There has been a recent shift towards using fuels produced from renewable resources in today's environmentally conscious world. Biomass found in municipal solid waste (MSW) provides an excellent opportunity as a major, near-term, carbon-neutral energy resource. MSW naturally biodegrades, producing landfill gas (LFG) composed mainly of methane and carbon dioxide, two major greenhouse gases, which can be used to generate the fuels.

Despite the potential benefits of processing MSW to generate liquid fuels, less than 15% of the over 243 million tons of MSW produced each year is used for that purpose. One reason for this is that landfills currently lack robust technologies that can efficiently convert environmentally harmful hydrocarbons produced in LFG into liquid fuels. Existing technologies are inhibited by high capital costs and low economic recovery and therefore require carbon capture credits to be economically feasible. Current technologies also require specific deliverables in order to function as designed. If feedstock flows are outside the required specifications, the LFG is flared and the resource is effectively wasted.

New LFG-to-liquids processes could provide high economic returns from an abundant and renewable feedstock. At the current prices of diesel and jet fuel, the end product would be an attractive alternative to power generation. Once a landfill is outfitted with an LFG-to-liquids plant, the fuel product could also be used to decrease fuel requirements needed to perform ordinary landfill tasks. Additionally, the fuel product could further be marketed to interested parties because it is compatible with existing infrastructure.

In view of the above discussion, it can be appreciated that it would be desirable to have alternative systems and methods for producing liquid fuels from MSW and/or LFG. Summary of the Invention

Viewed from a first aspect, the present invention provides a method for producing liquid fuel from landfill gas, the method comprising:

(i) providing 0 ₂ , steam and landfill gas comprising 30 to 70 %vol methane and 35 to 55 %vol C0 ₂ to a single tri-reformer reactor comprising a first catalyst; (ii) supplying energy to said single tri-reformer reactor;

(iii) performing a tri-reforming process on said landfill gas, wherein said tri- reforming process comprises carbon dioxide reforming, steam reforming, water-gas shifting, and methane oxidation, to produce synthesis gas having a H ₂ :CO ratio of approximately 2:1 ;

(iv) providing the synthesis gas to a Fischer-Tropsch synthesis (FTS) reformer comprising a second catalyst;

(v) converting said synthesis gas into liquid fuel, fuel gas and steam within said Fischer-Tropsch synthesis reformer;

(vi) separating said liquid fuel from said fuel gas and said steam; and

(vii) combusting at least some of said fuel gas to provide at least some of said energy to said tri-reformer reactor, whereby said method for producing liquid fuel is at least partially self-sufficient in terms of energy for said tri-reforming process.

Viewed from a further aspect, the present invention provides a method for producing liquid fuel from municipal solid waste comprising:

(i) biodegrading municipal solid waste in a landfill site to produce crude landfill gas;

(ii) separating particulate matter from said crude landfill gas to produce landfill gas; and

(iii) converting said landfill gas to liquid fuel as hereinbefore described.

Viewed from a further aspect, the present invention provides a system for producing liquid fuel from landfill gas, the system comprising:

an oxygen supply line;

a steam supply line;

a landfill gas supply line;

a fuel gas combustion unit that combusts fuel gas and provides heat to a single tri-reformer reactor;

a single tri-reformer reactor for performing a tri-reforming process on the landfill gas comprising carbon dioxide reforming, steam reforming, water-gas shifting, and methane oxidation, to produce synthesis gas having a H ₂ :CO ratio of approximately 2. , wherein said tri-reformer has one or more inlets fluidly connected to said oxygen supply line, said steam supply line and said landfill gas supply line and at least one outlet for said synthesis gas; a Fischer-Tropsch synthesis (FTS) reformer for converting the synthesis gas to liquid fuel, fuel gas and steam, wherein said FTS reformer has one or more inlets for said synthesis gas fluidly connected to said tri-reformer and an outlet for the liquid fuel, fuel gas and steam;

a separation unit for separating said liquid fuel, said fuel gas and said steam, wherein said separation unit has an inlet for the liquid fuel, fuel gas and steam fluidly connected to the Fischer-Tropsch synthesis reformer and separate outlets for each of liquid fuel, fuel gas and steam; and

a line connecting said fuel gas outlet of said separation unit to said fuel gas combustion unit.

Preferably the system is located on a landfill site.

Description of the Invention

The present invention relates to a method and system 'for producing liquid fuel from landfill gas. Landfill gas generally comprises 30 to 70 %mol methane and 35 to 55 %mol C0 ₂ . Preferably the landfill gas used in the method and system of the present invention comprises 40 to 65 %mol methane and 30 to 50 %mol C0 ₂ and still more preferably 45 to 60 %mol methane and 35 to 50 %mol C0 ₂ .

Preferably the landfill gas used in the method and system of the present invention further comprises at least one of:

0.5-20 %mol nitrogen;

0.1-2.5 %mol oxygen;

1-1700 ppm hydrogen sulfide;

1-400 ppm halides;

1-10 %mol water; and

200-15,000 ppm non-methane organic compounds.

More preferably the landfill gas additionally comprises at least two, still more preferably at least three and yet more preferably all of the above-mentioned components. Preferably the landfill gas used in the method and system of the present invention is produced from municipal solid waste by biodegradation.

In the method and system of the present invention the landfill gas undergoes tri- reforming in a single tri-reformer reactor to produce synthesis gas having a H ₂ :CO ratio of approximately 2:1 mole ratio. This ratio is the optimum ratio for subsequent Fischer- Tropsch synthesis of liquid fuel, i.e. useful hydrocarbons. The tri-reforming reaction requires three oxidants, specifically 0 ₂ , H ₂ 0 and C0 ₂ . Preferably 0 ₂ and H ₂ 0 are separately provided, e.g. injected, into the tri-reforming process. Preferably 0 ₂ is provided in the form of air.

