PROCESS FOR PRODUCING HYDROGEN AND GRAPHITIC CARBON FROM HYDROCARBONS

Abstract

In accordance with the present invention, there is provided a process for producing hydrogen and graphitic carbon from a hydrocarbon gas comprising: contacting at a temperature between 600° C. and 1000° C. the catalyst with the hydrocarbon gas to catalytically convert at least a portion of the hydrocarbon gas to hydrogen and graphitic carbon, wherein the catalyst is a low grade iron oxide.

Claims

1. A process for producing hydrogen and graphitic carbon from a hydrocarbon gas comprising: contacting at a temperature between 600° C. and 1000° C. the catalyst with the hydrocarbon gas to catalytically convert at least a portion of the hydrocarbon gas to hydrogen and graphitic carbon, wherein the catalyst is a low grade iron oxide.

2. A process for producing hydrogen and graphitic carbon according to claim 1, wherein the pressure is greater than atmospheric pressure.

3. A process for producing hydrogen and graphitic carbon according to claim 1, wherein the pressure is 0 bar to 100 bar.

4. A process for producing hydrogen and graphitic carbon according to any one of the preceding claims, wherein the temperature is between 700° C. and 950° C.

5. A process for producing hydrogen and graphitic carbon according to any one of claims 1 to 4, wherein the temperature is between 800° C. and 900° C.

6. A process for producing hydrogen and graphitic carbon according to any one of claims 1 to 4, wherein the temperature is between 650° C. and 750° C.

7. A process for producing hydrogen and graphitic carbon according to any one of the preceding claims, wherein the hydrocarbon gas is methane.

8. A process for producing hydrogen and graphitic carbon according to any one of claims 1 to 9, wherein the hydrocarbon gas is natural gas.

9. A process for producing hydrogen and graphitic carbon according to any one of the preceding claims, wherein, the step of contacting the catalyst with the hydrocarbon gas is performed in a plurality of pressurised reactors arranged in series.

10. A process for producing hydrogen and graphitic carbon according to claim 9, wherein the arrangement of the reactors in series allows gas to flow from a first reactor to a subsequent reactor and each subsequent reactor in the series operates at a lower pressure than the preceding reactor, allowing gas to travel to reactors of lower pressure.

11. A process for producing hydrogen and graphitic carbon according to claim 10, wherein each reactor is provided with unreacted catalyst.

12. A process for producing hydrogen and graphitic carbon according to claim 9, wherein the arrangement of the reactors in series allows catalyst to flow from a first reactor to a subsequent reactor and each subsequent reactor in the series is operated at a higher pressure than the preceding reactor, allowing catalyst to travel to reactors of higher pressure.

13. A process for producing hydrogen and graphitic carbon according to claim 12, wherein unreacted hydrocarbon gas is provided to each reactor.

14. A process for producing hydrogen and graphitic carbon according to claim 9, wherein the arrangement of the reactors in series allows for both the hydrocarbon gas and catalyst to flow between reactors in opposite directions.

15. A process for producing hydrogen and graphitic carbon according to claim 14, wherein unreacted catalyst is provided in the lowest pressure reactor and unreacted hydrocarbon gas is provided in the highest pressure reactor and catalyst is transferred between the chambers of increasing pressure counter-currently to the gas flow between the chambers.

16. A method for the beneficiation of catalytic metal containing ore, the method comprising contacting at a temperature between 600° C. and 1000° C. the catalytic metal containing ore with a hydrocarbon gas to form a carbon-coated metal species.

17. A method for the beneficiation of catalytic metal containing ore according to claim 16, wherein the carbon-coated metal species is a graphite coated metal species.

18. A method for the beneficiation of catalytic metal containing ore according to claim 16 or claim 17, wherein the catalytic metal containing ore is iron ore.

19. A method for the beneficiation of catalytic metal containing ore according to any one of claims 16 to 18, wherein the pressure is greater than atmospheric pressure.

20. A method for the beneficiation of catalytic metal containing ore according to any one of claims 16 to 18. wherein the pressure is 0 bar to 100 bar.

21. A method for the beneficiation of catalytic metal containing ore according to any one of claims 16 to 19, wherein the temperature is between 700° C. and 950° C.

22. A method for the beneficiation of catalytic metal containing ore according to any one of claims 16 to 21, wherein the graphite is removed from the graphite coated metal species by contacting at temperature of 700° C. to 900° C. the graphite coated metal species with hydrogen gas.

23. A method for the beneficiation of catalytic metal containing ore according to any one of claims 16 to 22, wherein the removal of the graphite from the graphite coated metal species is performed in a pressurised reduction reactor at a pressure of 0 bar to 100 bar.

24. A method for the beneficiation of catalytic metal containing ore according to any one of claims 16 to 23, wherein the step of contacting at temperature of 700° C. to 900° C., the graphite coated metal species with hydrogen gas produces methane which is recycled to produce hydrogen.

25. A method for the beneficiation of catalytic metal containing ore according to claim 24, wherein the hydrogen produced in the step of recycling the methane is used in the step of contacting at temperature of 700° C. to 900° C. the graphite coated metal species with hydrogen gas to remove the graphite from the graphite coated metal species.

