present invention, the equilibrium carbon dissolved in the iron can not exceed about 0.04%. The small amount of carbon dissolved in the iron is the source of CO that arises when the alloy (containing the dissolved carbon) is contacted with steam. The presence of CO in the high H ₂ syngas stream is undesirable for high purity applications such as making ammonia. According to the present invention, a high FeO content in the slag assures that the molten alloy can hold very little dissolved carbon and that the production of CO from dissolved carbon will be minimal. During production of the high H ₂ syngas, the metal oxide that is generated can advantageously be trapped by dissolution in the slag layer within the reactor. At the preferred temperatures, the iron oxide is molten and is incorporated into the slag, which is lighter than the metal mixture. Therefore, as the dissolved iron is depleted from the molten metal mixture, the molten iron oxide rises through the molten metal and contributes to the slag layer on top of the molten metal. It is an advantage that the

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oxide formed upon reaction of the reactive metal with the steam has a density that is less than the density of the molten metal mixture, whereby the oxide rises to the slag layer. Preferably, the metal oxide is at least about 20 percent less dense than the molten metal mixture. This also enables the metal to sink from the slag layer to the molten metal mixture upon reduction of the metal oxide. As is discussed above, this accumulation of iron oxide in the slag may require the addition of a flux such as SiO ₂ , CaO, MgO, Na ₂ O, K ₂ O or mixtures thereof to maintain the slag in the preferred condition with respect to viscosity, reactivity, foaming, and the like.

The molten metal must be contained within a suitable reactor to maintain the desired reaction conditions. Further, the reactants should be provided in a manner conducive to good mixing and high contact surface area. High-temperature reactors suitable for establishing good gas/liquid/solid contact are utilized in the chemical and especially metallurgical industries.

One reactor system that can be useful according to the present invention, referred to as a bath smelter, utilizes lances to inject the steam and other reactants into the molten metal. Examples of reactors utilizing top or side submerged lances to inject reactants are disclosed in U.S. Patent No. 3,905,807 by Floyd, U.S. Patent No.

4,251 ,271 by Floyd, U.S. Patent No. 5,251 ,879 by Floyd, U.S. Patent No. 5,282,881 by

Baldock et al., U.S. Patent No. 5,308,043 by Floyd et al. and U.S. Patent No. 6,066,771 by Floyd et al. Each of these U.S. patents is incorporated herein by reference in its entirety. Such reactors are capable of injecting reactants (e.g., hydrocarbons, steam and oxygen) into the molten metal at extremely high velocities, approaching Mach 1 , thereby promoting good mixing of the reactants.

The major function of the lance entering the bath smelter is to maximize contact between the solid, liquid and gas phases within the reactor. Fig. 5 schematically illustrates a cross-section of such a reactor. The reactor 500 includes sidewalls S02 that are adapted to contain the molten metal 504 and slag 506. The sidewalls 502 can optionally be cooled, such as by water cooling. A refractory lining 503 is provided to insulate the portion of the reactor containing the molten metal 504 and slag layer 506. Top-submerged lances 508 and 510 are disposed through the top of the reactor and are

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adapted to inject reactants such as steam into the metal 504 at a high velocity. Preferably, the lances 508 and 510 terminate and inject the reactants below the surface of the slag layer 506, such as near the interface of the molten metal 504 and the slag layer 506. For the oxidation of the reactive metal to proceed, a heat balance must be achieved around the reactor 500. Little heat is provided by the iron-steam reaction as it is only mildly exothermic. The steam supplied to the reactor could be superheated, for example, to a temperature of about 1200 ⁰ C. However, such high temperature steam flows are difficult to contain with standard processing equipment. A preferred method is to provide some heat by heating the steam to an elevated temperature, such as at least about 400 ⁰ C and more preferably at least about 500 ⁰ C. Temperatures substantially in excess of about 500°C may require special handling equipment due to the difficulty of containing steam at such high temperatures. Therefore, the steam is preferably heated to not greater than about 700 ⁰ C and preferably not greater than about 600°C. Since the heat of the incoming steam is not entirely sufficient to provide the required heat for the reaction to proceed, additional heat must typically be supplied to the reactor. In this regard, oxygen also can be injected into the reactor to burn a portion of the metal (or hydrogen) and provide the additional heat. For example, substantially pure O ₂ gas can be injected with the steam down either or both lances 510 and 508. The oxygen oxidizes the molten reactive metal in an exothermic reaction and creates the heat necessary to raise the temperature of the incoming superheated steam to the reactor operating temperature and to sustain the oxidation of the reactive metal by the steam. Preferably, a substantially pure oxygen-containing gas is provided and it is typically advantageous to minimize the amount of nitrogen (e.g., from air) injected into the reactor. However, it may be necessary to dilute the O ₂ gas with a carrier gas to reduce the possibility of burning the lances.

In the metal oxidation mode described above, a high H ₂ syngas is withdrawn from the top space of the reactor 512. The metal oxidation process is continued until the quantity of iron in the reactor decreases to a sufficiently low level.

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Thereafter, the injection of steam into the reactor is terminated and a hydrocarbon feedstock is introduced into the reactor containing the molten metal and the slag, which now includes an increased amount of iron oxide. This is referred to herein as the metal oxide reduction mode. The carbon in the hydrocarbon feedstock gasifies by reducing the metal oxide, now in the slag, back to the metal. The point at which the metal oxidation process is terminated and the metal oxide reduction process is begun can be determined based in-part on the quantity of iron and iron oxide in the reactors. These quantities can be pre-established on the basis of operating in the 1250 ⁰ C to 1400 ⁰ C temperature range above the freeze line for the slag (Fig. 3). In the metal oxide reduction step, which can be viewed as reductive cleaning of the slag, the reactive metal oxide (e.g., iron oxide) in the slag is reduced by the carbon in the hydrocarbon feedstock and returned to the melt. This is achieved by lowering the oxidation potential of the system by introducing additional hydrocarbon feedstock to the reactor thereby increasing the ratio of hydrocarbon to oxygen; simultaneously, this increases the ratio of CO:CO ₂ . The increased CO:CO ₂ ratio lowers the oxidation potential of the system thereby removing oxygen from the oxide and returning the metal to the melt. The particulate solid or liquid hydrocarbon feedstock is injected into the reactor, preferably at the slag/alloy interface under conditions of intense mixing. According to one embodiment, an iron-containing feedstock such as scrap tires can advantageously supply additional iron to the reactor to make up for incidental losses of the iron.

Preferably, the temperature when operating in metal oxide reduction mode is substantially identical to the conditions during metal oxidation mode. That is, it is preferred that the temperature of the reactor during metal oxide reduction is the same or very similar to the temperature of the reactor during metal oxidation. Thus, the temperature is preferably: (1) at least above the liquidus for the alloy (e.g., about 1134°C for the tin/iron system); and (2) at least above the slag freeze line (Fig. 3). Preferably, the temperature does not exceed about 1400 ⁰ C and more preferably does not exceed about 135O ⁰ C. In a particularly preferred embodiment, the temperature is in the range of from about 1300°C to about 1350 ⁰ C in both reactors.

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In the metal oxide reduction mode, the hydrocarbon feedstock is essentially subjected to gasification within the metal oxide reduction reactor. That is, the hydrocarbon feedstock is quickly pyrolized to release fuel-bound H ₂ and form carbon. The carbon gasifies while also acting as a reductant for the metal oxide. As is noted above, the slag composition can include a number of compounds, including silica (SiO ₂ ), calcia (CaO), alumina (AI ₂ O ₃ ) and magnesia (MgO). It has been advantageously found according to the present invention that when other oxides are contained within the slag, the iron oxide requires a higher reducing potential for producing iron metal. Therefore, an additional amount of hydrocarbon feedstock is necessary above the amount that would be required to reduce iron oxides in the absence of other oxides, and the equilibrium gas composition resulting from the reductive cleaning of the slag advantageously has a higher CO:CO ₂ ratio. According to the present invention, this advantageously decreases the amount of CO ₂ produced per unit of energy manufactured (e.g., CO ₂ /MW-hr of electricity produced or CO ₂ /tonne of hydrogen produced). A CO:CO ₂ ratio of 4, for example, contains 80 percent CO and 20 percent CO ₂ (exclusive of other gases) whereas a CO:CO ₂ ratio of 2 contains 66.7 percent CO and 33.3 percent CO ₂ . Combusting either gas produces the same amount of CO ₂ (100 percent). However, more useful work can be derived from the gas with the higher CO content. Therefore, the higher the CO:CO ₂ ratio, the greater the work that can be accomplished per unit of CO ₂ produced.

The metal oxide reduction reactor also can be maintained at atmospheric pressure or at an elevated pressure. Pressure in the metal oxide reduction reactor can be achieved by employing an air separation unit (ASU) that produces liquid (as opposed to gaseous) oxygen. It can be an advantage to pressurize the metal oxide reduction reactor because it mitigates the subsequent cost of compressing the high CO syngas ahead of the gas turbine. Periodically, however, the metal oxide reduction reactor must be tapped to remove slag, and tapping under pressure (or releasing pressure before tapping) can be difficult, time consuming and costly.

Referring back to Fig. 5, a description of the same reactor operating in metal oxide reduction mode will be given. The reactor includes side-wall(s) 502. In the case

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of the iron-tin system, the molten alloy 504 initially includes predominantly tin, although some iron will still be present. The slag layer 506 includes the iron oxide that was formed during the metal oxidation process described above along with other metal oxide compounds. A hydrocarbon feedstock can be injected through top-submerged lance 508. Oxygen, preferably diluted with a carrier gas (but no nitrogen) also can be injected through the same lance or through a separate lance 510. Under the reducing conditions achieved and maintained in the metal oxide reduction reactor by the addition of the hydrocarbon, CO is formed from the carbon, and some H ₂ is released from the hydrocarbon. The gases can be withdrawn from the reactor through port 512. The iron oxide is reduced back to iron and then re-dissolves within the molten metal 504. A crude (contaminated) syngas containing CO and the H ₂ released from the hydrocarbon feedstock (i.e., fuel-bound hydrogen) can be removed through outlet port 512. In addition, slag 506 can be periodically tapped through slag outlet port 514.

Controlling the oxidation potential of the reactor contents in metal oxide reduction mode comprises one factor that controls the rate of the reduction reaction and therefore the rate at which iron oxide in the slag is reduced to iron and reports to the melt during metal oxide reduction. This rate can be maximized by controlling the relative amounts of oxygen and hydrocarbon feedstock, such as coal, that are injected into the reactor.

In turn, these relative amounts of oxygen and hydrocarbon control the ratio of the partial pressure of the oxidized gases to the partial pressure of all the gases as expressed by the fraction:

(H ₂ O + CO ₂ ) (11)

(H ₂ + H ₂ O + CO + CO ₂ ).