Preferably the amount of 0 ₂ provided to the tri-reformer reactor is 1 -25 %mol, more preferably 3-15 %mol and still more preferably 5-10 %mol, based on the total moles of gases in the reactor. Preferably the H ₂ 0 is provided to the tri-reformer reactor in the form of steam. Preferably the amount of steam provided to the tri-reformer reactor is 2-45 %mol, more preferably 7-30 %mol and still more preferably 15-30 %mol, based on the total moles of gases in the reactor. Preferably the amount of C0 ₂ present in the landfill gas is sufficient for the tri-reforming process and no further addition of C0 ₂ is required.

Preferably the molar ratio of 0 ₂ to steam is 1 :15 to 1 :0.2, more preferably 1 : 10 to 1 :0.46 and still more preferably :5 to 1 : 1. Preferably the molar ratio of 0 ₂ to C0 ₂ is 1 : 10 to 1 :0.66, more preferably 1 :7 to 1 :1 and still more preferably 1 :5 to 1 :1.5. Preferably the molar ratio of methane to 0 ₂ is 1 :0.1 to 1 :0.5, more preferably 1 :0.15 to 1 :0.4 and still more preferably 1 :0.2 to 1 :0.33. Advantageously the amount of 0 ₂ and/or H ₂ 0 provided to the tri-reforming process may be altered to control the conversion achieved therein.

Preferably the molar ratio of methane:C0 ₂ :H ₂ 0:0 ₂ is 1 :0.33:0.1 :0.1 to 1 :1 :1.5:0.5 and more preferably 1 :1.33:0.23:0.2 to 1 :1 :1 :0.33.

A major advantage of the method and system of the present invention is that it is at least partially self-sufficient in terms of energy for the tri-reforming process. One of the major hurdles in reforming processes is the large amounts of energy required to drive the endothermic reactions that create the syngas. In preferred methods and systems of the present invention at least some of the fuel gas produced in the Fischer- Tropsch reformer is combusted to provide energy for the single tri-reformer reactor. Preferably at least 50 %vol, more preferably at least 75 %vol and still more preferably at least 85 %vol of the fuel gas produced in the Fischer-Tropsch reformer is combusted to provide energy for the single tri-reformer reactor. This may be achieved by transporting the fuel gas produced in the Fischer-Tropsch reformer to a fuel gas combustion unit, e.g. a furnace unit. Such units are conventional in the art.

In some preferred methods and systems of the invention a portion of landfill gas is combusted to provide energy to the single tri-reformer reactor. Thus preferred methods of the invention comprise the further step of combusting landfill gas to provide energy to the single tri-reformer reactor. The combustion may be carried out in a landfill gas combustion unit. Such units are conventional in the art. In preferred methods and systems of the invention the combustion of landfill gas is only required to provide energy to the single tri-reformer reactor upon initial start up until FTS product fuel gas is produced.

In preferred methods and systems of the invention the majority of the energy provided to the single tri-reformer reactor is from combustion of the fuel gas. Preferably 50-100 %, more preferably 75-100 % and still more preferably 85-100 % of the energy provided to the single tri-reformer reactor is from combustion of the fuel gas. Preferably 0-100 %, more preferably 0-75 % and still more preferably 0-50 % of the energy provided to the single tri-reformer reactor is from combustion of landfill gas. Preferably none of the energy provided to the single tri-reformer unit is from combustion of natural gas, coal, or petroleum derived fuel.

The recycling of the fuel gas produced in the Fischer-Tropsch reformer to provide energy is particularly beneficial because the conditions used in the Fischer- Tropsch synthesis are designed such that relatively few heavy hydrocarbons such as waxes are produced therein. This means that the energy content of the fuel gas is higher than in a conventional Fischer-Tropsch synthesis and thus that combustion of the fuel gas can provide significant amounts of energy to the single tri-reforming process.

In preferred methods and systems of the invention, the fuel gas comprises C _1-5 hydrocarbons, H ₂ , and CO. Particularly preferably the fuel gas comprises methane, H ₂ and CO.

In preferred methods of the invention, the method is substantially self-sufficient in terms of energy for the tri-reforming process. Thus no energy is derived from combustion of natural gas, coal, or other petroleum derived fuels. Preferably substantially all of the energy for the tri-reforming process is from combustion of the fuel gas and/or combustion of landfill gas. Still more preferably 50-100 % and even more preferably 85-100 % of the energy is from combustion of the fuel gas.

A further advantage of the method and system of the present invention is that the steam produced in the Fischer-Tropsch reformer and by the water/steam heat exchanger of the FTS reactor may also be recycled. Thus preferred methods of the invention further comprise recycling at least some of the steam from the Fischer- Tropsch reactor or reformer to the single tri-reformer reactor. In further preferred methods, substantially all of the steam from the Fischer-Tropsch reactor or reformer is recycled to the single tri-reformer reactor. Optionally the recycling may occur via one or more processing units. Hence in particularly preferred methods of the invention, the method is at least partially self-sufficient, and more preferably substantially self-sufficient, in terms of energy for the tri-reforming process and in terms of steam for the tri-reforming process.

In preferred methods and systems of the invention the first catalyst comprises one or more of nickel, magnesium, cerium, and zirconium. More preferably the first catalyst comprises Ce _(X) Zr ₍₁ _ _xr yNizMg and y and z are integers and x is less than one but greater than zero. Still more preferably the first catalyst is Ce ₀ . ₆ Zr ₀ .4- ₈ Ni ₈ Mg. The catalyst may be prepared by any method conventional in the art, e.g. deposition precipitation or wet impregnation, but is preferably prepared by wet impregnation. Preferably the surface area of the catalyst is 9-50 m ² /g and more preferably 22-50 m ² /g, as measured by BET analysis.