Description

BRIEF DESCRIPTION OF THE FIGURES

[0101] Further features of the present invention are more fully described in the following description of several non-limiting embodiments thereof. This description is included solely for the purposes of exemplifying the present invention. It should not be understood as a restriction on the broad summary, disclosure or description of the invention as set out above. The description will be made with reference to the accompanying drawings in which:

[0102] FIG. 1 shows a schematic representation of the process for producing hydrogen and graphite in accordance with the counter-current MPR of the present invention;

[0103] FIG. 2 shows a schematic representation of the process for producing hydrogen and graphite in accordance with the parallel gas MPR of the present invention;

[0104] FIG. 3 shows a schematic representation of the process for producing hydrogen and graphite in accordance with the parallel catalyst MPR of the present invention

[0105] FIG. 4 is a schematic representation of the process for the beneficiation of a catalytic metal containing ore in accordance with a first embodiment;

[0106] FIG. 5 is a graphical representation of XRD plots of the analytical grade iron oxides and iron ore catalyst samples;

[0107] FIG. 6 is a graphical representation of carbon purity (wt %) and carbon yield (grams of carbon per gram of iron—GC/GFe) of iron oxide catalysts post reaction;

[0108] FIG. 7 is a schematic representation of a three stage cascading counter-flow system;

[0109] FIG. 8 shows a schematic representation of the experimental conditions used to test the methane conversion of a MPR system with 3 reactors in series using static fixed bed reactors;

[0110] FIG. 9 is a graphical representation of the methane conversion results of hematite catalyst for different reaction pressures;

[0111] FIG. 10 is a graphical representation of carbon purity (wt %) and carbon yield (gram of carbon per gram of iron—GC/GFe) of hematite catalyst for different reaction pressures;

[0112] FIG. 11 is schematic shows a schematic representation of the variables for the mass balance calculation of a counter-current MPR system;

[0113] FIG. 12 is a graphical representation of the mass balance calculation results of both the counter-current MPR and parallel MPR showing the catalyst mass-flow required for a balanced system with a hydrogen production rate of 2000 m.sup.3/hr,

DETAILED DESCRIPTION OF THE PREFERRED EMBODIMENTS OF THE INVENTION

[0114] Those skilled in the art will appreciate that the invention described herein is susceptible to variations and modifications other than those specifically described. The invention includes all such variation and modifications. The invention also includes all of the steps, features, formulations and compounds referred to or indicated in the specification, individually or collectively and any and all combinations or any two or more of the steps or features.

[0115] Each document, reference, patent application or patent cited in this text is expressly incorporated herein in their entirety by reference, which means that it should be read and considered by the reader as part of this text. That the document, reference, patent application or patent cited in this text is not repeated in this text is merely for reasons of conciseness. None of the cited material or the information contained in that material should, however be understood to be common general knowledge.

[0116] Manufacturer's instructions, descriptions, product specifications, and product sheets for any products mentioned herein or in any document incorporated by reference herein, are hereby incorporated herein by reference, and may be employed in the practice of the invention.

[0117] The present invention is not to be limited in scope by any of the specific embodiments described herein. These embodiments are intended for the purpose of exemplification only. Functionally equivalent products, formulations and methods are clearly within the scope of the invention as described herein.

[0118] The invention described herein may include one or more range of values (e.g. size, concentration etc). A range of values will be understood to include all values within the range, including the values defining the range, and values adjacent to the range which lead to the same or substantially the same outcome as the values immediately adjacent to that value which defines the boundary to the range.

[0119] Throughout this specification, unless the context requires otherwise, the word “comprise” or variations such as “comprises” or “comprising”, will be understood to imply the inclusion of a stated integer or group of integers but not the exclusion of any other integer or group of integers.

[0120] Other definitions for selected terms used herein may be found within the detailed description of the invention and apply throughout. Unless otherwise defined, all other scientific and technical terms used herein have the same meaning as commonly understood to one of ordinary skill in the art to which the invention belongs.

[0121] Features of the invention will now be discussed with reference to the following non-limiting description and examples.

[0122] In a general form, the invention relates to a process for producing hydrogen and graphitic carbon from a hydrocarbon gas. In particular the present invention provides a process for catalytically converting hydrocarbon gas to hydrogen and graphitic carbon using a low grade iron oxide-containing catalyst.

[0123] The hydrocarbon gas may be any gas stream that comprises light hydrocarbons. Illustrative examples of hydrocarbon gas include, but are not limited to, natural gas, coal seam gas, landfill gas and biogas. The composition of the hydrocarbon gas may vary significantly but it will generally comprise one or more light hydrocarbons from a group comprising methane, ethane, ethylene, propane and butane.

[0124] In a preferred embodiment of the invention, the hydrocarbon gas is natural gas.

[0125] The process for producing hydrogen and graphitic carbon from natural gas comprises: [0126] contacting at a temperature between 600° C. and 1000° C. the catalyst with the hydrocarbon gas to catalytically convert at least a portion of the hydrocarbon gas to hydrogen and graphitic carbon,
wherein the catalyst is a low grade iron oxide.

[0127] Referring to FIG. 1, a counter-current MPR process 10, using three fluidised bed reactors, for producing hydrogen 12 and graphitic carbon 14 from a hydrocarbon gas, for example natural gas, 16 is described.

[0128] In the embodiment shown in FIG. 1, the process utilises three reactors operating at varying pressures, a high pressure reactor 18 at 18 bar, a medium pressure reactor 20 at 6 bar and a low pressure reactor 22 at 1 bar. The temperatures of the reactors are each 850° C. The reactors 18, 20 and 22 are arranged in series, such that hydrogen and unreacted hydrocarbon natural gas is transferred between adjacent reactors. i.e. from the high pressure reactor 18 to the medium pressure reactor 20 and from the medium pressure reactor 20 to the low pressure reactor 22.