The preferred value for this ratio according to the present invention is established through minimization of Gibbs' free energy for the reduction reaction. That is, the process is most effective and produces the highest COiCO ₂ ratio and the highest quality syngas stream from the metal oxide reduction reactor when the value of the ratio of the oxidizing gases to total gases is determined by minimization of the Gibbs' free energy for the particular hydrocarbon feedstock that is being fed to the reactor. Process control can also be based on approaching (targeting) the pre-calculated preferred ratio of oxidized gases to total gases (Equation 11), which is unique for each different

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hydrocarbon feedstock and which value when approached insures rapid reduction of iron oxide to iron. According to one embodiment, the value of Equation 11 is maintained at about 0.229 to maximize the rate of the reduction reaction.

In accordance with the foregoing, two or more reactors can be operated in parallel for the production of a high H ₂ syngas and a high CO syngas in a continuous manner. As the reactive metal is depleted from the molten metal in a metal oxidation reactor, and as the metal oxide is reduced to metal in a metal oxide reduction reactor, their functions can be reversed by switching the flows into the reactors and out of the gas purification trains. Although the functionality of the reactors is reversed, there is no movement of metal or metal oxide into or out of the reactors, except as may be required for make-up of incidental losses.

The cycle time for the metal oxide reduction reactor will be greater than the cycle time for the metal oxidation reactor. Therefore, it may be desirable to utilize more reactors operating in the metal oxide reduction mode than reactors operating in the metal oxidation mode at any given point in time. According to one embodiment, at least two reactors are utilized operating in metal oxide reduction mode for every reactor that is operating in metal oxidation mode. The metal oxide reduction reactors can be staggered in time to produce a substantially continuous flow of product gases.

A flowsheet illustrating the gasification of a hydrocarbon feedstock using three reactors according to the present invention is illustrated in Fig. 6. The process employs two or more reactors to provide substantially continuous production of two or more separate syngas streams. Reactors 602 and 604a/604b are preferably bath smelters similar to those described above with respect to Fig. 5.

Gas purification trains 680a/680b and 682 are in gaseous communication with reactors 604a/604b and 602, respectively. It is an advantage of the present invention that a major portion of each of the gas purification trains 680a/680b and 682 can be substantially identical to enable: (1) reduction of capital cost and maintenance cost; and (2) the production mode of each reactor to be switched to continually produce both a high H ₂ syngas stream and a high CO syngas stream.

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After the syngas purification trains, valves 670 control the direction of the syngas streams, depending on whether the syngas stream is a high H ₂ syngas stream or a high CO syngas stream, and depending on the desired end products. For example, as is illustrated in Fig. 6, the high H ₂ syngas stream from reactor 602 can be further refined such as by processing through a water gas shift reactor 634 and/or a pressure swing adsorption (PSA) unit 636, after which the high H ₂ syngas stream can be transported or placed in storage 650. Likewise, the gas purification trains 680a/680b direct the refined high CO syngas stream to the valves 670. All or a portion of this gas stream can be taken to an electricity generating turbine 642 to optionally burn CO and fuel-bound H ₂ and generate electricity, or employ the water gas shift reaction 634 and/or the PSA unit 636 to manufacture additional H ₂ .

A processing alternative to burning the high CO syngas stream from the gas purification trains 680a/680b, is to combine the high CO syngas stream with the high H ₂ syngas stream 682 using valves 670 to create a precursor gas for manufacturing hydrogen-containing commodities in unit operation 690, for example Fischer-Tropsch synthesis. It is an advantage of the present invention that some amount of the high CO syngas stream will remain after creating the precursor gas stream, and it can advantageously be used to generate electricity and heat, particularly heat to raise steam for injection into the metal oxidation reactor 602. As illustrated in Fig. 6, three reactors are utilized for the substantially continuous production of reaction products. Since the reaction of steam with a molten metal to produce H ₂ is more rapid than the metal oxide reduction process, at any given time two or more reactors are operating in metal oxide reduction mode for each reactor that is operating in metal oxidation mode.

As illustrated in Fig. 6, reactors 604a/604b are operating in metal oxide reduction mode while reactor 602 is operating in metal oxidation mode. Reactors 604a/604b can be staggered in time such that when one reactor reaches the end of the cycle it is immediately switched to metal oxidation mode, while the other metal oxide reduction reactor continues to operate. At the same time, the reactor functioning as the metal oxidation reactor is switched to metal oxide reduction mode.

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As is described more fully below, the gas trains can include multiple units-of- operation for: (a) rapid gas cooling by water quenching, such as when the temperature of the high CO syngas stream is rapidly decreased from about 1350°C to 700 ⁰ C by adding liquid water to preclude metal "dusting", a type of corrosion that occurs with high concentrations of CO and high temperatures; (b) gas cooling and heat recovery by heat exchangers; (c) removal of solid pollutants such as fine particles, for example by filtration; (d) catalytic conversion of carbonyl sulfide (COS) to hydrogen sulfide (H ₂ S); (e) condensation of water vapor and simultaneous removal of soluble halogen acid gases such as HCI; (f) compressors to compress the syngas streams to meet the requirements of the amine scrubbers and other downstream operations, such as an integrated cycle gas turbine 642 to generate electricity and to provide heat required to raise the steam for production of the high H ₂ syngas; (g) amine scrubbers or solvents to capture hydrogen sulfide (H ₂ S); and (h) other purification unit operations for capturing pollutants originating within the hydrocarbon feedstock, such as activated carbon for capturing volatile species of mercury, a common pollutant emitted by coal-fired electrical generating plants.

After purification, the gas streams can be introduced to PSA units 636 to purify the hydrogen. Additional equipment can include one or more steam turbines to generate electricity utilizing the excess steam that is raised above what is required for producing hydrogen; an air separation plant adapted to isolate O ₂ and N ₂ from air; a pelletizer or briquetter to agglomerate furnace dust and condensed tin sulfide (SnS) and a roaster (e.g., a fluid bed or multiple hearth roaster) for oxidizing the solid sulfur species extracted from the syngas streams to sulfur dioxide (SO ₂ ) and calcine (largely SnO ₂ ); an amine regeneration section for recovering H ₂ S from amine solutions; a cooling tower or other means of condensing spent steam; and a water purification system capable of producing boiler-quality feed water.

The process can include providing the reactants including at least water, a hydrocarbon feedstock (which can be represented by C _y HχO _z N _a S _b Ash _c ) and O ₂ obtained from an air separation plant. The feed can also include a flux to control slag properties and make-up metals to replace incidental metal losses. In a metal oxide reduction reactor, pyrolysis of the solid hydrocarbon releases fuel-bound hydrogen (H _x ) as a gas

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and carbon (C _y ) as a solid; other constituents of the hydrocarbon fuel (O _H , N ₃ , S _b and ash) are also released (i.e., molecular bonds broken) by the pyrolysis. Also in the reactors 604a/604b, at least a portion of the C _y is oxidized to carbon oxides with oxygen derived from molten iron oxide (FeO _x ), an endothermic reaction, and at least a second portion of the carbon is oxidized to carbon oxides with oxygen from an air separation plant, an exothermic reaction that at least partially balances the above endothermic reaction. In one embodiment, the molten metal alloy in the reactor contains tin, and sulfur, if any is contained in the hydrocarbon feedstock, principally reacts with tin to form the volatile species SnS, although some sulfur can react with hydrogen forming a small concentration of H ₂ S. The ash is also fused and migrates into the slag.

In a separate reactor 602 a portion of the steam (H ₂ O) is reduced to H ₂ while a portion of the iron is oxidized to FeO _x . Sufficient O ₂ can also be introduced, burning either iron to FeO _x or H ₂ to H ₂ O, to produce sufficient heat to achieve an energy balance about the reactor 602. Thus, a high CO syngas stream is recovered from the reactors 604a/604b, and a high H ₂ syngas stream is recovered from the reactor 602.

The temperature of the high CO syngas stream can be rapidly reduced by water quenching to preclude corrosion issues that arise from high temperatures and high CO concentrations. Heat can be recovered from both syngas streams utilizing conventional heat exchangers and the recovered heat can be used to raise steam. Fine particulates, such as furnace dust and condensed SnS, can be removed from the high CO syngas stream, and the particulates can be agglomerated and roasted with oxygen. The roasted product (calcine) from the first purification train can be re-injected into the reactors 604a/604b to conserve metal values and the SO ₂ from the roasting operation can be directed to a Claus plant for recovery of sulfur. Fine particulates, such as furnace dust, can also be removed from the high H ₂ syngas stream and after agglomeration and roasting can then be re-injected into the reactors. Carbonyl sulfide (COS) that may be present in both syngas streams can be catalyzed to H ₂ S and the H ₂ S can be removed from both syngas streams by amine scrubbing or other processes. The H ₂ S released from the amine regeneration unit

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operation can be directed to a Claus plant where it reacts with the SO ₂ derived from roasting furnace dust to form elemental sulfur.

2H ₂ S + SO ₂ = 3S + H ₂ O (12)

If insufficient H ₂ S is available to react with the SO ₂ , H ₂ can be provided to the Claus plant, as needed to meet the reduction requirement for making elemental sulfur.

Slag from the reactors 604a/604b can be tapped (removed), preferably in an amount approximating the amount of ash plus flux materials that are added, or in an amount that precludes the rapid build-up in concentration of some compound in the slag, such as vanadium pentoxide, that is frequently present when petroleum coke is utilized.

According to the present invention, burning relatively low-ash carbon feedstocks is preferable to burning high-ash carbon feedstocks. This is because a flux, usually CaO, SiO ₂ or both must be added in proportion to the amount of ash in the feedstock to control slag properties. Thus, for high-ash carbon, there must be a large slag tap so that an equivalent amount of slag is removed as ash and flux are added. At the end of the metal oxide reduction cycle, the slag typically contains about 2% tin and about 30% iron, and this translates into a significant economic loss if not recovered and recycled to the reactor.

However, coal is plentiful and widely distributed and can be used as a hydrocarbon feedstock for producing clean gasoline and diesel fuels in accordance with one embodiment of the present invention. The relatively higher ash content of coal and consequently the high losses of iron and tin with each slag tap, however, is problematic. Aggressive coal cleaning such as froth flotation is a widely practiced approach known for minimizing ash content when using coal. Even with such measures, however, there may be more ash than is economically desirable.

Therefore, one aspect of the present invention anticipates recovering iron and tin from the slag.

Tin recovery can be achieved by mixing elemental sulfur, a by-product of the present method, into the slag just after removing slag from the reactor and while the slag is still hot. Sulfur will react with either elemental tin (Sn) or tin dioxide (SnO ₂ ),

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whichever form is present in the slag, to create volatile tin sulfide (SnS). If SnO ₂ is present, SO ₂ also will be formed. If excess sulfur is added over and above the stoichmetric requirement to volatilize the tin compounds, the excess sulfur will react with the iron oxide to form FeS. The expected chemistry and associated thermodynamic parameters are listed below, and all three reactions are expected to proceed to the right.

(Delta Hi3oo-c = -46.1 1 kcal, Delta G ₁₃₀₀ -C = -41.33 kcal)

SnO ₂ + S -> SnS(g) + SO _2(g ) (14)

(Delta Hi3oo°c = -49.84 kcal, Delta Gi ₃ oo°c = -70.45 kcal)

FeO + S = FeS + SO ₂ (15) (Delta Hi ₃ oo°c = -29.80 kcal, Delta Gi∞crc = -23.32 kcal)

The above three reactions are exothermic, and the heat generated from the vigorous reactions will help off-set thermal losses to the atmosphere and keep the slag from prematurely freezing.