In preferred methods and systems of the invention, the tri-reforming process is carried out at 650-900 °C and more preferably 750-850 °C. Preferably the gas hourly space velocity (GHSV) of the tri-reforming process is 10,000 to 61 ,000 h ^"1 , more preferably 20,000 to 40,000 h ^"1 and still more preferably 25,000 to 30,000 h ^"1 .

In preferred methods and systems of the invention, the tri-reforming process achieves a C0 ₂ conversion of at least 27%, more preferably at least 45% and still more preferably at least 60%. Preferably substantially no coke is formed.

The H ₂ :CO ratio of the synthesis gas produced in the tri-reforming is approximately 2:1. This refers to a molar ratio. The ratio may be, for example, 1.5: 1.4 to 2.4:0.6. Preferably, however, the H ₂ :CO ratio is 2: 1.

The synthesis gas is converted to liquid fuel in Fischer-Tropsch synthesis (FTS) reformer comprising a second catalyst. In preferred methods and systems of the invention, the second catalyst is a cobalt-silica catalyst, more preferably an eggshelll cobalt-silica catalyst, still more preferably an eggshell cobalt-silica/titania catalyst, and yet still more preferably a promoted eggshell cobalt-silica/titania catalyst. Preferred promoter species may be selected from any one, or a combination of, Au, Ba, Ca, Ce, Cr, Cs, Cu, Fe, Hf„ K, La, Mg, Mn, Mo, Na, Nd, Pd, Pt, Re, Rh, Ru, Sr, Th, Ti, U, Zr. Preferably the surface area of the catalyst is 100-300 m ² /g and more preferably 150- 250 m ² /g, as measured by BET analysis.

In preferred methods and systems of the invention, the FTS process is carried out at 200-250 °C and more preferably 215-245 °C. Preferably the gas hourly space velocity (GHSV) of the FTS process is 1 ,600 to 17,158 h ^" , more preferably 2,000 to 10,000 h ^"1 and still more preferably 2,500 to 6,000 h ^"1 . The unique combination of eggshell catalyst and process conditions overcome heat and mass transfer limitations typically seen in FTS processes. This is done by shortening the diffusional length across the active catalyst surface to approximately 0.2mm-0.3mm allowing for reactants and hydrocarbon products to easily diffuse into and out of pores while limiting heat build up within the catalyst pores created by the exothermic reactions. Therefore heat is more readily removed in the FTS reactor allowing for precise control of the temperature within the catalyst bed with minimal temperature gradient along the length of the catalyst bed. This combination allows for a narrow boiling point range of liquid hydrocarbons to be produced especially concentrated around the middle distillate region while avoiding wax formation.

Preferably the liquid fuel produced in the method of the invention comprises C ₅ _ 30 hydrocarbons and more preferably C ₆ . ₂₈ hydrocarbons. Preferably the liquid fuel comprises paraffin and iso-parrafin hydrocarbons that make up at least 95.4 %wt of the liquid fuel product and with carbon numbers ranging from C5-C28. This carbon number distribution produces a fuel that can pass the flash point fuel specification for diesel, according to ASTM D93, while also passing the distillation specification for diesel, according to ASTM D86. The low aromatic content of the fuel produced contributes to an exceptional net heat of combustion (ASTM D3338) and contributes to less soot formation when combusted in diesel engines. These fuel qualities make the fuel produced an excellent candidate for a drop-in diesel fuel that can be utilized by today's diesel engines with no modifications to the engine needed. Preferably the Fischer-Tropsch synthesis (FTS) does not produce a significant proportion of waxes. Preferably less than 0.5 %wt and more preferably less than 0.2 %wt of the product of the Fischer-Tropsch synthesis (FTS) is wax. The definition of wax in this instance is taken to be hydrocarbons with more than 27 carbons in the chain. The production of wax is disadvantageous in the present method and system, because it requires extensive and expensive post treatment processes to separate and convert the waxes into usable transportation fuel (e.g. high temperature distillation, water gas shift reactor, hydrocracking, and other hydrotreatment type processing). The production of wax also depletes the amount of energy-containing fuel gases produced and thus curtails the ability of the process to be self-sufficient in terms of energy for the tri-reforming process.

In preferred methods of the invention the heat released in the Fischer-Tropsch synthesis is used to heat water and the heated water is recycled to the single tri- reformer reactor in the form of steam. This may be achieved, for example, by pumping water around the outside of the FTS reformer tubes to remove heat therefrom. During this heat exchange process, steam is produced that is preferably transported to the steam supply line of the single tri-reformer.

The present invention also relates to a method for producing liquid fuel from municipal solid waste comprising:

(i) biodegrading municipal solid waste in a landfill site to produce crude landfill gas;

(ii) separating particulate matter from said crude landfill gas to produce landfill gas; and

(iii) converting said landfill gas to liquid fuel as hereinbefore described.

In preferred methods and systems of the invention the particulate matter is separated from the crude landfill gas and the resulting gas is cleaned. These steps are conventional and may be carried out by any conventional method. The landfill gas may then be converted to liquid fuel by the method hereinbefore described.

The present invention also relates to a system for producing liquid fuel from landfill gas. Like the method of the invention, the system of the invention

advantageously recycles at least some of the energy of the fuel gas produced in the Fischer Tropsch reformer to the single tri-reformer reactor by combusting the fuel gas to produce heat. The system comprises:

an oxygen supply line;

a steam supply line;

a landfill gas supply line;

a fuel gas combustion unit that combusts fuel gas and provides heat a single tri- reformer reactor;

a single tri-reformer reactor for performing a tri-reforming process on the landfill gas comprising carbon dioxide reforming, steam reforming, water-gas shifting, and methane oxidation, to produce synthesis gas having a H ₂ :CO ratio of approximately 2:1 , wherein said tri-reformer has one or more inlets fluidly connected to said oxygen supply line, said steam supply line and said landfill gas supply line and at least one outlet for said synthesis gas;

a Fischer-Tropsch synthesis (FTS) reformer for converting the synthesis gas to liquid fuel, fuel gas and steam, wherein said FTS reformer has one or more inlets for said synthesis gas fluidly connected to the tri-reformer and an outlet for the liquid fuel, fuel gas and steam; a separation unit for separating said liquid fuel, said fuel gas and said steam, wherein said separation unit has an inlet for the liquid fuel, fuel gas and steam fluidly connected to the Fischer-Tropsch synthesis and separate outlets for each of liquid fuel, fuel gas and steam; and

a line connecting said fuel gas outlet of said separation unit to said fuel gas combustion unit.