[0129] Each reactor 22, 20 and 18 is respectively loaded with low grade iron oxide catalyst, for example iron ore 24″, 24′ and 24. Where 24 is fresh unreacted catalyst and 24′, 24″ are progressively more utilised, having more graphitic carbon attached and less overall catalytic activity remaining. Stream 14 contains only trace amounts of fully spent catalyst, with the vast majority (>90% wt at reaction temperatures of 850° C.) of this stream being graphitic carbon.

[0130] The amount of catalyst required for this reaction is relative to the quantity of hydrogen required, the process conditions and the type of catalyst. A 2000 m.sup.3/hr hydrogen production plant operating under the conditions above with 3 reactors would require approximately 14 kg/hr of iron

[0131] Natural gas 16 is directed through the reactors in series from the high pressure reactor 18, to the medium pressure reactor 20 and low pressure reactor 22. Each reactor converts a portion of the natural gas into hydrogen, with each successive gas stream 28, 30, 12 containing higher portions of hydrogen. Fresh natural gas 16 initially contacts the catalyst 24″ in the high pressure reactor 18 at a temperature of 850° . and a pressure of 18 bar to convert a portion of the natural gas into hydrogen, thus the corresponding gas stream 28 being a mixture comprising of hydrogen and unreacted natural gas. This reactor also deposits some graphitic carbon onto catalyst 24″ contributing to the total graphitic carbon in steam 14.

[0132] Gas stream 28 is transferred to the medium pressure reactor 20 where it contacts the catalyst 24′ at a temperature of 850° C. and a pressure of 6 bar to convert the natural gas into hydrogen and carbon. The lower pressure of the medium pressure reactor 20 enables conversion of the gas stream 28, thus contributing to the total hydrogen steam 12. The process deposits graphitic carbon onto the catalyst 24′ and in so contributes to the total graphitic carbon stream 14. A portion of the natural gas in gas stream 28 remains unreacted and mixes with the produced hydrogen gas to form gas stream 30.

[0133] Gas stream 30 is transferred to the low pressure reactor 22 where it contacts the catalyst 24 at a temperature of 850° C. and a pressure of 1 bar (atmospheric pressure). The lower pressure of the low pressure reactor 22 enables the thermodynamic equilibrium of the reaction to favour the decomposition direction of the reaction, thereby allowing more conversion of the second gas stream 30 into carbon and hydrogen gas . The process deposits graphitic carbon onto the catalyst 24, and in so contributes to the total graphitic carbon stream 14. This reactor also contributes to the hydrogen gas in the total hydrogen steam 12 which exits the reactor for use or further processing.

[0134] Theoretical empirical calculations dictate that the reactors 18, 20, 22 have conversion efficiencies of 54%, 75% and 94% respectively, and correspondingly the gas streams 28, 30 and 12 have hydrogen concentrations of 70%, 86% and 97% wt respectively.

[0135] The proportion of graphitic carbon in the iron oxide streams 24, 24′, 24″ and 14 are 0%, 91%, 95%, 98% respectively.

[0136] In the embodiment shown in FIG. 1, when the natural gas 16 contacts the catalyst 24 at a high temperature to produce hydrogen gas 12 and carbon 14, the catalyst 24 depletes to form a partially deactivated catalyst 24′. The partially deactivated catalyst 24′ is transferred between the reactors counter-currently to the natural gas 16 flow. The catalyst 24 is introduced into the lowest pressure reactor 22, and is subsequently passed to higher pressure reactors. The partially deactivated catalyst 24′ would therefore retain activity in the higher pressure reactor 20, and the resultant graphitic carbon 14 has the higher carbon-purity (as % of mass) with correspondingly higher value.

[0137] In FIG. 2, a parallel gas MPR process 50 is shown. Parallel gas MPR process 50 shares common features with counter-current MPR process 10 and like numerals denote like parts.

[0138] In the embodiment shown in FIG. 2, the process utilises three reactors operating at varying pressures, a high pressure reactor 18 at 18 bar, a medium pressure reactor 20 at 6 bar and a low pressure reactor 22 at 1 bar. The temperatures of the reactors are each 850° C. The reactors 18, 20 and 22 are arranged in series, such that unreacted hydrocarbon natural gas can be transferred between adjacent reactors. i.e. from the high pressure reactor 18 to the medium pressure reactor 20 and from the medium pressure reactor 20 to the low pressure reactor 22.

[0139] Each reactor 18, 20 and 22 is respectively loaded with an iron ore catalyst 52. In contrast to the counter-current MPR process 10 shown above, each reactor 22, 20 and 18 is provided with unreacted catalyst 52 prior to being contacted with the hydrocarbon gas.

[0140] The amount of catalyst required for this reaction is relative to the quantity of hydrogen required, the process conditions and the type of catalyst. A 2000 m.sup.3/hr hydrogen production plant operating under the conditions above with 3 reactors would require approximately 27 kg/hr of iron.

[0141] Natural gas 16 is directed through the reactors in series from the high pressure reactor 18, to the medium pressure reactor 20 and low pressure reactor 22. Each reactor converts a portion of the natural gas into hydrogen, with each successive gas stream 28, 30, 12 containing higher portions of hydrogen. Unreacted natural gas 16 initially contacts at a temperature of 850° C. and a pressure of 18 bar the catalyst 34 in the high pressure reactor 18 to convert a portion of the natural gas into hydrogen, producing a gas stream 28 which is a mixture of hydrogen and unreacted natural gas. Graphite is also deposited onto catalyst 34, producing partial graphite stream 54.