For example, the slag can be directed into a crucible or converter immediately after tapping. Sulfur, irrespective of its state, can be blown into (through) the hot slag with a non-reactive gas using lances. The expected off-gas will contain the volatile species SnS and SO ₂ . Iron sulfide will be formed in the slag if excess sulfur over the stoichiometric requirement is used and the tin species have already been volatized.

The volatile compounds SnS and SO ₂ can then be directed to the dry bottom quench, which can be directly connected to the metal oxide reduction reactor being tapped. Any

SnS entering the quench as a gas will be cooled to form solid SnS, which can be captured in a baghouse and recycled.

Any SO ₂ entering the quench will react with H ₂ , which is released from the hydrocarbon feedstock in the reactor, to form H ₂ S and H ₂ O. The H ₂ S can be captured by a downstream amine scrubber and ultimately processed back to elemental sulfur in a Claus plant. The water can be removed by chilling.

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As noted above, any sulfur introduced into the slag which does not react with a tin species will react with the iron oxide content of the slag to form FeS and SO ₂ . The sulfur will appear as a Claus plant product. Any H ₂ that binds with excess sulfur constitutes a minor loss of H ₂ . A second step of this process relates to iron recovery. Addition of the hydrocarbon and oxygen will reduce the FeO _x content of the slag to iron. Further, to the extent that FeS is present in the slag due to tin recovery, it too will be reduced to metallic iron. The relevant equations and thermodynamic properties are given below: FeO + C = Fe + CO(g) (16) (Delta H ₁₃₀₀ -C = 35.23 kcal, Delta dsoo-c = -21.19 kcal)

FeS + C +1.5 O ₂ (g) = Fe + CO(g) + SO ₂ (g) (17)

(Delta H _130OX = -84.83 kcal, Delta Gi ₃₀₀ - _C = -102.97 kcal)

The endothermic nature of Equation 16 may require providing heat to the converter by either burning natural gas or introducing oxygen, and possibly additional (hydro)carbon into the converter.

After the "sulfur blow" and iron reduction, the lances can be removed (retracted) and the slag transferred and quenched in a wet-bottom quench. This quickly freezes the slag and controls particle size, minimizing crushing requirements. Referring back to Fig. 6, the high CO syngas stream and the high H ₂ syngas stream can be mixed together to produce a new blended precursor syngas stream with the desired ratio of H ₂ :CO that is necessary to produce a gaseous or liquid hydrogen- containing fuel in a hydrocarbon synthesis step 690. Some amount of high CO syngas will remain after the new blended precursor syngas stream is produced; this residual high CO syngas can be used for generating electricity via the IGCC route. Alternatively, the high H ₂ syngas stream (or both syngas streams) can be directed to a purification step to produce pure H ₂ , such as by water gas shift 634 followed by pressure swing adsorption (PSA) 636. As another alternative, N ₂ from the air separation plant can be blended with purified H ₂ to make a precursor gas for creating ammonia (NH ₃ ). At least one H ₂ -containing commodity can be recovered from the blended precursor syngas stream, or ammonia gas can be recovered from available N ₂ and

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purified H ₂ , and the H ₂ -containing commodity can be selected from the group consisting of gaseous hydrocarbon fuels, liquid hydrocarbon fuels and ammonia.

Further, at least a portion of the high CO syngas stream, a portion the high CO syngas stream and the reject stream from PSA of the high H ₂ syngas, or the H ₂ reject streams from both PSA systems can be burned in an integrated gas turbine combined cycle (IGCC) with either air or oxygen to provide efficiently generated electricity and thermal energy to raise steam for the process.

The components of the gas purification trains will be understood with reference to Fig. 7 and Fig. 8. A flowsheet illustrating the co-generation of H ₂ and electricity according to the present invention is illustrated in Fig. 7, which is configured to produce large amounts of electricity.

The process typically employs two or more molten metal reactors. As illustrated in Fig. 7, reactor 602 is operating in metal oxidation mode and generates a high H ₂ syngas stream (i.e., H ₂ » CO) and reactor 604 is operating in metal oxide reduction mode and produce a high CO syngas stream (i.e., CO » H ₂ ). As is discussed above, it is believed that the rate of metal oxidation will always be faster than the rate of metal oxide reduction within the reactors. Therefore, although illustrated as one reactor and one purification train for each reaction mode, more than one reactor (e.g., two or more reactors) and its associated purification train can be employed in the reducing mode for every reactor being used in the metal oxidation mode.

Steam for input to the metal oxidation reactor 602 is provided by heating purified water in the fluid bed roaster 626 and waste heat exchangers 610, 611 and 648. Prior to heating, the water should be subjected to purification 612 such as by using reverse osmosis and de-ionization to remove contaminants that can affect boiler operation or introduce impurities into the high H ₂ syngas stream. A specified quantity of steam produced in the waste heat exchangers is provided to the reactor 602, preferably at a super-heated temperature such as at least about 400 ⁰ C and more preferably at least about 500 ⁰ C. However, the steam temperature should not be greater than about 700°C, more preferably not greater than about 600 ⁰ C, and even more preferably not greater than about 55O ⁰ C. Higher temperatures are preferred to the extent that

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equipment comprised of common materials of construction can withstand the temperature and pressure of the steam without degradation.

The steam is preferably injected into the reactor 602 through a submerged lance, as is discussed above. Injecting the steam through a submerged lance(s) provides good mixing and a high contact surface area between the steam and the molten metal to promote the metal oxidation reaction. As the injection velocity is increased toward sonic velocity, at some point there is a dramatic decrease in the size and increase in the number of bubbles; this emulsion-like condition and associated high surface area of contact favors approaching the theoretical conversion of steam to hydrogen. Reactor 602 provides for an egress of H ₂ and un-reacted steam from the reactor, but otherwise is designed to contain the liquid alloy, slag and gases. Also, the reactor 602 may be placed under pressure to provide a sufficient contact time for the steam (to maximize H ₂ production) and to facilitate the further downstream pressurization of the high H ₂ syngas stream. Heat must be supplied to the metal oxidation reactor 602 to raise the temperature to the reaction temperature, since the modest amount of heat delivered by the steam- iron reaction at 1300 ⁰ C is insufficient to elevate the incoming steam to the reaction temperature. According to a preferred embodiment, oxygen (O ₂ ) from an air separation plant 614 is supplied by valves 615 to the reactor 602 to provide that heat by various oxidation reactions.

In the reactor 602, O ₂ generates heat by reacting with (burning) the metal to form a metal oxide. Another method for generating heat is by reacting O ₂ with the H ₂ just produced, yielding steam, which in turn can re-react with the metal to re-form the hydrogen and metal oxide. Both reactions decrease the amount of metal available to react with steam to form H ₂ ; it therefore it is preferred to provide the reactor with sufficient metal, over and above that required to produce the desired amount of H ₂ , to react with only sufficient O ₂ to provide the necessary heat.

Simultaneously, a hydrocarbon feed is provided to the metal oxide reduction reactor 604 to reduce the iron oxide contained in the slag back to iron which re- dissolves into the molten metal. In the metal oxide reduction reactor 604, O ₂ generates

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heat by supporting the partial oxidation of the carbon from the hydrocarbon feedstock to CO in the highly reducing environment of the reactor 604. By controlling the carbon-to- oxygen ratio, which in turn controls the ratio of oxidizing gases to total gases as is discussed above, the oxygen present can oxidize the carbon predominately to CO, while at the same time minimizing the formation of CO ₂ . To preclude potential corrosion, a carrier gas such as purified high CO syngas (preferably devoid of N ₂ ) may be used to dilute the O ₂ prior to injection into the reactor 604.

In the metal oxide reduction reactor 604, CO derives from two reactions: (1) the gasification of carbon with FeO _x (endothermic); and (2) the partial oxidation of carbon (exothermic). Employing gas-fired turbines 642, this CO can be advantageously used to generate electricity. The advantage is two-fold: (1) the CO to CO ₂ conversion (the second oxygen atom accepted by the carbon) embodies approximately two-thirds of the energy available from the complete oxidization of carbon; and (2) gas fired turbines, especially when operated as an Integrated Gas-Turbine Combined Cycle (IGCC) 652, comprise a means of generating electricity that is about 66 percent more efficient in converting thermal energy to electric energy than the conventional coal-fired, steam- driven-turbines that currently generate approximately 51 percent of the electrical needs of the United States.

Fluxes can be added to control the properties of the slag layer that forms above the molten metal mixture as the metal oxidation reaction oxidizes the metal in reactor 604. Additionally, other materials such as tin compounds, cassiterite ore or other materials such as iron compounds or ore may be added to make-up for losses of metal values. According to one embodiment, cassiterite ore (SnO ₂ ) is injected into the reactor to make-up for tin losses. A crude high H ₂ syngas stream is removed from metal oxidation reactor 602.

This crude hydrogen-containing syngas stream can include:

• H ₂ and un-reacted steam;

• minor amounts of CO and H ₂ S (which derive, respectively, from carbon and sulfur dissolved in the metal reacting with the injected steam);

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• furnace dust; and

• gaseous tin sulfide (SnS).

The crude high H ₂ syngas and crude high CO syngas streams also carry substantial heat values. These hot, crude syngas streams are passed through a quench 608 and 609 respectively where liquid water rapidly cools the syngas stream but without substantial loss of recoverable heat. Preferably, the quench 608 and 609 is a dry bottom quench wherein a controlled amount of liquid water rapidly cools the syngas stream to a reduced temperature, such as about 700 ⁰ C. Particularly for the high CO syngas stream, the rapid cooling is designed to minimize the Boudouard reaction that is favored above 700 ⁰ C and which consumes carbon (from the steel of the equipment walls) by reacting with CO ₂ to form CO. However, the farther the temperature is dropped below about 700 ⁰ C, the less heat that is available downstream for producing electricity. Thus, it is preferred to reduce the syngas stream to a temperature in the range of from about 900°C to about 700 ⁰ C. The use of such a quench to cool the two syngas streams advantageously:

1. Minimizes metal "dusting", i.e., the destruction of the containing ductwork, by rapidly dropping the temperature through the temperature range where the "dusting" reaction occurs (thought to be caused by the Bodouard reaction, wherein carbon contained in the steel plus CO ₂ yields CO; 2. Minimizes the potential for the inadvertent deposition of carbon from the reverse Boudouard reaction;

3. Shifts some CO to H ₂ by the water gas shift reaction. It is believed that this conversion may advantageously be catalyzed by the nascent iron oxide dust simultaneously expelled with the gases; 4. Cools the hot gases to more manageable temperatures and volumes without substantial loss of heat; and

5. Condenses any gaseous SnS to solid SnS.

The high H ₂ syngas and high CO syngas streams can then be passed through waste heat exchangers 610 and 611 to further cool the syngas streams and to provide

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heat for additional steam, thereby conserving heat values. For example, the temperature of the gas streams can be dropped to about 25O ⁰ C and the recovered heat used to generate steam.