Preferably the system comprises no other reactors that assist in generating the liquid fuel from the landfill gas. Preferably the system also does not comprise a hydrocracking unit, hydrotreatment unit, or a water gas shift reactor to produce H ₂ needed for the hydrotreatment processes. Such units are not required because the liquid fuel produced in the process of the invention does not require cracking.

A preferred system of the invention further comprises a line connecting the steam outlet of the separation unit to the steam supply line. Optionally the line is interrupted by one or more processing units. Thus, again like the method of the invention, the system of the invention advantageously recycles steam produced in the Fischer Tropsch reformer to the single tri-former reactor.

Preferably the tri-reformer reactor contains a first catalyst and the FTS reformer contains a second catalyst. Preferably the first is as described above in relation to the process. Preferably the second catalyst is as described above in relation to the process.

A preferred system of the invention further comprises a landfill gas combustion unit that combusts landfill gas and provides heat to the tri-reformer reactor. A further preferred system of the invention comprises a source of solar energy that provides heat to the tri-reformer. Preferably the source of solar energy comprises heliostats and a central receiver that contains heating media that is used to provide heat to the tri- reformer reactor. Optionally the system may further comprise a heat storage unit in which the heating media can be stored at an elevated temperature.

A further preferred system of the invention comprises a gasifier that extracts crude landfill gas from landfill biomass and provides the crude landfill gas to the tri- reformer.

Preferably the system of the invention is located on a landfill site.

Brief Description of the Drawings The present disclosure may be better understood with reference to the following figures. Matching reference numerals designate corresponding parts throughout the figures, which are not necessarily drawn to scale.

Fig. 1 is a block diagram of a first embodiment of a liquid fuel production system.

Fig. 2 is a graph that shows x-ray diffraction (XRD) profiles of example catalyst support materials.

Fig. 3 is a graph that shows temperature-programmed reduction (TPR) profiles of catalyst support materials.

Fig. 4 is a graph that shows TPR profiles of Ceo _.6 Zr _0.4 0 ₂ -8Ni (wetness impregnation), Ceo.6Zro _.4 0 ₂ -8Ni8Mg (deposition precipitation), and Ceo.e ro^O^s isMg (wetness impregnation).

Fig. 5 includes graphs that show energy dispersive x-ray spectroscopy (EDS) results for Ce ₀ . ₆ Zr _0.4 -8Ni8Mg loaded by wetness impregnation and deposition precipitation.

Fig. 6 is a graph that shows the surface spectrum by x-ray photoelectron spectroscopy (XPS) of the Ni2p _3/2 peak for reduced Ce ₀ . ₆ Zro.4-8Ni8Mg loaded by wetness impregnation.

Fig. 7 is a graph that shows H ₂ production for initial catalytic testing using steadily ramped temperature at 10°C/min with a gas composition of (CH ₄ :C0 ₂ :0 ₂ :H ₂ ) = (1 :1 :0.1 :0.1).

Fig. 8 is a graph that shows XRD profiles of Ce _0.6 Zr ₀ . ₄ -8Ni8Mg loaded by wetness impregnation before and after tri-reforming.

Fig. 9 is a gas chromatography distribution of liquid hydrocarbons using the HP- 5 column.

Fig. 10 is a block diagram of a second embodiment of a liquid fuel production system.

Fig. 1 1 is a bar chart showing the carbon number distribution of the fuel produced by the method of the invention compared to commercial diesel.

Detailed Description

As described above, it would be desirable to have alternative systems and methods for producing liquid fuels from municipal solid waste (MSW) and/or landfill gas (LFG). Disclosed herein are examples of systems and methods designed for this purpose. In some embodiments, the systems and methods use only two reactors to convert LFG into liquid fuel. In other embodiments, the systems further utilize solar energy to assist in the conversion.

In the following disclosure, various specific embodiments are described. It is to be understood that those embodiments are example implementations of the disclosed inventions and that alternative embodiments are possible. All such embodiments are intended to fall within the scope of this disclosure.

Fig. 1 illustrates a first embodiment of a fuel production system 10. As is shown in that figure, raw LFG is input into an LFG separating and cleaning unit 12 that separates particulate matter from the gas and removes trace level contaminants. This minimal pretreatment process greatly improves the overall process economics and eliminates any C0 ₂ removal steps. The LFG is primarily composed of methane (CH ₄ ) and carbon dioxide (C0 ₂ ).lt is important to distinguish LFG from natural gas or other gaseous hydrocarbons as LFG is not gaseous hydrocarbons. The CH ₄ component of LFG is highly diluted with typical compositions of LFG having CH ₄ in the range of 35- 60% and C0 ₂ in the range of 35-55%. Therefore typical reforming technologies applicable to natural gas or normally gaseous hydrocarbon streams are unable to tolerate the high amounts of C0 ₂ found in LFG for extended periods of time and deliver the desired H ₂ :CO ratios needed for FTS Cobalt based catalyst. By way of example, the LFG can flow into the cleaning unit 12 at a rate of approximately 3,000 standard cubic feet per minute (scf/min) to remove trace contaminants such as sulfides, halides, and siloxanes known to cause deactivation to Ni based reforming catalysts. After being cleaned, the LFG leaves the cleaning unit 12 and passes into a tri-reformer 18 described below. In addition, some of the LFG can pass into an LFG combustion unit 14 in which the LFG is combusted by mixing it with oxygen (0 ₂ ) or air to provide heat for the reaction in the tri-reformer 18. By way of example, the LFG is heated within the LFG combustion unit 14 to a temperature of approximately 600°C to 800°C. As is also shown in Fig. 1 , C0 ₂ and water can be output from the LFG combustion unit 14 as byproducts of the combustion process. In addition, excess LFG that is not needed for the production of liquid fuel can be delivered from the cleaning unit 12 to a flare unit 16 to be flared (i.e., combusted). The by-products of such flaring are C0 ₂ and water.