[0142] Gas stream 28 passes to the medium pressure reactor 20 where it contacts at a temperature of 850° C. and a pressure of 6 bar the catalyst 52 to convert the natural gas into hydrogen and carbon. The lower pressure of the medium pressure reactor 20 enables further conversion of the gas stream 28, thus contributing to the total hydrogen steam 12. The process deposits carbon onto the catalyst 52, producing partial graphite stream 56. A portion of the natural gas in gas stream 28 remains unreacted and mixes with the produced hydrogen gas to form gas stream 30.

[0143] Gas stream 30 passes to the low pressure reactor 22 where it contacts at a temperature of 850° C. and a pressure of 1 bar (atmospheric pressure) the catalyst 52. The lower pressure of the low pressure reactor 22 enables the thermodynamic equilibrium of the reaction to favour the decomposition direction of the reaction, thereby allowing conversion of the natural gas in second gas stream 30 into carbon and hydrogen gas. The process deposits carbon onto the catalyst 52, producing partial graphite stream 58. This reactor also contributes to the hydrogen gas in the total hydrogen steam 12, and exits the reactor for use or further processing.

[0144] Partial graphite streams 54, 56 and 58 contain a mixture of unreacted iron ore and graphitic material. Given the varying pressures of the each reactor 22, 20 and 18 each partial graphite stream will have different conversion rates. Partial graphite stream 58 will have the highest iron impurity, followed by partial graphite stream 56 and then partial graphite stream 54.

[0145] Empirically, the reactors 18, 20, 22 have conversion efficiencies of 54%, 75% and 94% respectively, and correspondingly the gas streams 28, 30 and 12 have hydrogen concentrations of 70%, 86% and 97% wt respectively.

[0146] In FIG. 3, a parallel catalyst MPR process 60 is shown. Parallel MPR process 60 shares common features with counter-current MPR process 10 and like numerals denote like parts.

[0147] In the embodiment shown in FIG. 3, the process utilises three reactors operating at varying pressures, a high pressure reactor 18 at 18 bar, a medium pressure reactor 20 at 6 bar and a low pressure reactor 22 at 1 bar. The temperatures of the reactors are each 850° C.

[0148] Each reactor 18, 20 and 22 is respectively loaded with a low grade iron oxide containing catalyst, for example iron ore 24″, 24′ and 24. Where 24 is unreacted catalyst and 24′, 24″ are progressively more utilised, having more carbon attached and less overall catalytic activity remaining. Stream 14 contains only trace amounts of fully spent catalyst, with the vast majority (>90% wt at reaction temperatures of 850° C.) of this stream being graphite.

[0149] The reactors 18, 20 and 22 are arranged in series, such that catalyst 24″, 24′ and 24 is transferred between adjacent reactors. i.e. from the low pressure reactor 22 to the medium pressure reactor 20 and from the medium pressure reactor 20 to the high pressure reactor 18.

[0150] In contrast to the counter-current MPR process 10 shown above, unreacted natural gas 16 is provided to each reactor 22, 20 and 18.

[0151] In the embodiment shown in FIG. 8, when the natural gas 16 contacts the catalyst 24 at a high temperature to produce hydrogen gas 12 and carbon 14, the catalyst 24 depletes to form a partially deactivated catalyst 24′. The partially deactivated catalyst 24′, 24″ is transferred between the reactors. The catalyst 24 is introduced into the lowest pressure reactor 22, and is subsequently passed to higher pressure reactors 20, 18. The partially deactivated catalyst 24′ would therefore retain activity in the higher pressure reactor 20, and the resultant carbon 14 has the higher carbon-purity (as % of mass) with correspondingly higher value.

[0152] Natural gas 16 contacts the catalyst 24 in the low pressure reactor 22 at a temperature of 850° C. and a pressure of 18 bar to convert a portion of the natural gas 16 into hydrogen to produce a gas stream 68, being a mixture comprising of hydrogen and unreacted natural gas. Graphitic carbon is deposited onto catalyst 24 to produce catalyst 24′ contributing to the total carbon in stream 14.

[0153] Natural gas 16 contacts the catalyst 24′ in the medium pressure reactor 20 at a temperature of 850° C. and a pressure of 6 bar to of the natural gas 16 into hydrogen to produce a gas stream 64, being a mixture comprising of hydrogen and unreacted natural gas. Graphitic carbon is deposited onto catalyst 24′ to produce catalyst 24″ contributing to the total carbon in stream 14.

[0154] Natural gas 16 contacts the catalyst 24″ in the high pressure reactor 18 at a temperature of 850° C. and a pressure of 18 bar to convert a portion of the natural gas 16 hydrogen to produce a gas stream 62, being a mixture comprising of hydrogen and unreacted natural gas. Graphitic carbon is deposited onto catalyst 2″ to produce catalyst graphitic carbon 14.

[0155] Referring to FIG. 4, a process 100 for the beneficiation of a catalytic metal containing ore for example iron ore 102 is described.

[0156] Low grade iron ore 102 is passes through a surge bin 104 and into a dusting reactor 106. In the dusting reactor 106, the iron ore 102 is contacted at a temperature between 850° C. and a pressure between 10 and 20 bar with a hydrocarbon gas 108 to produce a graphite coated iron stream 110 and a waste stream 112 comprising of larger (>1 mm) gangue particles. The size difference between the graphite coated iron stream 110 and a waste stream 112 separates the streams. The graphite coated iron stream 110 is passed through a gas/solids separator 114 to separate the gas stream 116 from the solids stream 118, which is passed to a reduction reactor 120.