It is also an advantage that the quench 608 uses water, which converts to steam and promotes the water gas shift (WGS) reaction, wherein CO "contaminating" the high

H ₂ syngas reacts with H ₂ O to form H ₂ and CO ₂ . It is also believed that this reaction is advantageously catalyzed by the presence of nascent, particulate metal oxide (e.g., iron oxide) in the reduced syngas stream, thereby increasing the amount of H ₂ produced.

After the quench, heat from the high H ₂ syngas can then be conserved by heat exchanger 610 and heat from the high CO syngas stream can be conserved by heat exchanger 611.

The high H ₂ syngas stream can also include some contaminants, such as CO and H ₂ S, both arising from carbon and sulfur dissolved in the metal reacting with the injected steam, and entrained particulates of (frozen) slag which are ejected from the molten metal bath and slag. Similarly, the high CO syngas stream can also include some contaminants including volatile tin sulfide (SnS) and entrained particulates of (frozen) slag which volatilize or are ejected from the molten metal bath and slag. The particulate contaminants can be removed from both syngas streams, such as by metal filters 616 and 617. Alternatively, other means such as electrostatic precipitators or bag houses can be used to separate particulates. For example, the volatile tin compounds (e.g., SnS), are condensed from the high CO syngas stream during the quench 609 and are captured in the metal filter 617 as solid SnS along with particulate slag.

After removal of contaminants, if any, both the high H ₂ syngas and high CO syngas streams can be treated in catalytic reactors 618 and 619 to convert carbonyl sulfide (COS) in the syngas streams to H ₂ S. This reaction is typically carried out at a temperature of about 200 ⁰ C. Other means to remove COS, such as physical solvents, can be used.

Thereafter, the high H ₂ syngas and high CO syngas streams can be cooled in a chillers 620 and 621 to condense excess steam and the water can be recovered and recycled. The chiller 621 can also advantageously remove soluble chlorides (and other

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halogen) compounds from the high CO syngas stream. Chlorine is a common contaminant in many of the types of hydrocarbons suitable as a feedstock for this process. The chlorine, released during pyrolysis, can react with the hydrogen to form gaseous HCI. The resulting high CO syngas stream from the chiller 621, except for trace amounts of H ₂ S, comprises a refined and relatively pure high CO syngas stream.

Similarly, the resulting high H ₂ syngas stream from chiller 620, except for trace amounts of H ₂ S and CO, comprises a refined and relatively pure high H ₂ syngas stream. Considerable heat is released as water is condensed from both the high H ₂ and high CO syngas streams, and this heat may be captured within a hot water header 640 for recycling within the steam system. The Cl ^" ion concentration in the condensed water is preferably controlled to preclude corrosion problems and assure continued adsorption of the extremely water soluble HCI gas.

Amine scrubbers and pressure swing adsorption units require elevated pressure for best operation. Therefore, compressor 654 follows chiller 621 and compressor 624 follows chiller 620. Chillers 620 and 621 are the last place where the equipment in the two (or more) gas trains is identical, since compressors for H ₂ and CO differ due to vast physical differences in the gases being compressed.

After compression, the high H ₂ syngas stream can then be passed through an amine scrubber 622 to remove H ₂ S. The H ₂ S-rich amine solution from the scrubber can be passed to an amine regenerator 630 to regenerate the amine solution which is then passed back to the amine scrubber 622 or 623. Other means for removing the H ₂ S, such as physical solvents (e.g., methanol), can also be used.

As previously mentioned, trace amounts of CO can arise from the reaction of the injected steam with the residual carbon dissolved in the molten metal in the metal oxidation reactor 602. This CO can optionally: (1) be removed in a pressure-swing absorption (PSA) unit 636; or (2) be converted to methane in a methane synthesis loop (not illustrated). Methane is advantageous as it is substantially benign with respect to most catalysts that may be used for subsequent hydrocarbon or ammonia synthesis. If the high H ₂ syngas stream is purified by a PSA unit 636, the rejected CO (and

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inadvertently about 10 to 15 percent of the H ₂ ) can then be transferred to the IGCC generator 652 and burned for electricity, as is discussed below.

The refined high H ₂ stream exiting the PSA unit 636 is high purity H ₂ and preferably comprises at least about 99 vol.% H ₂ , more preferably at least about 99.9 vol.% H ₂ . The high purity H ₂ stream can be moved to storage 650 or can be immediately delivered to an end user. Alternatively, the high purity H ₂ can be immediately utilized to produce high-value hydrocarbon end products, or other products such as ammonia.

The metal oxide that is produced by the metal oxidation reaction must be converted back to the reactive metal, and it can be reduced by carbon from the hydrocarbon feedstock. As illustrated in Fig. 7, metal oxides are reduced in reactor 604, preferably simultaneously with the metal oxidation in reactor 602.

Both carbon and H ₂ are released by pyrolysis of the hydrocarbon feedstock in reactor 604. One portion of the carbon serves as the reductant to render the metal oxides back to the metals and simultaneously generate CO by gasifying the solid carbon. Another portion of the carbon reacts with O ₂ which, by controlling the oxygen partial pressure, expressed as the ratio of oxidized gases to total gases by the fraction:

(H ₂ O + CO ₂ ) (H ₂ + H ₂ O + CO + CO ₂ ) controls the rate at which the reactive metal oxide is reduced back to the reactive metal. Advantageously, the highly reduced atmosphere that is required to reduce the reactive iron oxide (FeO _x ) contained in a slag of mixed oxides is also high in CO relative to CO ₂ and high in H ₂ relative to steam (H ₂ O). Advantageously, a high CO syngas stream comprised of CO and H ₂ , with lesser amounts of CO ₂ and trace impurities, results. A preferred hydrocarbon feedstock is a low-ash hydrocarbon high in its percentage of both carbon and hydrogen, and low in its percentage of moisture, oxygen and ash. Low ash reduces slag losses, low oxygen enhances the available (fuel bound) hydrogen, and low moisture permits a rapid rate for the reduction of the iron oxide. That is, fuels with a higher oxygen content such as municipal waste will consume some of the available H ₂ . In this instance, the hydrogen released as elemental hydrogen will be

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the total hydrogen minus about 1/8 of the oxygen. High carbon and hydrogen values minimize the total amount of fuel required while simultaneously producing a desirable syngas stream. A low moisture content in the hydrocarbon feed maximizes the rate of reduction of the iron oxide (and therefore the production of syngas). The hydrocarbon feedstock can be injected into the reactor 604 using a submerged lance or similar device. When so injected, the particulate feedstock can be entrained in a reducing gas such as CO, and during steady-state operation a portion of the purified and compressed high CO syngas stream can advantageously be recycled and used as a carrier gas. It is also possible, although less desirable, to add the hydrocarbon feedstock to the reactor 604 by other means, such as by simply dropping it into the reactor 604. The reactor 604 is preferably operated at a temperature that is similar to that of reactor 602 discussed above, such as from about 1300 ⁰ C to about 135O ⁰ C. As with reactor 602, O ₂ can be introduced into the reactor 604 and the flow of O ₂ from the air separation plant 614 can be controlled by valves 615. Introduction of oxygen may require a carrier gas, which desirably has little or no nitrogen. The amount of O ₂ introduced into reactor 604 is preferably just sufficient to support combustion of enough fuel to supply the endothermic heat required for the conversion of metal oxide to metal. Only minimal O ₂ is introduced into the reactor 604; off-gas from the metal oxide reduction reactor 604 is in a reduced state and includes mostly the same gases and contaminants listed above, but in different relative amounts, plus contaminants peculiar to the hydrocarbon feedstock that are not captured either in the alloy or slag.

Other materials such as fluxes can be injected into the reactor 604, for example to control the properties of the slag such as slag fluidity or tendency to foam. The ash- forming minerals that can be part of the hydrocarbon feedstock contribute to the slag layer within the reactor 604. When coal is used as a hydrocarbon feedstock and there is adequate calcium oxide (CaO) in the slag (either inherent in the coal or added as flux), the slag layer can be tapped from the metal oxide reduction reactor 604 and sold as a pozzolanic by-product. Preferably, the slag is tapped at a point when the iron oxide and tin oxide content in the slag are low so that iron and tin losses are minimized. As is discussed above, one aspect of the present invention anticipates recovering the residual tin and iron from the tapped slag. The process comprises

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mixing elemental sulfur, a plant by-product from the Claus plant 632, into the slag just after tapping and while the slag is still hot. Sulfur will react with either elemental tin (Sn) or tin dioxide (SnO ₂ ), whichever form is present, to create the volatile tin sulfide, SnS. If SnO ₂ is present, SO ₂ also will be formed. Both reactions are exothermic and the heat derived from the reactions can advantageously off-set thermal losses to the environment.

Immediately after tapping, the slag can be directed into a heated crucible or converter. Using lances, sulfur, irrespective of its state, can be blown through the hot slag using a recycle gas that is non-oxidizing with respect to sulfur and contains little or no nitrogen. The expected off-gas from the heated crucible will contain the volatile species, SnS and SO ₂ . These can be directed to the dry bottom quench 609, which is directly connected to the reactor 604 being tapped. Any excess sulfur above the stoichiometric requirement to react with any Sn and/or SnO ₂ in the slag will react with the iron to form iron sulfide. After recovery of residual tin from the slag, iron may be recovered from the FeO _x and FeS, if any, by admitting hydrocarbon and oxygen to the converter. To preclude slag freezing, additional heat may be introduced into the converter by burning natural gas.

Any SnS entering the quench can be captured in a baghouse and recycled after oxidation to SnO ₂ . Any SO ₂ entering the quench will react with the H ₂ to form H ₂ S and H ₂ O. The H ₂ S can be captured by the amine scrubber 623 and ultimately processed back to elemental sulfur in the Claus plant 632. The water can be removed by chilling.

After the "sulfur blow" to recover tin and addition of hydrocarbon and oxygen to recover the iron, the lances can be removed (retracted) and the slag quenched in a wet- bottom quench. This quickly freezes the slag and controls particle size, minimizing crushing requirements.

The series of unit operations from the reactor 604 through the amine scrubber

623 are substantially identical to those downstream of reactor 602 through the amine scrubber 622. Specifically, in terms of size, throughput and function, the following pairs of equipment can be identical: reactors 604 and 602; dry-bottom quench 609 and 608; heat exchanger and superheater 611 and 610; metal filter 617 and 616; catalysis of

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COS to H ₂ S 619 and 618; chiller 621 and 620 and amine scrubber 623 and 622. This is because, at the end of the cycle, when the desired amount of metal alloy is oxidized and the desired amount of metal oxide is reduced, the functions of the reactors 604 and 602 switch. The size of all unit operations is advantageously designed to handle the larger gas flow, which will be the high CO syngas stream from metal oxide reduction reactor 604

Thus, the high CO syngas stream from reactor 604 is treated in a quench 609 to reduce the temperature of the syngas stream, preferably to not greater than about 700 ⁰ C. This quench step advantageously prevents a series of otherwise untoward events, as previously listed, such as "dusting" of metals due to the presence of CO and or deposition of carbon on the surface of ductwork containing the gases. Particulates from the reduced off-gas are then removed in a filter unit 617. These particulates can include metal oxides (e.g., iron oxide), carbon and tin sulfide (SnS) when tin is used as a diluent metal in the reactor 604. Preferably, the SnS is conveyed to a dryer and pelletizer 628 and the agglomerated pellets are then treated in a roaster such as a fluidized bed roasting unit 626 to convert the SnS to SnO ₂ and SO ₂ through the introduction of O ₂ from the air separation unit 614. The SnS is preferably roasted in the roasting unit 626 in a manner that the O ₂ remaining in the roasting gas is minimized, so that little or no O ₂ is mixed with the SO ₂ coming off the roasting unit 626.