As noted above, LFG from the separating and cleaning unit 12 and the combustion unit 14 can be delivered to the tri-reformer 18 for processing. The tri- reforming process involves a combination of C0 ₂ reforming (Equation 1 ), steam reforming (Equation 2), water-gas shift (Equation 3), and methane oxidation (Equations 4 and 5) in a single reactor. CH ₄ + C0 ₂ = 2CO + 2H ₂ ΔΗ° = 247.3 kj/mol [Equation 1]

CH ₄ + H ₂ 0 = CO + 3H ₂ ΔΗ° = 206.3 kj/mol [Equation 2]

CO + H ₂ 0 = C0 ₂ + H ₂ ΔΗ° = -41 kj/mol [Equation 3]

CH ₄ + / ₂ 0 ₂ = CO + 2H ₂ ΔΗ ⁰ = -35.6 kj/mol [Equation 4]

CH ₄ + 20 ₂ = C0 ₂ + 2H ₂ 0 ΔΗ ⁰ = -880 kj/mol [Equation 5]

Use of the tri-reforming process eliminates the serious problem of carbon formation and high energy consumption commonly seen in C0 ₂ reforming by incorporating H ₂ 0 and 0 ₂ (see Equations 6, 7, 8, 9, and 10 below). Heat is generated in-situ by the addition of 0 ₂ , in the form of pure 0 ₂ or air, that can be used to increase energy efficiency and achieve a thermo-neutral balance of reactions. H ₂ and CO selectivity can also be adjusted by controlling the amount of steam and C0 ₂ added to the reaction. The flexibility of tri-reforming using three oxidant streams is an important advantage that distinguishes it from other bi-reforming technologies only utilizing two oxidant streams. The significance of performing tri-reforming vs. bi-reforming plays a major role in the energy, economical, and environmental impacts of the process. Compared to C0 ₂ reforming, tri-reforming consumes 45.8% less energy and produces 92.8% less C0 ₂ . Also compared to steam reforming, tri-reforming consumes 19.7% less energy and produces 67.5% less C0 ₂ . These are significant energy and environmental benefits that contribute to the overall efficiency and profitability of the tri- reforming process using substantially less energy and less C0 ₂ generation adding to the overall profitability and environmentally friendliness of the process. This provides an important role in both industrial and environmental applications allowing production of high-value chemicals via oxo-synthesis, electricity via solid oxide fuel cells or molten carbonate fuel cells, and clean-burning hydrocarbon fuels via Fischer-Tropsch synthesis (FTS).

Coke formation CH ₄ = C + 2H ₂ ΔΗ ⁰ = 74.9 kj/mol [Equation 6]

2CO = C + C0 ₂ ΔΗ° = 172.9 kj/mol [Equation 7] Coke Destruction

C (ads) + C0 ₂ = 2CO AH" = 172.2 kj/mol [Equation 8]

C (ads) + H ₂ 0 = CO + H ₂ AH =° 131 .4 kj/mol [Equation 9]

C (acte) + 0 ₂ = C0 ₂ ΔΗ ⁰ = -393.7 kj/mol [Equation 10]

The tri-reforming catalyst used in the tri-reforming process must be thermally stable, have a high surface area, have high oxygen storage capacity (OSC), have good redox properties, provide resistance to coke formation, and be economically advantageous. Ni-based catalysts have shown good potential for reforming methane and provide a more economically friendly option over noble metals. However, Ni has the disadvantage of being susceptible to coke formation. Deactivation is directly related to the catalyst structure and composition and, therefore, research has been aimed at producing a suitable catalyst in the upgrading of MSW synthesis gas (or "syngas"). Ce0 ₂ has a high oxygen storage capacity (OSC) and can be used as a promoter with Ni for methane conversion to syngas. The addition of Zr0 ₂ to Ce0 ₂ has been shown to improve OSC, redox property, thermal stability, metal dispersion, selectivity, and catalytic activity. These improved characteristics are attributed to the formation of a (Ce, Zr)0 ₂ solid solution. Others have attributed the increased OSC from modifying the local oxygen environment around Ce and Zr and generating active oxygen. This result may be explained by the introduction of undersized Zr ions into the Ce framework that helps compensate for the volume increase associated with the valence change of Ce ⁴⁺ to Ce ³⁺ , easing the transition. Research has shown that the Ce/Zr mixed oxides consistently perform with higher activity compared to the pure oxide supports and Al ₂ 0 ₃ due to its ability to promote partial oxidation of methane (POM) and steam-reforming reactions. Because of this, Ce _1-x Zr _x 0 ₂ support materials have received much attention with 0.6<x<0.8 being preferred for catalytic applications. Basic oxides, such as magnesia and zirconia, have been shown to catalyze the gasification of coke with steam and help prevent deposition of carbon in dry reforming. This phenomenon may be attributed to the low concentration of Lewis sites and increase of oxygen vacancies by introducing Zr0 ₂ and MgO into the catalyst composition. By coupling these basic oxides with Ni, catalysts promote C0 ₂ and H ₂ 0 adsorption leading to enhanced C0 ₂ conversion and H ₂ production. The enhanced C0 ₂ conversion has been attributed to a higher interface between Ni, MgO, and Zr0 ₂ resulting from NiO/MgO and Zr0 ₂ /MgO solid solutions.