[0157] In the reduction reactor 120, the graphite coated iron particles of the solids stream 118 contact with hydrogen gas 122 at a temperature between 800° C. and 900° C. and a pressure between 10 and 20 bar in order to remove the carbon coating, leaving a iron concentrate stream 124. The reaction also forms a methane gas stream 126, which is recycled into other parts of the process. In the embodiment shown in FIG. 2, the methane gas stream 126 is contacted at a temperature between 800° C. and 900° C. with further iron ore 102 passed through a surge bin 130 in a hydrogen reactor 127 to produce hydrogen gas 122 and a graphite powder 128. As shown in FIG. 4, the hydrogen gas 122 is transferred back into the reduction reactor 120.

EXAMPLES

Example 1

[0158] The use of iron ore as the catalyst for the economical production of hydrogen and graphite via the thereto-catalytic decomposition of methane.

Experimental Details

[0159] The present invention provides a method which enables the use of low grade iron oxide as a catalyst for the decomposition of methane. In order to demonstrate the catalytic activity of the low grade iron oxide catalyst of the present invention, samples of low grade iron oxide were compared to high grade iron oxide samples. Two types of high grade iron oxide were tested; hematite (99%, <5 μm, Sigma-Aldrich) and magnetite (95%, <5 μm, Sigma-Aldrich); as well as two iron ore samples: Hematite ore (Pilbara mine) and goethite ore (Yandi mine). The ore samples were milled to <150 μm but otherwise untreated. The ‘as received’ compositional data, particle size distribution, and surface area of all the samples are detailed in Table 1.

TABLE-US-00001 TABLE 1 Compositional, particle size and surface area data for the iron oxide samples. Ca Mn Mg Ti Fe P Si Al Oxide Oxide Oxide Dioxide Iron Oxide Type % % % % % % % % Analytical Grade Fe.sub.2O.sub.3 69.9 — — — — — — — Analytical Grade Fe.sub.3O.sub.4 72.3 — — — — — — — Hematite ore (Pilbrara) 62.9 0.1 4.0 2.2 0.0 0.1 0.0 0.1 Goethite ore (Yandi mine) 57.9 0.0 5.5 1.4 0.0 0.0 0.0 0.1 K Balance Surface S Oxide Cl Oxide D.sub.10 D.sub.50 D.sub.90 area Iron Oxide Type % % % % um um um m.sup.2/g Analytical Grade Fe.sub.2O.sub.3 — — — 30.1 0.2 0.4 1.65 5.37 Analytical Grade Fe.sub.3O.sub.4 — — — 27.7 0.17 1.42 3.24 6.24 Hematite ore (Pilbrara) 0.0 0.0 0.0 30.5 3.2 44.5 141.0 7.62 Goethite ore (Yandi mine) 0.0 0.0 0.0 34.9 3.1 47.8 156.8 29.76

[0160] Each sample was placed in a separate single stage reactor. The reactors were vertical 1/2″ diameter stainless steel (SS316 Swagelok) tube, with ⅜″ quartz tube internal liners. The quartz tube internal liners reduce the catalytic effect of the stainless steel reactor walls by restricting contact with the reacting methane gas. 20 g catalyst samples were contained within a ⅜″ ‘test-tube like’ quartz chamber.

[0161] The XRD plots of the high grade iron oxide catalyst samples, namely analytical grade (hematite and magnetite) and low grade iron oxide catalyst samples (hematite ore and goethite ore) are shown in FIG. 5.

[0162] Each sample was reacted at temperatures ranging from 750-950° C., using 10 sccm pure methane (UHP), and a reaction pressure between 1-9 bar (absolute). After complete deactivation (approximately 19 hr) the reaction was terminated and the samples were cooled with 20 sccm of pure nitrogen (UHP). The resulting carbon (and embedded catalyst particles) was weighed to determine the total carbon yield per gram of iron catalyst used.

[0163] FIG. 6 shows the results of these experiments at reaction conditions of 850° C. and atmospheric pressure. The results show that the low grade iron ore samples performed almost as well as the high grade oxides, with carbon yields ranging from 9.2 to 8.9 grams of carbon per gram of iron, corresponding to carbon purities of 90 wt % to 89 wt % respectively. These values are shown to closely correlate with the quantitative XRD derived values, with differences of less than 2 wt % (represented as hollow shapes in FIG. 6).

[0164] As would be understood by a person skilled in the art, a common way to increase the activity of a catalyst is to make it very high purity in order to increase the reactive area. Iron oxide catalysts, such as the high grade iron oxide sample tested must be must be specifically synthesized to have a purity of >99%. The results of this experiment indicate that the particular process condition of the present invention allow for the use of a low grade catalyst, whilst still obtaining high conversion rates and yield.

Example 2

[0165] Thermo-Catalytic methane Decomposition using Counter-Current MPR.

Counter-Current.

[0166] A three reactor counter-current MPR was set up in a cascade arrangement as shown in the schematic of FIG. 7.

[0167] Experimental evaluation of the counter-current MPR system was undertaken using a static (non-continuous) system. This was done by testing the effect of pressure on the methane conversion efficiency and the carbon yield. The results confirmed that an increase in pressure lowered the methane conversion, and increased the carbon yield, and conversely a lower pressure increased the methane conversion and lowered the total carbon yield.