The SO ₂ can then be transferred to a Claus plant 632 where it is combined with the H ₂ S from the amine regenerator 630 or, if sufficient H ₂ S is not available for the

Claus reaction, H ₂ exiting the PSA system 636 can be used as the reductant. The

Claus plant 632 produces sulfur which is a salable by-product of the process. The SnO ₂ can advantageously be recycled to the reactor 604 to reduce tin losses from the system.

The high CO syngas stream is then treated in a catalytic unit 619 to convert COS in the gas to H ₂ S. This reaction can take place, for example, at about 25O ⁰ C. The high

CO syngas stream is then treated in a chiller unit 621 to reduce the temperature to preferably not greater than about 100°C, more preferably to not greater than about 8O ⁰ C.

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After compression 654 the high CO syngas stream is then treated in an amine scrubbing unit 623 to remove H ₂ S. The amine scrubber removes H ₂ S from the syngas stream and the H ₂ S-rich amine from the scrubbing unit 623 is passed to an amine regenerator 630 to regenerate the amine which is then passed back to scrubbing unit 622 or 623

The H ₂ S can then be combined with SO ₂ in a Claus plant 632 for the production of sulfur. It may also be desirable to divert a portion of the H ₂ to the Claus plant 632 since there may not be enough H ₂ S available to stoichiometrically match the SO ₂ from the roaster 626. The tail gas from the Claus plant 632 may be directed to the quench downstream from the reduction reactor for final gas clean-up (not shown in Fig. 7).

After amine scrubbing, the high CO syngas stream preferably includes predominately CO, H ₂ and CO ₂ . In the embodiment illustrated in Fig. 7, this reduced off-gas (previously compressed in unit 654) can be burned in a gas-fired turbine 642 to produce electricity via generator 644, and heat from gas-fired turbine 642 recovered in heat exchanger 648 can be used to raise steam to power steam turbine 646, thereby increasing electrical output. Some of this electricity can be used to operate different unit operations, such as air separation plant 614 and the various compressors. Excess electricity can be sold into the power grid. As is illustrated in Fig. 7, the unit operations 642, 644, 646 and 648 comprise an integrated gas turbine combined cycle (IGCC) generator 652.

Fig. 8 illustrates another embodiment of the present invention that, as compared to Fig. 7, is adapted to increase the production of high purity H ₂ gas while reducing the amount of electricity that is produced. In this embodiment, the high CO syngas stream from the metal oxide reduction reactor 604 is treated in a PSA unit 637 to recover the H ₂ contained in the high CO syngas stream and mix it with the high H ₂ syngas stream from the metal oxidation reactor 602. Specifically, a pressure swing absorption (PSA) unit 637 is provided to separate the H ₂ from the CO and CO ₂ in the high CO syngas stream. Thereafter, the H ₂ from PSA 637 can be mixed with the high H ₂ syngas stream from PSA unit 636 and the remaining high CO syngas stream can be provided to gas-fired

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turbine 642 to produce electricity. The high CO syngas may require compression in compressor 625 before entering the gas fired turbine 642.

In yet another embodiment of the present invention, the high CO syngas stream from the metal oxide reduction reactor 604 and the high H ₂ syngas stream from the metal oxidation reactor 602 can be commingled after the amine scrubbing units 622 and

623. The commingled gas stream can then be used to produce electricity, such as by providing the commingled syngas stream directly to a gas fired turbine, to maximize the production of electricity. Alternatively, the commingled gas stream can be used to produce additional hydrogen by utilizing the water gas shift reaction and PSA to isolate H ₂ from CO ₂ . Further, the two syngas streams can also be treated and/or commingled in various ratios to achieve the desired objective of the plant, i.e., to produce electricity,

H ₂ gas and or hydrogen containing products such as high-value fuels.

Specifically, the high H ₂ and high CO syngas streams can be utilized to produce other high-value products, such as methane (CH ₄ ) in a methanation unit. Other high value hydrocarbons such as gasoline, diesel fuel, jet fuel and the like can be produced using known Fischer-Tropsch processes or modifications thereof. To provide the correct 3:1 molar ratio of H ₂ to CO, a portion of the high CO syngas stream (containing CO, CO ₂ and H ₂ ) can be mixed with the high H ₂ syngas stream such that the combined stream has the proper ratio of H ₂ :CO. If the two syngas streams are blended to make a tailored syngas preparatory to synthesis of a hydrogen-containing commodity, then the PSA unit operation(s) is superfluous.

One aspect of the present invention is directed to the production of ammonia using the manufactured low-cost, high-purity hydrogen gas and nitrogen gas from the air separation plant as reactants. One of the important aspects of the method according to the present invention is the in-situ manufacture of large quantities of H ₂ at a relatively low cost. It is believed that one of the primary hindrances to the methods disclosed in the prior art for the production of ammonia is the need for large volumes of H ₂ gas and the high cost associated with the H ₂ gas. According to the present invention, high volumes of hydrogen gas can be economically generated in-situ.

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The nitrogen and hydrogen are combined in a H ₂ :N ₂ molar ratio of about 3:1 in order to maximize the production of ammonia (NH ₃ ). In a typical ammonia production method, a gas including hydrogen and nitrogen is compressed to about 200 atmospheres of pressure and passed over an iron catalyst at a temperature of from about 380 ⁰ C to about 450 ⁰ C.

The methods for gasification of the present invention can provide numerous advantages as compared to prior art gasification methods. Among these are:

■ reduced CO ₂ (a greenhouse gas) is produced per unit H ₂ or energy produced as compared to conventional gasification. CO ₂ emissions per unit of H ₂ produced are: conventional gasification about 22 tons CO ₂ per ton H ₂ ; the method of the present invention about 14 tons CO ₂ per ton H ₂ . Steam methane reformation (SMR) emits 13 tons CO ₂ per ton H ₂ , however, SMR is not desirable due to the high cost of its feedstock, natural gas. ■ low-value hydrocarbon fuels, including high-sulfur hydrocarbons or hydrocarbons containing chlorine, can be utilized.

■ solid or liquid hydrocarbons can be gasified to hydrogen and syngas in a fashion that permits tailoring the H ₂ :CO ratio of a combined gas to the preferred value for downstream operations. ■ the off-gas from the metal oxide reduction reactor is kept in a highly reducing state., typically with values of CO:CO ₂ in excess of those found in gases from conventional gasification.

■ overall H ₂ generation efficiency is extremely high.

■ the separate stream of H ₂ , uniquely produced by the present process, requires minimal or no gas clean-up, depending upon final use.

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Removal and Preclusion of Pollutants

In accordance with the foregoing method, two (or more) substantially identical trains of equipment can be used for (1) recovering heat; (2) removing pollutants; and (3) precluding the formation of pollutants from components comprising the syngas stream. Pollutants contained in the feed material which distribute to the gas phase (as opposed to the alloy or slag phase) determine what unit operations are required for removing the pollutants.

By way of example, listed below is a sequence of unit operations (with reference to Figs. 7 and 8) designed to remove pollutants that might be expected when utilizing coal as the carbon source for the process of the present invention. Pollutants removed include fine solid particulates such as furnace dust and SnS, water, chlorine, sulfur, mercury and CO. Pollutants whose formation is advantageously precluded include nitrogen oxides (NO _x ) and furans and dioxins.

1. Dry bottom quench 608/609: this unit is designed to rapidly decrease the gas temperature from 1300 ⁰ C to 700 ⁰ C by injecting liquid water. The purpose of the quench is to preclude a potential metallurgical problem known as dusting, which is the deterioration of the metal that contains the gas and is known to occur at temperatures above about 700 ⁰ C in syngas streams with a high concentration of CO.

2. Heat exchanger and steam super heater 610/611: This unit operation is a conventional heat exchanger and super heater designed to recover the sensible heat of the gases.

3. Metal filter 616/617: After the gas is cooled, particulates are removed by a filter, such as a candle filter employing a metal filter medium. The fine solids that are recovered are comprised of furnace dust and tin sulfide.

4. Fluid bed roaster 626: In a unit external to the purification trains, the particulates collected from metal filter 616/617 are roasted in a

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fluid bed roaster in oxygen to form SO ₂ and SnO ₂ . The SnO ₂ is returned to the furnace with the dust. The SO ₂ advances to a Claus plant 632, where the SO ₂ is combined with H ₂ S, recovered from the amine regeneration system or H ₂ from the product line to form elemental sulfur.

5. COS to H ₂ S 618/619: In this unit operation, a catalyst is used to hydrolyze the carbonyl sulfide to hydrogen sulfide and carbon dioxide, and is part of the sulfur removal system. Sulfur is a ubiquitous contaminant of coal and other hydrocarbons. 6. Chiller 620/621: The chiller is designed to remove steam originating from two sources: (i) steam that was not converted to H ₂ in the reactor oxidizing iron to iron oxide, where the theoretical conversion of steam to hydrogen is about 75%; and (ii) water added to the dry bottom quench. Acid gases such as HCI or HF also will be removed by the chiller due to their high solubility in water. Their removal is related to precluding formation of furans and dioxins.

7. Amine scrubber 622/623: This standard unit operation is part of the system for removing sulfur, and it operates in conjunction with the amine regeneration unit 630. 8. Activated carbon adsorber (not illustrated): Mercury can be adsorbed by activated charcoal, and can be recovered from the loaded carbon and the charcoal reactivated and reused.

9. Pressure swing adsorption (PSA) system 636/637: This system is a standard means for disengaging commingled gases and typically is used to separate H ₂ from CO and CO ₂ to yield a pure H ₂ stream. Removal of CO is required for ammonia production since it fouls the ammonia conversion catalyst. An alternative to removing the CO is to convert it to methane (CH ₄ ) in a methanation synthesis loop. Methane characteristically does not foul catalysts.

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Dioxins are a family of compounds known as polychlorinated dibenzo-dioxins (PCDD), and furans are a family of compounds known as polychlorinated dibenzofurans (PCDF). There are about 210 compounds in these two families, and they have a wide range of environmental, chemical and physical properties. Two methods are postulated for their formation. Both methods are believed to require the presence of all of the following precursor conditions: (1) the presence of solid particles containing carbon structures; (2) the presence of organic or inorganic chlorine; (3) the presence of iron, copper, manganese or zinc ions; (4) an oxidizing atmosphere; and (5) a temperature window of 250 ⁰ C to 400 ⁰ C. According to the present invention, the formation of dioxins and furans is substantially precluded because oxygen is absent as the hot syngas streams are cooled down through the temperature window of 400°C to 250 ⁰ C, and chlorine (Cl) and fluorine (F) are removed before the gases are reheated up through that temperature window.