Experimentation was performed to determine which catalysts would be best for liquid fuel production. This experimentation and catalyst synthesis is described in the following paragraphs.

Ce/Zr oxide supports were prepared using the co-precipitation method reported by Rossignol et al. using Ce(N0 ₃ ) ₃ x 6H ₂ 0 and ZrO(N0 ₃ ) ₂ x H ₂ 0 as precursors. Pure ceria and zirconia oxides, as well as the mixed oxides with Ce:Zr molar ratios of 0.16:0.84, 0.6:0.4, and 0.8:0.2 were all prepared using the same method. Appropriate quantities of the precursor salts were dissolved in deionized (Dl) water and precipitated by the addition of NH ₄ OH to form hydrous zirconia, ceria, or Ce/Zr solution. This precipitate was vacuum-filtered and re-dispersed into a 0.25 M NH ₄ OH solution. This dilute, basic solution was again vacuum-filtered and dried in an oven at 120°C overnight. The dried powder was then calcined at 800°C for 4 hours.

The loading of Ni and Mg to the oxide support was carried out using two different loading procedures: wet impregnation (Wl) and deposition precipitation (DP). All metals were loaded on a mass basis to achieve desired weight percentage of metal on the catalyst. For the Wl method, appropriate amounts of Mg(N0 ₃ ) ₂ x H ₂ 0 and Ni(N0 ₃ ) ₂ x 6H ₂ 0 were dissolved in deionized water to form a homogeneous solution. This solution was then added drop-wise to the support until incipient wetness and dried at 120°C for 2 hours. This step was repeated until all of the metal nitrate solution had been added to the support. Following the final drying step, the catalyst was calcined at 500°C for 4 hours. DP was performed using a modified method adapted from Li et al. Appropriate amounts of Mg(N0 ₃ ) ₂ x H ₂ 0 and Ni(N0 ₃ ) ₂ x 6H ₂ 0 where added to a volumetric flask and dissolved in 25 ml Dl water. The powder support was added to the metal-salt solution and mixed with a stir plate to form a slurry. In a separate beaker, CO(NH ₂ ) ₂ (urea) was added in excess to 10 ml of Dl water to achieve a 1 :4 ratio of total metal nitrates: urea. The urea solution was added drop-wise to the metal-salt solution while stirring. The top of the volumetric flask was sealed to prevent evaporation of the solution and heated to 1 15°C while stirring at 600 rpm on a heated stir plate. Urea hydrolyzes slowly at temperature allowing hydroxyl groups to react rapidly as they form, maintaining a constant pH and allowing precipitation on the surface and interior of pores. The solution was aged for 24 hours and then cooled to room temperature before vacuum-filtering with a Buchner funnel. Cold DI water was used to wash any remaining precursors and impurities from the filtered catalyst. The catalyst was then dried at 120°C for 4 hours followed by calcination at 500°C for 4 hours.

Braunauer-Emmett-Teller (BET), x-ray diffraction (XRD), temperature- programmed reduction (TPR), scanning electron microscopy-energy dispersive x-ray spectroscopy (SEM-EDS), and x-ray photoelectron spectroscopy (XPS) were used to characterize catalysts. The combination of these techniques provides valuable data that aids in the catalyst design by providing insights into physical and chemical structure. Physisorption experiments were performed using a Quantachrome Autosorb- iQ. The BET surface area was calculated using data in the P/P ₀ range of 0.05-0.3, where a linear relationship for the BET isotherm is maintained.

XRD analysis was performed with a Philips X'pert XRD using a powder x-ray diffraction technique. The machine was operated in a Bragg angle (2Θ) range of 15-80°. The step size was 0.06° and a dwell time of 1 s was used for each step. X'pert Highscore software was used to assist in data analysis.

TPR was performed using the Quantachrome Autosorb-iQ, mentioned above, using 50 mg of catalyst loaded into a quartz sample cell. Each sample was pretreated with helium while ramping the temperature 10°C/min from 25°C to 1 10°C and holding at temperature for 30 minutes. The sample was then cooled to 50°C. Following pretreatment, the carrier gas was switched to 5% H ₂ /N ₂ and the temperature ramped to 1 100°C at 10°C/min. Gas analysis was performed using a thermal conductivity detector (TCD) measuring the conversion of H ₂ under the temperature-programmed conditions.

A Hitachi S-800 SEM coupled to an Ametek EDAX was utilized to conduct SEM-EDS experiments. An excitation energy of 10 keV, a magnification of 1010, and a tilt angle of 30° were used in this analysis.

XPS measurements were taken with a Perkin-Elmer PHI 560 ESCA/SAM system under vacuum using an Mg filament. Binding energies were scanned in the 0- 1030 eV range initially. A high resolution scan was performed on the Ni 2p ₃ peak in the binding energy range of 849-869 eV. RBG AugerScan 3 software was used for data analysis of the resulting spectrum. Catalytic reactions were performed in a fixed-bed quartz u-tube reactor (ID = 4mm) at 1 atmosphere. Feed gas composition was controlled using Alicat Scientific mass flow controllers and adjusting the flow rates accordingly. Online analysis of the product gas was taken with a MKS Spectra (Cirrus) mass spectrometer (MS) connected in-line with the reactor. Before each experiment, the quartz reactor was loaded with 75.2 mg of catalyst into the bottom third of the quartz tube and supported on either side by inert quartz wool. The reaction vessel was positioned inside a Thermoscientific Thermolyne tube furnace. Reaction temperature was controlled by adjusting the furnace temperature program to the desired ramp rate or fixed temperature. Heat tape was used to heat reactant and product lines to prevent condensation from occurring prior to the catalyst bed and MS detector. Water was delivered to the reactant gas mixture through a heated water bubbler using helium as a carrier gas. All catalysts were first reduced with 10% H ₂ in He while ramping the temperature from room temperature to 800°C at 10°C/min and holding for 2 hours. After reducing the catalyst, bypass valves were used to stop flow through the reactor while the reforming gas mixture was adjusted to the desired composition. The valves were then reopened after the MS gave stable responses for each of the reactants. A gas hourly space velocity (GHSV) of 61000 hr ^"1 was employed for all tri-reforming reactions, unless otherwise specifically stated. Conversion of CH ₄ and C0 ₂ were calculated using the following formulae:

CH _A conv. = 1 - {mol CH ₄ in product ÷ mol CH ₄ in feedj [Equation 1 1 ]

C0 ₂ conv. = 1 - [mol C0 ₂ in product ÷ mol C0 ₂ in feed j [Equation 12]

Immediately following each reaction, a temperature-programmed oxidation (TPO) was performed to quantify any coke present on the surface of the catalyst. After reactions, catalysts were quickly cooled to 1 15°C under an inert (He) environment. The temperature was then ramped at 10°C/min to 700°C and held for 1 hour as flow rates of 0 ₂ = 2.5 SCCM and He = 50 SCCM were used to oxidize the catalyst and convert surface coke to C0 ₂ . Essentially all carbon was converted to C0 ₂ with insignificant amounts of other carbon-containing species produced. The product gas was analyzed by a MS detector and quantified by integrating the peak areas to determine the amount of carbon present as coke. TPO was used to measure the amount of coking and is reported in this study as average rate of coke formation per mass catalyst. This number is given as the mass of carbon deposited as coke/mass catalyst reaction time.

To understand how support composition, metal loading, and preparation influenced the surface area of the catalyst, BET analysis was performed and compared on multiple samples (Table 1). The pure oxide species had significantly lower surface areas than any of the mixed oxide supports, which suggests that the mixed oxide supports are not simply a mechanical mixture of the two species. Instead, a new oxide material with different physical properties from either of its pure components had been synthesized. This suggests a solid solution of Ce and Zr oxides formed using the co- precipitation technique. As more Ce is introduced into the structure of the catalyst, the surface area also increased. This effect reaches a maximum at a Ce:Zr ratio between 0.8:0.2 and 1 :1 because the pure Ce oxide material has a dramatically lower surface area than the highest Ce content sample tested here (Ce:Zr=0.8:0.2). Upon loading of the Ni and Mg metals to the surface of each mixed oxide support material, the surface area decreased slightly. This is attributed to metal crystals forming within pores of the support and, in some cases, blocking the pathway.

Table 1: BET surface area for various supports and catalysts (8Ni8Mg refers to 8 % metal loading by wt. for each).

Catalyst support materials were analyzed by XRD (Fig. 2) and compared to elucidate structural differences as the composition was altered. This figure compares the crystal structure of pure oxide species and the mixed oxide support with a Ce:Zr ratio of (0.6:0.4) . Miller indices are also represented for each peak in Fig. 2.

Pure Zr0 ₂ is known to exist in the tetragonal and monoclinic phases. The XRD pattern of pure Zr0 ₂ obtained from this experimentation closely resembles characteristic peaks of the monoclinic phase. This is typical of Zr0 ₂ samples that have been calcined at higher temperatures. The XRD pattern from the pure Ce0 ₂ shows characteristic peaks for a cubic fluorite structure. However, when these two pure oxide species were co-precipitated, no peaks could be identified that indicated a monoclinic Zr0 ₂ species and all peaks resembled the cubic fluorite structure found in pure Ce0 ₂ . This suggests that Zr0 ₂ is incorporated into the Ce0 ₂ lattice and that a solid solution formed from the combination of these two oxide species. Peak broadening is seen in the mixed oxide sample compared to the pure oxides and is most likely due to lattice defects from the insertion of the smaller Zr cation into the Ce0 ₂ lattice.

As shown in Fig. 3, different reduction peaks existed for the pure and mixed oxide supports. Both the pure Ce0 ₂ and Zr0 ₂ show much higher temperatures needed to reduce these species compared to the mixed oxide. The pure Ce0 ₂ support shows a maximum reduction peak around 865°C, while the pure zirconia support shows no reduction occurring at temperatures up to 1100°C. When these pure species were combined to form a mixed oxide support, a much lower reduction peak is seen to occur between 300-650°C with a max adsorption peak at 555°C. This lower reduction temperature is attributed to a (Ce, Zr)0 ₂ solid solution forming with similar trends seen for other Ce:Zr ratios. The first and second reduction peaks in the mixed oxide are due to the surface and bulk reduction, respectively, and can be explained by the Binet et al. model for Ce reduction. Incorporation of Zr ions facilitates the valence change of Ce by enabling the volume change associated with the reduction of Ce. By incorporating Zr within support framework, oxygen mobility is increased, allowing oxygen migration between nearby cation channels. From the TPR experiments, it is seen that incorporating Zr0 ₂ into Ce0 ₂ to form a mixed oxide improves oxygen storage capacity (OSC) and redox properties. Zr0 ₂ is also a more thermally stable compound that improves the mixed oxides' catalytic activity at the elevated temperatures used in reforming reactions.

TPR was also utilized to gain a better understanding of how the Ni interactions between Mg and the support are affected when using different metal loading techniques. The TPR profiles (Fig. 4) of catalysts loaded with Ni and Mg using Wl and DP methods are compared to a catalyst with only Ni loaded via Wl. Interestingly, when Ni and Mg were loaded by DP, the reduction profile closely resembled that of the catalyst with only Ni loaded onto the surface. When Ni and Mg were loaded by Wl, most of the reduction occurred at higher temperatures. The lower temperature reduction peaks seen are associated with isolated Ni and weakly interacting Ni with the support and Mg. The higher temperature reduction seen in the Wl catalyst is indicative of a strong interaction occurring between Ni and Mg. This result was found as a surprise since DP is usually associated with higher dispersion of smaller particles and thus stronger interactions. The lower reduction peak in the DP-prepared catalyst could be attributed to higher dispersion causing fewer interactions between the Ni and Mg. However, upon further experimentation using SEM coupled with EDS (Fig. 5), it was determined that less Mg had been loaded onto the DP-prepared catalysts compared to the Wl-prepared catalyst and explains why the reduction peak of the DP catalyst resembled the catalyst with only Ni loaded onto the surface by Wl. The high temperature reduction peak in the Wl catalyst containing Ni and Mg species is thus attributed to more interfaces between Ni and Mg with stronger interactions between them.