Experimental Details

[0168] Effect of Reaction Pressure on the methane Conversion Limit.

[0169] The reactor set-up comprised three independent reactor stages (3×½″ OD 316SS Swagelok, 700 mm length) with different set back-pressures (12 bar, 4 bar and atmospheric) and an isothermal temperature of 850° C. Instead of linking the reactors in series, each was fed and analysed independently in order to assess their individual performances. The feed gas compositions of each reactor were set to simulate their operation in series, where each of the reactors were operating at their theoretical maximum possible conversion at the reaction pressure (Table 2). The performance of each stage was determined by monitoring the effluent from each reactor using a Gas Chromatograph (GC). A schematic of this process is shown in FIG. 8. An excess amount of iron oxide was used to simulate continuous catalyst flow conditions within this static system to allow a momentary glimpse of the expected steady-state continuous operation.

TABLE-US-00002 TABLE 2 Process conditions for the MPR experimental trial Reactor Stage R1 R2 R3 Pressure (bar) 1 4 12 Temperature (° C.) 850 850 850 Theoretical TEL (%) 91.9 72.3 54.4 Methane input flow (sccm) 2.8 4.8 10 Hydrogen input flow (sccm) 14.4 10.4 0

[0170] The results obtained from this experiment are shown in FIG. 9 and are in good correlation with theoretical expectations and validate the theory. All three reactor stages correlated well with the expected thermodynamic equilibrium limit (shown as dashed lines) for a period exceeding 20 hr, after which the reaction was terminated. It is clear that the high hydrogen concentration did not affect the ability for the reaction to attain conversions at the thermodynamic equilibrium limit when the MPR system was used.

[0171] These results indicate that a continuous MPR system can sustain stable conversions at the thermodynamic equilibrium limit regardless of the level of hydrogen.

Effect of Pressure on Product Yield

[0172] The effects of reaction pressure were tested using 20 mg of catalyst at pressure intervals of 1 bar, with all other reaction conditions remaining the same as the previous experiments (namely 850° C., 20 sccm methane, auto-reduction, 19 hr duration).

[0173] The results indicate that there is a positive linear relationship between the reaction pressure and the total carbon yield. The profile, as shown in FIG. 10 shows that the carbon yield per gram of iron increases from −9 g to 22 g across the pressure range of atmospheric pressure to 9 bar absolute, corresponding to carbon purities of −90-96% respectively.

Empirical Catalyst Flow Rate Calculations

[0174] The overall feasibility of the MPR systems are dependent on the balance of the mass flows. This is of particular importance for the counter flow MPR because of the strict interdependence of the catalyst mass flow with (1) the number of reactor pressure stages, (2) the range of pressures, and (3) the catalyst carbon capacity profile. An empirical mass balance calculation was done to determine the feasibility of attaining balance.

[0175] The catalyst flow rate within each reactor can be determined by dividing the carbon deposition rate by the catalyst utility for each reactor; which are both bounded by the reactor pressure range.

[00001]${\stackrel{.}{M}}_{\mathrm{cat}\left(\mathrm{Rn}\right)}=\frac{\stackrel{.}{M_{c\left(\mathrm{Rn}\right)}}}{\Delta C_{\left(\mathrm{Rn}\right)}}$

Where

[0176] {dot over (M)}.sub.catt(Rn) is the mass flow rate of catalyst, M.sub.c(Rn) is the carbon deposition rate, and [0177] ΔC.sub.(Rn) is the catalyst utility through reactor ‘Rn’.

Parallel MPR

[0178] Parallel flow MPRs have the advantage of design simplicity and fewer constraints determining the number of stages and pressure limits.

Counter-Current MPR

[0179] The biggest constraint for the counter-current MPR arrangement is balancing the catalyst mass flow between all reactor stages in order to enable continuous operation. The catalyst flow rate required at each stage is dependent on (1) the number of reactor stages and (2) the catalyst carbon capacity profile relative to pressure. This balance is illustrated in FIG. 11.

[0180] The purpose of this calculation is to determine the number of pressure stages that balances the catalyst flow rate between all stages, for a given catalyst carbon capacity profile and reaction temperature (assuming isothermal conditions).

[0181] If the catalyst mass flow rate is set so that it is fully deactivated when exiting each reactor stage, the catalyst utility at each stage is the difference between the total catalyst utility at the reactor pressure and the adjacent lower pressure reactor. Thus:

ΔC.sub.Rn=C.sub.n−C.sub.n−1

[0182] Where ‘n’ is the reactor number (n=1 is the lowest pressure reactor).

[0183] The catalyst mass flow rate through each reactor stage then becomes:

[00002]${\stackrel{.}{M}}_{\mathrm{cat}\left(\mathrm{Rn}\right)}=\frac{{\left(\frac{P}{\mathrm{RT}}\right)}_{\mathrm{STP}}.Math.{\stackrel{.}{Q}}_{{\mathrm{CH}}_4\left(I\right)}.Math.\left[{\xi }_{\mathrm{Rn}}-{\xi }_{\mathrm{Rn}+1}\right].Math.12}{\left[\alpha .Math.\frac{K_T}{4}.Math.\left(\frac{1}{{\xi }_{\mathrm{Rn}}^2}-1\right)+\beta \right]-\left[\alpha .Math.\frac{K_T}{4}.Math.\left(\frac{1}{{\xi }_{\mathrm{Rn}}^2}-1\right)+\beta \right]}{\xi }_{\mathrm{Rn}}$

[0184] Where P, R, and T are the STP pressure, gas constant and temperature respectively, {dot over (Q)}.sub.CH.sub.4.sub.(I) is the initial methane feed rate, ξ.sub.Rn is the TEL at reactor ‘n’, α and β are coefficients relating to the pressure effects on the carbon capacity, KT is the equilibrium constant at temperature T.