The formation of nitrogen oxides can also be precluded or reduced by reducing the oxidation temperature inside the gas fired turbine. Water or other oxidized gases such as CO ₂ can be used as a temperature control method.

There are numerous methods for removing elements that are considered pollutants from the two syngas streams. In terms of what potential pollutants to remove, the starting point is an analysis of the solid hydrocarbon feed being used and all other materials that enter the process. The method of the present invention advantageously partitions all elements admitted to the process to one of four locations - the slag, the alloy, the dust, or the syngas streams.

Slag

Refractory oxides such as SiO ₂ , AI ₂ O ₃ , CaO and MgO report to the slag. Control of slag properties depends upon taking a sufficiently large slag tap to preclude the buildup of potentially detrimental elements. For example, the ash from petroleum coke typically has a high percentage of vanadium. Vanadium is expected to report to the slag as vanadium oxide. The vanadium content of the slag can be kept within pre- established limits, preferably less than 20 percent for satisfactory slag properties, by adjusting the amount of flux added and the size of the tap taken.

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Alloy

Some elements that may be associated with the solid hydrocarbon feed, such as oxides of nickel or copper, are expected to be reduced by the carbon and report to the alloy. These elements dilute the alloy but do not render it ineffective. After some time, elements (other than tin and iron) accumulate in the alloy, and the entire alloy will have to be changed out. Value received from the "contaminated" alloy likely will exceed the cost of a fresh alloy system.

Reactor Dust

This material, extracted from the syngas stream by filtration, is agglomerated and then roasted to produce dry SO ₂ , for sulfur production by the Claus process, and calcine, for returning iron and tin oxides to the reactor that ejected the dust.

Syngas streams

Solid hydrocarbons derive from living materials and they are comprised principally of carbon, hydrogen, nitrogen, oxygen, sulfur, chlorine and ash. Ash may be inherent, comprised of inorganic elements commonly associated with the living material, or the ash may be adventitious, washed in from another source. Most ash components are expected to partition to the slag with a few partitioning to the alloy.

The only solid hydrocarbon currently available in North America in sufficient quantity to off-set the use of imported oil is coal. Coal can be augmented with pet coke, municipal waste, and rubber tires, either to enhance the quality of the hydrocarbon or to consume waste thereby reducing or eliminating landfills and their associated ills. The major potential contaminants arising from these hydrocarbons are considered below.

From solid hydrocarbon, water and air, the method of the present invention creates both a high CO syngas stream, issuing from the "reduction" reactor, and a high H ₂ syngas stream, issuing from the "oxidizing" reactor. Both syngas streams are highly reducing (low partial pressure of oxygen), hot, dusty and can be contaminated with chlorine, sulfur compounds and various other elements. Also, both gas streams are passed through a series of gas treatments that can render purified hydrogen and syngas

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Precluding Corrosion. A high concentration of CO at high temperatures can cause corrosion or "dusting" of the metal ductwork containing the gas. For this reason the 1300 ⁰ C high CO syngas from the furnace is rapidly cooled by injecting sufficient liquid water to reduce the temperature to 700 ⁰ C. Conserving Heat. This is critical for maintaining good efficiency. Standard heat exchangers are used for this purpose.

Particulate Removal. Devices effective in removing particulates from a syngas stream can include: ESPs (electro-static precipitators), metal (candle) filters and (cloth) bag-houses. Metal filters are generally preferred because they can withstand the (relatively high) gas temperature. Two types of (intermingled) particulates are removed; furnace dust and tin sulfide. Tin sulfide along with associated furnace dust is subsequently roasted to recover the tin as tin dioxide, which, along with the furnace dust, recycles to the reactor, and sulfur dioxide which reports to the Claus plant.

Acid Gas Removal. Acid gases include CO ₂ (as H ₂ CO ₃ ); hydrogen sulfide (H ₂ S) and hydrogen chloride (HCI) and/or hydrogen fluoride. Various methods for their removal include:

CO _∑ . Scrubbing with a solvent, such as the RECTISOL process (Lurgi, Frankfurt, Germany), which removes the CO ₂ from the syngas stream, to be subsequently released (upon regeneration of the solvent) and optionally can be compressed and sequestered. The RECTISOL process uses cold methanol as a physical solvent and the CO ₂ (as well as H ₂ S, COS and other sulfur compounds) are removed from the syngas stream. Alternatively, the high CO syngas stream can be burned in pure oxygen yielding a more-or-less pure CO ₂ , again for compression-sequestration. Other methods for removing CO ₂ exist including adsorption on activated carbon. Sulfur. Carbonyl sulfide (COS) is not removed by conventional amines. For this reason it typically is first hydrolyzed into H ₂ S:

COS + H ₂ O → H ₂ S + CO ₂ (18) This reaction proceeds well in the presence of a catalyst.

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Some solvents can remove both H ₂ S and COS. An example is SELEXOL (Union Carbide), a physical solvent made of a dimethyl ether of polyethylene glycol.

Halogen Acids. Halogen acids such as HCI can be removed in the method of the present invention due to their extreme solubility in water. In the chiller, the large amount of water (added at the quench where it is flashed to steam) is condensed. Formation of the cloud of condensed water will dissolve and remove the halogen acids from the syngas stream (and also some H ₂ S). If halogen acids are not removed before amine scrubbing, they will react (destructively) with the amine.

NO _x The formation of appreciable amounts of NO _x is precluded during combustion of the gas in the gas turbine by the addition of sufficient water to control flame temperature to below the temperature that is required for its formation.

Mercury. Mercury (Hg) is not present in the pet coke. Mercury, however, does exist in coal. Commercial methods for its removal have been developed employing

(powdered or granular) activated charcoal for adsorption with regeneration of the activated charcoal achieved by the application of mild heat to the sorbent. Such methods can be employed when mercury is present in the hydrocarbon feed.

Dioxins and Furans. These toxic compounds (collectively about 210 of them) do not exist in the feed; rather, they can form during cooling or heating of the syngas stream as it passes through the temperatures window of 25O ⁰ C to 400 ⁰ C, when all four of the following are present - oxygen, a carbon structure, chlorine and iron (as a catalyst). Absence of any one of these components will preclude formation.

The method of the present invention is advantageously arranged so that O ₂ is absent as the temperature drops from 400 ⁰ C to 25O ⁰ C and chlorine is absent as the temperature is raised from 250 ⁰ C to 400 ⁰ C in the IGCC circuit. The ability to preclude furan and dioxin formation is critical if municipal waste or coal with high chloride content, Illinois coal for example, is used as a feedstock. Municipal waste can be especially high in chlorine content (from PVC, household bleach and other sources).

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In summary, there is often more than one means of removing a pollutant from a reducing syngas stream; however, once selected, integration into the gas purification train is required.

In another embodiment of the present invention, CO ₂ can be removed from the atmosphere and sequestered. This embodiment can potentially create revenue in the form of CO ₂ credits that are available in several industrialized nations.

According to the foregoing description, methane (CH ₄ ) can be manufactured from low-cost hydrocarbons, water and air, such as by combining the high H ₂ gas stream and the high CO gas stream in the proper ratio. The low-cost hydrocarbon can be biomass. It is well established that methane can be decomposed into carbon black (or carbon fibers) and hydrogen, particularly at temperatures above 600 ⁰ C. High temperature (e.g., greater than 800°C) catalysts and membrane technology all favor the decomposition reaction.

The hydrogen released by the decomposition is high purity and can be used for many purposes, such as the manufacture of ammonia, refining and upgrading of crude oils, and the like. Further, the carbon black can be used in commerce, as a structural material and filler for plastics and rubber, or can be stored underground.

Thus, biomass (that otherwise would decompose by oxidation in a landfill) is converted to H ₂ and carbon, with very few by-products. Every ton of carbon sequestered is equivalent to removing 3.7 tons of CO ₂ from the atmosphere. Since oxidizing biomass does not create CO ₂ emissions (oxidizing biomass simply returns CO ₂ to the atmosphere that was first removed to create the biomass), sequestering carbon derived from biomass essentially removes CO ₂ from the atmosphere.

Even if the hydrocarbon feedstock is not biomass or municipal waste, the process as described will reduce or preclude CO ₂ from entering the atmosphere.

EXAMPLES

The following three examples are each based on the objective of producing 100 million standard cubic feet of hydrogen per day (100 MM scfd). In meeting that

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objective, all three examples produce a high CO syngas stream and a high H ₂ syngas stream. The distribution of the output energy from the process between H ₂ and electricity is dependent upon the treatment of the two syngas streams downstream of the amine scrubbers 622/623.

Example 1

This example illustrates an embodiment of the process that delivers most of the output energy as electricity (726 MW net), however, still produces 100 MM scfd of H ₂ . This example follows the sequence of unit operations as depicted in Fig. 7 and will be described with reference to Fig. 7.

All reactors have a 6.5 meter hearth diameter and 8.6 meter barrel diameter.

There are two metal oxide reduction reactors (e.g., reactor 604) and both have a 50 minute cycle for reducing FeO to Fe and they are staggered in time. There is one metal oxidation reactor (e.g., reactor 602) that has a 25 minute cycle time for oxidizing Fe to FeO. The plant operates 7884 hours per year (i.e., 90 percent availability).

At the start of the cycle, two metal oxide reduction reactors (referred to as 604a and 604b) each contain 293 tonnes (metric tons) of an alloy comprised of 40% iron and 60% tin, and 540 tonnes of slag comprised of 65% FeO, 12.4% CaO, 15.4% SiO ₂ and 7.2% ash (all percentages in these examples are in wt.%, unless otherwise indicated). These mass and composition conditions are also the conditions at the end of a cycle for metal oxidation reactor 602 which oxidizes Fe to FeO.

At the end of a metal oxide reduction cycle, each reactor 604 contains 503 tonnes of an alloy comprised of 65% iron and 35% tin, and 271 tonnes of slag comprised of 30% FeO, 24.6% CaO, 30.7% SiO ₂ and 14.7% ash. These conditions (mass and composition) are also the conditions at the beginning of the cycle for metal oxidation reactor 602.

During the process, feed materials are input to each of the reactors. The feed materials going into each of the two metal oxide reduction reactors 604 that produce the high CO syngas stream are illustrated in Table 4.

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Table 4 Input to each Metal Oxide Reduction Reactor (per reactor)

The temperature in reactors 604a and 604b is 1300 ⁰ C. The carrier gas is a portion of the flue gas extracted downstream during "steady state" operations, and, after compression, is used to carry the pet coke into the reactor. The total amount of carrier gas used is 65 Nm ³ (normal cubic meters) per tonne of solids. To insure that the desired reactions occur, the partial pressure of oxygen in reactors 604a and 604b is kept low. More specifically, the reactors 604a and 604b are operated such that the overall oxidizing potential, based upon the partial pressures of the gases:

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(H ₂ O + CO ₂ )

(H ₂ + H ₂ O + CO + CO ₂ ).

is maintained at 0.229, a value that is based on maximizing the rate of the reduction reaction (Gibb's free-energy minimization), which reaction also depends upon the chemical composition of the pet coke. Control of this ratio is established by adjusting the ratio of oxygen to hydrocarbon (pet coke) injected into the reactor.

The materials exiting each of the reactors 604 are illustrated in Table 5.

Table 5 Output for each Metal Oxide Reduction Reactor (per reactor)

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Under these conditions, a high CO syngas stream exits each reactor 604 at 298,450 Nm ³ /hr (normal cubic meters per hour). These outputs are each transferred to a dry bottom quench 609, where water is introduced to cool the materials.

For simplicity of illustration, the high CO syngas departing the two reactors 604a and 604b will be treated as a single gas stream, although in practice the two gas streams will be individually treated until the gases are sufficiently cooled that they can be combined. Table 6 illustrates the quantity and quality of the products entering the dry bottom quench 609 from the combined output of the two metal oxide reduction reactors.

Table 6 Combined Input to Dry Bottom Quench

In the quench, the temperature of the products drops to 700 ⁰ C, the high CO syngas stream picks up steam, and the dust picks up SnS from condensation of volatile tin compounds. Table 7 illustrates the output from the dry bottom quench 609.

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These materials are introduced to the heat exchanger and superheater 611, and Table 8 illustrates the quantity and quality of the products departing the heat exchanger and superheater 611.

Table 8 Output from Heat Exchanger and Superheater

Steam at 500 ⁰ C is generated by separately introducing water to the heat exchanger and superheater 611, which extracts heat from the high CO syngas. The

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discharge from the heat exchanger and superheater 611 is fed to a metal filter 617. Table 9 illustrates the quantity and quality of the products departing the metal filter 617.

Table 9 Output from Metal Filter

The high CO syngas stream can then be treated in unit 619 with the aid of a catalyst to convert any trace amounts of COS to H ₂ S The high CO syngas stream is then directed to a chiller and condenser 621 to remove H ₂ O as liquid water, and Table 10 illustrates the quantity and quality of the products exiting the chiller 621.

Table 10 Output from Gas Chiller and Condenser

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After compression 654 the high CO syngas stream is moved to an amine scrubber 623 to remove H ₂ S. The quantity and quality of the products exiting the amine scrubber 623 are illustrated in Table 11.

Table 11 Output from Amine Scrubber

The high CO syngas stream from the amine scrubber 623 has a molar (volumetric) CO: H ₂ ratio of 5.4. A portion (about 26 tonnes/hr total) of this high CO syngas stream is returned to the two reactors operating in metal oxide reduction mode as a carrier gas for carrying solids into the reactor.

In this example, the remaining high CO syngas stream is combined with a reject gas stream from the pressure swing absorption (PSA) operation 636 (discussed below) and with air. The composition of this combined gas stream is listed in Table 12.

Table 12 Input to Combined-Cycle Turbine

The combustion of this gas in the combined cycle turbine produces 829.2 MW of electricity. Table 13 illustrates the composition of the flue gas that is emitted by the combined cycle turbine.

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Table 13 Flue Gas Composition from Combined Cycle Turbine

The previously described equipment that produces the high CO syngas stream, reactors 604a/604b through the amine scrubber 623, is identical to the equipment that produces high H ₂ syngas stream, reactor 602 through the amine scrubber 622.

Tables 14 through 21 illustrate the quantity and composition of materials entering and/or leaving the equipment that produces the high H ₂ syngas stream, beginning with reactor 602 through the amine scrubber 622. Table 15 illustrates the feed material going into metal oxidation reactor 602.

Table 14 Input to the Metal Oxidation Reactor

The total flow rate of the O ₂ into the reactor 602 is 18,400 NrrrVhr. Assuming a 70% steam to H ₂ conversion ratio, the amount and composition of the output from reactor 602 and the input to the dry bottom quench 608 is illustrated in Table 15.

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Table 15 Input to Dry Bottom Quench

The total flow rate of the high H ₂ syngas stream from the reactor 602 is 235,060 Nm ³ /Hr. The quench water is added separately for reducing the temperature of the high H ₂ syngas stream and the dust. The output from the dry bottom quench is illustrated in Table 16.

Table 16 Output from Dry Bottom Quench

In the quench, the temperature of the products drops to 700 ⁰ C and the high H ₂ syngas stream picks up additional steam. As compared to dry bottom quench 609, little or no additional SnS from volatile tin compounds is expected in the gas stream.

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These materials are introduced to the heat exchanger and superheater 610 and Table 17 illustrates the quantity and quality of the products departing the heat exchanger and superheater 610.

Table 17 Output from Heat Exchanger and Superheater

Superheated steam at 500 ⁰ C is generated by separately introducing steam to the heat exchanger and steam superheater 610, which extracts heat from the high H ₂ syngas and the dust. The syngas stream and dust from the heat exchanger and superheater 610 are fed to a metal filter 616. Table 18 illustrates the quantity and quality of the products departing the metal filter 616.

Table 18 Output from Metal Filter

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The products exiting the metal filter can be treated in unit operation COS to H ₂ S 618, which, with the aid of the catalyst, converts carbonyl sulfide to hydrogen sulfide.

The high H ₂ syngas stream is then directed to a chiller and condenser 620 to remove H ₂ O as liquid water and Table 19 illustrates the quantity and quality of the products exiting the chiller and condenser 620.

Table 19 Output from Gas Chiller and Condenser

The high H ₂ syngas stream is then moved to an amine scrubber 622 to remove H ₂ S. The quantity and quality of the products exiting the amine scrubber 622 are illustrated in Table 20.

Table 20 Output from Amine Scrubber

The molar (volumetric) H ₂ :CO ratio of the high H ₂ syngas stream removed from the amine scrubber 622 is 70.4. Continuous production of H ₂ and electricity requires cyclic operation. In this example, the cycle time is approximately 25 minutes. At the end of the cycle, all feed materials entering reactor 604 are transferred to enter reactor

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602 and visa versa. Similarly, after a slight delay, compressed gas from compressor 654 enters amine scrubber 622; and compressed gas from compressor 624 enters amine scrubber 623.

The purified high CO syngas stream from amine scrubber 623 (Table 11 , CO: H ₂ molar ratio = 5.4) and the purified high H ₂ syngas stream from amine scrubber 622

(Table 20, H ₂ :CO molar ratio = 70.4) can be combined to form various precursor gases with a target H ₂ :CO ratio in the range of 0.19 (reciprocal of 5.4) to 70.3 for manufacturing gaseous fuels, liquid fuels, fertilizer, or the like.

In this example, however, CO is removed from the high H ₂ syngas stream in a pressure swing adsorption (PSA) system 636. The resulting pure H ₂ can advance to storage 650. The CO that is rejected by the PSA system 636 (which typically also includes 10 to 15 percent H ₂ ) is routed to the high CO syngas stream entering compressor 654 and then gas-fired turbine 642.

With respect to other unit operations, an air separation unit 614 is provided to produce industrial grade oxygen used in the process and air which is compressed as it enters the gas-fired turbine 642.

Also, dust from metal filters 617 and 616, comprised of fine particles ejected from the reactors and condensed SnS, enters a pelletizer and dryer 628 before proceeding to the fluid bed roaster 626. In the fluid bed roaster 626, the SnS is oxidized with oxygen from the air separation unit 614 to SnO ₂ and SO ₂ . The calcine (furnace dust and SnO ₂ ) is routed to reactor 604. The SO ₂ advances to the Claus plant 632 where it is joined with H ₂ S recovered from the amine regeneration unit 630 and sufficient additional H ₂ from the PSA system 636 to reduce the SO ₂ to elemental sulfur.

All water used to produce steam (condensate from the steam turbines 638 and 646, cooling tower 634 return water and make-up water) is processed in water purification unit 612 to make boiler-quality feed water. That portion of 500 ⁰ C, 1800 psi steam, raised from heat recovered from heat exchangers 610, 611 and 648, fluid bed roaster 626, Claus plant 632 and chillers 620 and 621, that is not directed toward making H ₂ advances to steam turbine 638 to produce additional electricity by generator 640.

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The generation of electricity by the process is summarized in Tables 21a and

21b.

Table 21a Electricity Generation - Combined Cycle Turbine

Table 21b - Electricity Generation - Excess Steam Turbine

The total electricity generated by the combined cycle turbine and the excess steam turbine is 870 MW, or 6,859,080 MW-hrs. The amount of electricity consumed by the process is summarized in Table 22.

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Table 22 Total Electricity Consumed by Process

The total electricity consumed by the process is 144 MW, or 1 ,136,424 MW-hrs. Table 23 summarizes the total salable co-products from the process.

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Table 23 Net Salable Co-Products from Process

Example 2 Example 2 illustrates an embodiment of the process that also makes 100 MM scfd of H ₂ . Through the recovery of fuel-bound H ₂ (derived from the hydrocarbon feed), this example focuses on recovering all available H ₂ ; consequently, less electricity is produced as compared to Example 1. This example follows the sequence of unit operations as depicted in Fig. 8. All reactors have a 4.2 meter hearth diameter and 5.6 meter barrel diameter.

These reactors are smaller than the reactors of Example 1 as the total pet coke feed (illustrated below) is less for this Example 2. There are two metal oxide reduction reactors 604 (referred to as 604a and 604b); both have a 50 minute cycle for reducing FeO to Fe, and they are operated in a staggered mode. There is one metal oxidation reactor 602; it has a 25 minute cycle time for oxidizing Fe to FeO. The plant operates 7884 hours per year (i.e., 90 percent availability).

At the start of the cycle, metal oxide reduction reactor 604 contains 121 tonnes of an alloy comprised of 40% iron and 60% tin and 223 tonnes of slag comprised of 65% FeO, 12.5 % CaO, 15.7% SiO ₂ and 6.8% ash. The prior listed conditions (mass and composition) are also the conditions at the end of a cycle for metal oxidation reactor 602 which converts Fe to FeO.

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At the end of a cycle, reactor 604 contains 208 tonnes of alloy comprised 65% iron and 35% tin and 112 tonnes of slag comprised of 30% FeO, 25.0% CaO, 31.2% SiO ₂ and 13.8% ash. The prior listed conditions (mass and composition) are also the conditions at the beginning of the cycle for reactor 602.

Feed materials going into the metal oxide reduction reactor 604 that produces the high CO syngas stream are listed in Table 24.

Table 24 Input to each Metal Oxide Reduction Reactor (per reactor)

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The temperature in reactors 604 is 1300 ⁰ C. The carrier gas (the high CO syngas extracted before the gas turbine) is provided at a rate of 65 Nm ³ per tonne of solids and carries the pet coke, dust recycle and make-up into the reactor. The same oxidation potential is maintained as in Example 1. The materials exiting each reactor 604 are shown in Table 25.

Table 25 Output from each Metal Oxide Reduction Reactor (per reactor)

Under these conditions, the high CO syngas stream exits each reactor 604 at 135,374 Nm ³ /Hr. These outputs are each transferred to a dry bottom quench 609, where water is introduced to cool the materials.