[0185] For a reactor system with only one reactor (n=1), reactor (n−1) and (n+1) reactor stages do not exist; thus this can be simplified to:

[00003]${\stackrel{.}{M}}_{\mathrm{cat}\left(R1\right)}=\frac{{\left(\frac{P}{\mathrm{RT}}\right)}_{\mathrm{STP}}.Math.{\stackrel{.}{Q}}_{{\mathrm{CH}}_4\left(I\right)}.Math.{\xi }_{R1}.Math.12}{\alpha .Math.\frac{K_T}{4}.Math.\left(\frac{1}{{\xi }_{R1}^2}-1\right)+\beta }$

[0186] Similarly for a two stage reactor, reactor (n−1) and (n+2) reactors do not exist, and for a three stage reactor (n−1) and (n+3) reactors do not exist.

Catalyst Mass Flow Balance

[0187] For a multistage process to be continuous the catalyst flow rate must equate:

{dot over (M)}.sub.catt(R1)={dot over (M)}.sub.catt(R2)

[0188] Thus for a two stage reactor:

[00004]$\frac{{\left(\frac{P}{\mathrm{RT}}\right)}_{\mathrm{STP}}.Math.{\stackrel{.}{Q}}_{{\mathrm{CH}}_4\left(I\right)}.Math.\left[{\xi }_{\mathrm{Rn}}-{\xi }_{\mathrm{Rn}+1}\right].Math.12}{\alpha .Math.\frac{K_T}{4}.Math.\left(\frac{1}{{\xi }_{R1}^2}-1\right)+\beta }=\frac{{\left(\frac{P}{\mathrm{RT}}\right)}_{\mathrm{STP}}.Math.{\stackrel{.}{Q}}_{{\mathrm{CH}}_4\left(I\right)}.Math.{\xi }_{R2}.Math.12}{\left[\alpha .Math.\frac{K_T}{4}.Math.\left(\frac{1}{{\xi }_{R2}^2}-1\right)+\beta \right]-\left[\alpha .Math.\frac{K_T}{4}.Math.\left(\frac{1}{{\xi }_{R1}^2}-1\right)+\beta \right]}$

[0189] For isothermal conditions this can be simplified to:

[00005]$\frac{{\xi }_{R1}-{\xi }_{R2}}{\left(\frac{1}{{\xi }_{R1}^2}-1\right)+\frac{\beta .Math.4}{K_T.Math.\alpha }}=\frac{{\xi }_{R2}}{\frac{1}{{\xi }_{R2}^2}-\frac{1}{{\xi }_{R1}^2}}$

[0190] If reactor ‘R1’ is operating at atmospheric pressure and operating at the TEL then ξ.sub.R1 is known, and ξ.sub.R2 can be solved using the above equation. This solution would solve for the pressure required for reactor 2 to equate the catalyst flow rate.

[0191] This can similarly be done for larger numbers of reactors:

Three Reactor Stages

[0192] [00006]$\frac{{\xi }_{R1}-{\xi }_{R2}}{\left(\frac{1}{{\xi }_{R1}^2}-1\right)+\frac{\beta .Math.4}{K_T.Math.\alpha }}=\frac{{\xi }_{R2}-{\xi }_{R3}}{\frac{1}{{\xi }_{R2}^2}-\frac{1}{{\xi }_{R1}^2}}=\frac{{\xi }_{R3}}{\frac{1}{{\xi }_{R3}^2}-\frac{1}{{\xi }_{R2}^2}}$

Four Reactor Stages

[0193] [00007]$\frac{{\xi }_{R1}-{\xi }_{R2}}{\left(\frac{1}{{\xi }_{R1}^2}-1\right)+\frac{\beta .Math.4}{K_T.Math.\alpha }}=\frac{{\xi }_{R2}-{\xi }_{R3}}{\frac{1}{{\xi }_{R2}^2}-\frac{1}{{\xi }_{R1}^2}}=\frac{{\xi }_{R23}-{\xi }_{R4}}{\frac{1}{{\xi }_{R3}^2}-\frac{1}{{\xi }_{R2}^2}}=\frac{{\xi }_{R4}}{\frac{1}{{\xi }_{R4}^2}-\frac{1}{{\xi }_{R3}^2}}$

[0194] The above can then be extrapolated to additional reactor stages.

Results

[0195] The empirical results using the linear extrapolated values for pressures above 9 bar(abs) are graphically demonstrated in FIG. 12. These results show that the counter flow MPR would consume considerably less catalyst than the parallel flow process under all scenarios. The counter-flow process with 5 reaction stages requires only 19% of the catalyst required for a single reactor, whereas the parallel-flow process requires 42% with the same number of stages. The counter-flow process however is only able to have a maximum of 5 reactor stages for the catalyst mass flow to be constant for all stages. In contrast, the number of parallel-flow process stages is limitless; however it can be seen that each stage has diminishing returns and overall require significantly more catalyst than the counter-flow option. The catalyst mass flow rates are calculated based on an assumed hydrogen output flow rate of 2000 m.sup.3/hr.

Example 3

[0196] Beneficiation of Iron Ore.