For simplicity of illustration, the high CO syngas streams departing the two reactors 604a and 604b will be treated as a single gas stream, although in practice the two gas streams will be individually treated until the gases are sufficiently cool that they can be combined. Table 26 illustrates the quantity and quality of the products entering

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the dry bottom quench 609 from the combined output of the two metal oxide reduction reactors.

Table 26 Input to Dry Bottom Quench

Table 27 illustrates the output from the dry bottom quench 609. In the quench, the temperature of the products drops to 700 ⁰ C, some of the high CO syngas stream reacts with steam to form additional H ₂ and the dust gains particulate SnS from condensation of volatile tin compounds.

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Table 27 Output from dry bottom quench

These materials are introduced to the heat exchanger and superheater 611, and Table 28 illustrates the quantity and quality of the products departing the heat exchanger and superheater 611.

Table 28 Output from Heat Exchanger and Superheater

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Steam superheated to 500 ⁰ C is generated by introducing lower temperature steam into the heat exchanger and superheater 611, which extracts heat from the high CO syngas. Table 29 illustrates the quantity and quality of the products departing the metal filter 617.

Table 29 Output from Metal Filter

These products can then be treated to remove COS, and Table 30 illustrates the quantity and quality of the products exiting the unit operation COS to H ₂ S 619, which, with the aid of a catalyst, converts carbonyl sulfide to hydrogen sulfide.

Table 30 Output from COS to H ₂ S Unit Operation

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The high CO syngas stream is then directed to the chiller and condenser 621 removes H ₂ O as liquid water. Table 31 illustrates the quantity and quality of the products exiting the chiller 621.

Table 31 Output from Gas Chiller and Condenser

The high CO syngas stream, and the reject stream from PSA System 636, are compressed in unit 654 which occurs before the combined gases are moved to an amine scrubber 623 to remove H ₂ S. The quantity and quality of the products exiting from the amine scrubber 623 are shown in Table 32.

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Table 32 Output from Amine Scrubber

The molar (volumetric) ratio of CO: H ₂ in the high CO syngas stream exiting the amine scrubber is 5.4. The previously compressed high CO syngas stream from the amine scrubber 623 and the previously compressed reject gas stream from the pressure swing absorption (PSA) operation 636 of the high H ₂ syngas stream (discussed below) advance to the PSA system 637.

The compressor 654 requires 26.6 MvV of electricity to compress the gases to about 2750 kPa (400 psi). The composition of high CO syngas stream exiting amine scrubber after joining the reject stream from PSA system 636 is illustrated in Table 33.

Table 33 High CO Syngas Stream Exiting PSA

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In this Example 2, the output from the compressor 654 and amine scrubber 623 is then subjected to a PSA unit operation 637. The PSA operation 637 separates the H ₂ from the remaining syngas stream and the H ₂ can be advanced to storage 650. Table 34 illustrates the output from this PSA operation 637.

Table 34 Output from PSA

It should be noted that the process illustrated in this Example 2 differs from the process of Example 1 in this regard. Specifically, the process of Example 2 which produces additional H ₂ incorporates a second PSA system to extract H ₂ from the high CO syngas stream, as well as recovers hydrogen "lost" from the high H ₂ syngas stream by PSA unit 636 to the reject CO stream. Thus, additional hydrogen can be provided for other purposes.

After compression (not shown), the CO-rich reject stream from PSA 637 can then advance to the gas turbine. As shown in Table 35, a portion of the compressed gas exiting the PSA system 637 is used as a carrier gas to carry materials into the metal oxide reduction reactors.

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Table 35 Output from PSA system 637

Output from the PSA system 637, less carrier gas recirculating to reactors 604a and 604b, advances to the IGCC system 652 where it is combined with compressed air and combusted.

Table 36 Input to Turbine

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The combined cycle turbine generates 244 MW of electricity from these input gases. Also, although not listed above, some water is input to the turbine to control temperature and preclude NO _x formation. The composition of the flue gas exiting the IGCC system 652 is illustrated in Table 37.

Table 37 Output from the Combined Cycle Turbine System

The equipment that was previously described and which produces the high CO syngas stream, reactors 604 through the amine scrubber 623, is duplicated by equipment that produces the high H ₂ syngas stream, reactor 602 through the amine scrubber 622.

Tables 38 through 44 present the quantity and composition of materials entering and/or leaving the equipment that produces the high H ₂ syngas stream, beginning with reactor 602 through the amine scrubber 622.

Table 38 Input to Metal Oxidation Reactor

It is assumed that the conversion rate of steam to H ₂ in the reactor is 70 mol.%. The flow rate of O ₂ into the reactor is 18,400 NM ³ /hr. The amount and composition of the output from reactor 602 that is input to the dry bottom quench 608 is illustrated in Table 39.

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Table 39 Input to Dry Bottom Quench

Table 40 illustrates the output from the dry bottom quench.

Table ; 40 Output from Dry Bottom Quench

H ₂ - 11.1 %

H ₂ O - 86.2%

High H ₂ Syngas stream CO - 2.5%

(700 ⁰ C) N ₂ - 0.04%

H ₂ S - 0.03%

Total 53.2

FeO - 65.0%

CaO - 12.4%

Dust

SiO ₂ - 15.6% (700 ⁰ C)

Ash - 7.0%

Total 0.95

In the quench, the temperature of the products drops to 700 ⁰ C and the high H ₂ syngas stream picks up additional steam. As compared to dry bottom quench 609, little or no additional SnS from volatile tin compounds is expected in the gas stream. These materials are introduced to the heat exchanger and superheater 610 and

Table 41 illustrates the quantity and quality of the products departing the heat exchanger and superheater 610.

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Table 41 Output from Heat Exchanger and Superheater

Steam superheated to 500 ⁰ C is generated by introducing lower temperature steam into the heat exchanger and superheater 610, which extracts heat from the high H ₂ syngas and the dust. The syngas stream and dust from the heat exchanger and superheater 610 are fed to a metal filter 616. Table 42 illustrates the quantity and quality of the products departing the metal filter 616.

Table 42 Output from Metal Filter

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The products exiting the metal filter can be treated in unit operation COS to H ₂ S 618 which, with the aid of a catalyst, converts carbonyl sulfide to hydrogen sulfide.

The high H ₂ syngas stream is then directed to a chiller and condenser 620 to remove H ₂ O as liquid water, and Table 43 illustrates the quantity and quality of products existing the chiller and condenser 620.

Table 43 Output from Gas Chiller and Condenser

After compression 624 the high H ₂ syngas stream advances to an amine scrubber 622 to remove H ₂ S. The quantity and quality of the products exiting the amine scrubber 622 are illustrated in Table 44.

Table 44 Output from Amine Scrubber

The high H ₂ syngas stream has a molar (volumetric) H ₂ :CO ratio of 70.4.

Continuous production of hydrogen and electricity requires cyclic operation. In this example, cycle time is approximately 25 minutes. At the end of the cycle, all feed materials entering reactor 604 are transferred to enter reactor 602 and visa versa.

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Similarly, after a slight delay, the gas exiting amine scrubber 623 is directed to PSA unit 636 and the gas exiting amine scrubber 622 is directed to PSA unit 637.

The purified high CO syngas stream from amine scrubber 623 (CO: H ₂ molar ratio = 1.49) and the purified high H ₂ syngas stream from amine scrubber 622 (H ₂ :CO molar ratio = 61.5) can be combined to form various precursor gases for manufacturing gaseous fuels, liquid fuels, fertilizer or the like.

Characteristic of this Example 2, CO is removed both from the high CO syngas stream in its pressure swing adsorption (PSA) system 637 as well as from the high H ₂ syngas stream by PSA 636. The resulting pure H ₂ , the final product, advances to storage 650.

The CO (and inadvertently a small amount of H ₂ ) that is removed by PSA system 636 is routed to the high CO syngas stream, and then proceeds to PSA 637 (where some of the inadvertently lost hydrogen may be recaptured) and then to the integrated gas turbine combined cycle (IGCC) circuit 652. An air separation unit 614 is provided to produce industrial grade oxygen used in the process and air provided to the gas-fired turbine. Preferably, the oxygen gas from the (cryogenic) air separation unit comprises at least 98.5 vol.% O ₂ and not greater than 1.5 vol. % N ₂ .

Dust from metal filters 617 and 616, comprised of fine particles ejected from the reactors and condensed tin sulfide, enters a pelletizer and dryer 628 before proceeding to the fluid bed roaster 626. In the fluid bed roaster 626, the tin sulfide is oxidized with oxygen from the air separation unit 614 to tin dioxide and sulfur dioxide. The calcine (furnace dust and tin dioxide) is routed to reactor 604. The sulfur dioxide advances to the Claus plant 632 where it is joined with hydrogen sulfide, recovered from the amine regeneration 630 unit and sufficient H ₂ S, if available from "across the fence", or, if not, additional hydrogen from the PSA system 636 to reduce the sulfur dioxide to elemental sulfur.

All water used to produce steam (condensate from the steam turbines, cooling tower 634 return water and make-up water) are all processed in water purification unit 612 to make boiler-quality feed water. That portion of 500 ⁰ C, 1800 psi steam, raised from heat recovered from heat exchangers 610, 611 and 648, fluid bed roaster 626,

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Claus plant 632 and chillers 620 and 621, that is not directed toward making hydrogen, advances to steam turbine 638 to produce additional electricity by generator 640. The total generation of electricity is summarized in Tables 45 and 46.

Table 45 Electricity Generation - Combined Cycle Turbine

Table 46 Electricity Generation - Excess Steam Turbine

The total electricity generated by the combined cycle turbine and the excess steam turbine is 265 MW, or 2,089,260 MW-hr. The amount of electricity consumed by the process is summarized in Table 47.

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Table 47 Total Electricity Consumec i by Process

The total electricity consumed by the process is 81 MW, or 639,392 MW- hrs.

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Table 48 summarizes the total salable co-products from the process of this Example 2.

Table 48 Net Salable Co-products

Example 3

This Example 3 is directed to maximizing the amount of electricity produced. It derives from Example 1 (Figure 7). However, the high CO syngas stream from amine scrubber 623 and the high H ₂ syngas stream from amine scrubber 622 are combined and sent to turbine unit 642, air compressor and gas-fired turbine. Table 49 summarizes the net electricity generated by this Example 3.

Table 49 Net Salable Product

Net Electrical 922 MW

Energy for Sale 7,269,048 MW'hrs

Example 4

This example is directed at maximizing the amount of hydrogen that can be produced from a plant initially designed to produce 100 million scfd of hydrogen from the reduction of steam by the oxidation of liquid metallic iron. Example 4, like Example 3, derives from Example 1 (Figure 7). However, additional H ₂ is produced from the CO

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(in lieu of electricity) by employing the water gas shift reaction using steam to convert CO to H ₂ and co-mingled CO ₂ . Table 50 details the hydrogen production.

Table 50 Net Salable Product

Process versatility of the present invention is illustrated by the above four examples. The distribution of output energy for each of the four examples is summarized in Table 51.

Table 51 Process Energy Distribution

While various embodiments of the present invention have been described in detail, it is apparent that modifications and adaptations of those embodiments will occur to those skilled in the art. However, it is to be expressly understood that such modifications and adaptations are within the spirit and scope of the present invention.