Experimental Details

[0197] Typical low grade iron ore rock consists of distinct sections of high grade iron oxide and low grade counterpart. This type of rock is known as banded iron formation (BIF). A 6.39 g sample of BIF iron ore was prepared, an analysis of the characteristics are shown in Table 3.

TABLE-US-00003 TABLE 3 Sample Analysis OXIDE SiO.sub.2 TiO.sub.2 Al.sub.2O.sub.3 Fe.sub.2O.sub.3 Mn.sub.3O.sub.4 MgO CaO Na.sub.2O K.sub.2O P.sub.2O.sub.5 Iron ore rich section 10.62 0.06 0.11 88.8 0.01 0.12 0.02 0.23 0.01 0.03 Iron ore poor section 84.1 0.04 0.14 12.9 <0.01 0.07 0.21 0.11 0.03 0.02 OXIDE SO.sub.3 Cr.sub.2O.sub.3 ZrO.sub.2 SrO ZnO CuO NiO BaO PbO L.O.I. TOTAL Iron ore rich section <0.01 0.13 0.02 0.01 <0.01 0.01 0.27 <0.01 <0.01 1.98 102.43 Iron ore poor section <0.01 0.21 <0.01 <0.01 <0.01 <0.01 0.43 <0.01 <0.01 1 99.26

[0198] The sample was loaded into a static reactor bed and was contacted at 900° C. with methane gas and atmospheric pressure for a period of 4 hours. Following reaction, the high grade iron oxide band had fragmented whereas the low grade counterpart was largely unaffected.

[0199] Without wishing to be bound by theory it is understood by the inventors that the first reaction that occurs is the reduction of the aggregate iron oxide species ore into iron carbide, emitting water vapour, H.sub.2, CO.sub.2 and trace CO. Continued reaction causes the aggregate iron carbide to fragment via metal dusting (as described earlier) and in the absence of oxides the system emits H.sub.2 gas only. This dusting causes all the iron species to disintegrate into micron and nano fragments due to the encapsulating graphitic layers. The gangue of the iron ore (typically highly stable minerals containing SiO.sub.2 and Al.sub.2O.sub.3) are unaffected by these process conditions and remain intact and unaltered. The product of the process is therefore left with larger aggregates of gangue and tiny particles of graphite encapsulated ferric iron/iron carbide. The size and density difference between the iron species and the gangue can then be exploited to separate the two in cyclonic separator.

[0200] The compositional data of the samples after the reaction and physical separation by size is shown in Table 4.

TABLE-US-00004 TABLE 4 Sample Analysis OXIDE WT % SiO.sub.2 TiO.sub.2 Al.sub.2O.sub.3 Fe.sub.2O.sub.3 Mn.sub.3O.sub.4 MgO CaO Na.sub.2O Sample A 11 0.05 0.12 89.05 <0.01 0.13 0.02 0.21 Sample B 79.13 <0.01 0.14 19.72 <0.01 0.11 0.06 0.08 OXIDE WT % K.sub.2O P.sub.2O.sub.5 SO.sub.3 Cr.sub.2O.sub.3 ZrO.sub.2 SrO Sample A 0.02 0.04 <0.01 0.05 0.02 0.01 Sample B 0.05 0.03 <0.01 <0.01 <0.01 <0.01 OXIDE WT % ZnO NiO BaO PbO CuO Total Sample A <0.03 0.08 <0.01 <0.01 0.07 100.87 Sample B <0.01 <0.01 <0.01 <0.01 <0.01 99.29

[0201] The analysis showed that the size separation was able to separate the majority of the iron species, with sample A corresponding to the majority of iron. The compositional data was determined by XRF analysis, which requires the sample to be oxidised beforehand, thus showing all iron species as oxides instead of ferrite. Energy-dispersive X-ray spectroscopy analysis prior to calcination showed the iron species to be ferrite. Empirically, removing this oxide from the iron composition we are able to calculate that the process is able to extract a product that is 85 wt % iron from an original total rock composition of approximately 35 wt %

[0202] It is envisaged that the graphitic carbon can then be removed from the graphite encapsulated ferric iron/iron carbide by a process called methanation. In this reaction the iron/carbon particles are contacted with hydrogen gas at elevated temperatures, to form methane gas by way of the following reaction 2.

C+2H.sub.2.fwdarw.CH.sub.4   (2)

[0203] As the iron particles are very small and this reaction is exothermic the iron particles agglomerate to form larger particles of pure iron.

[0204] It is envisaged that the graphitic carbon can then be removed from the graphite encapsulated ferric iron/iron carbide by contacting at 800° C. and 20 bar the graphite encapsulated ferric iron/iron carbide with hydrogen gas.

[0205] The advantages of the beneficiation method of the present invention over classical methods of iron ore beneficiation is that the produced iron oxide species are reduced (oxygen removed leaving ferric iron) in addition to the gangue being removed. This reduced iron is 90-95% wt iron whereas high grade iron ore is typically 55-63% wt (70% theoretical maximum). Reduced iron is a premium product compared to iron ore and thus commands a higher price. Also, the reduced iron product potentially has lower transportation costs because ballast oxygen is not transported—a saving of 30-40% by weight and −50% by volume. Classical beneficiation processes used in industry for iron ore include milling, magnetic separation, floatation, gravity concentration, thickening/filtering and agglomeration.

[0206] Those skilled in the art will appreciate that the invention described herein is susceptible to variations and modifications other than those specifically described. The invention includes all such variation and modifications. The invention also includes all of the steps, features, formulations and compounds referred to or indicated in the specification, individually or collectively and any and all.