XPS was utilized to measure the binding energies of various components present in the reduced catalysts prepared by Wl with the support makeup of Ce ₀ . ₄ Zr ₀ .6O ₂ . An initial broad range scan was performed to identify the major species present and the binding energies associated with these species.

A high resolution scan of the Ni 2p _{3 2} peak centered at a binding energy near 856 eV was then performed to identify the interactions between the Ni, mixed oxide support, and MgO of the catalyst loaded with 8 wt% Ni and Mg. A curve-fit summary was produced from this scan, the results of which can be found in Fig. 6 where the majority of Ni is oxidized. The major peak (856 eV) is associated with oxidized Ni and could be associated with interactions to the mixed oxide support, MgO, or hydroxyl groups. The second (near 862 eV) and third largest (near 865 eV) peaks are attributed to a satellite peak of the main peak. The small peak near 853 eV is the only signature of metallic Ni. These results indicated strong interactions with the Mg and the mixed oxide support, but only limited conclusions can be made because of the complex spectrum of Ni and possible oxidation at the surface.

Various catalyst formulations were tested to study the consequences of altering the support mixture and the ratios/amounts of metals loaded onto the catalyst. These catalysts were each tested under the same conditions while steadily ramping the temperature. All catalysts were prepared using the same Wl preparation technique. Results were compared and are shown in Fig. 7, which illustrates the H ₂ production from various catalysts tested while steadily ramping the temperature at 10°C/min. H ₂ production is only shown in Fig. 7 for ease of interpretation, but C0 ₂ and CH ₄ conversions were also analyzed. The effects of varying the support makeup ratio were compared by holding the metal weight percentage and ratios constant. The lower Ce:Zr ratios of 0.16:0.84 in the support led to lower H ₂ production, CH conversion, and C0 ₂ conversion. When Ce:Zr ratios were increased to 0.8:0.2, the H ₂ production, CH ₄ conversion, and C0 ₂ conversion slightly increased. Adjusting Ce:Zr ratios to a more even ratio of 0.6:0.4 gave the best results with the highest H ₂ production, CH conversion, and C0 ₂ conversions. This finding is in agreement with the charge channeling effect created by nearby cations. By incorporating a more even ratio of Ce:Zr, oxygen mobility and redox properties are improved, allowing transport of oxygen to appropriate sites and preventing coking on the Ni metal surface. Thermal stability was seen in all mixed oxide support ratios and is attributed to Zr0 ₂ high thermal stability.

The impacts of metal loading ratios and weight percentage on the catalyst were explored by holding the support ratio constant and varying the metal loading quantities. Different Ni amounts (4% and 8%) were loaded onto the same support composition. In all cases, the lower weight percentage of Ni lead to a plateau effect and a low H ₂ production was seen. The plateauing effect describes the tendency of the production of a certain compound, in this case H ₂ , to remain unchanged even when temperature is increased. When 8% Ni was loaded onto the catalyst, the plateauing effect on H ₂ production at higher temperatures is no longer seen. The amounts and ratios of Mg were also varied to study its effect on catalyst performance. Again, catalysts with the same support composition and metal-loading technique were compared while varying the metals loaded onto the surface. The catalyst with no Mg loaded, suggesting coking on this sample onto the surface, had the slowest rates of H ₂ production and quickly plateaued even when higher Ni amounts were loaded onto the surface. At a Ni:Mg weight percentage ratio of 2:1 , H ₂ production rates were increased and no plateauing of the H ₂ production was seen with a steady rise in production as temperatures were increased. Higher amounts of H ₂ and the fastest rate of H ₂ production were seen when Ni:Mg weight percentage ratios of 1 :1 were loaded onto the catalyst surface. This effect can be explained by the facilitation of the redox mechanism involved in methane reforming with increased interface between Ni and Mg. Metal weight percentage ratios approaching unity gave more interfaces between Ni and Mg. This facilitates C0 ₂ adsorption/dissociation and oxygen movement to the reduced Ni surface where it could react with the adsorbed carbon from CH , Basic promoters like MgO have an affinity for C0 ₂ due to its acidic nature. This is an added advantage in C0 ₂ reforming because C0 ₂ is normally a very stable molecule and a catalytic reaction is needed for quick dissociation.

Reactions were studied under controlled temperature programs with the optimum temperature range found between 750-850°C. At the lower end of this range, higher H ₂ :CO ratios were produced due to the steam reforming and water-gas shift (WGS) reactions (Equations 2 and 3) being more favorable at these temperatures. However, lower C0 ₂ conversions were obtained at the lower temperatures. Because C0 ₂ reforming is favorable at high temperatures, it was determined that C0 ₂ conversion increased with increasing temperature in this range. At 800°C, C0 ₂ conversions remained high and desired H ₂ :CO ratios could be achieved without catalyst deactivation. At this temperature, coke gasification reactions can occur while maintaining high levels of steam reforming and POM to produce desired H ₂ :CO ratios. At higher temperatures, C0 ₂ conversion increases but H ₂ :CO ratios dropped not only due to the increase in CO production, but also to less H ₂ production. This result occurs because H ₂ production decreases as C0 ₂ reforming dominates the reaction making steam reforming and POM reactions less favorable at higher temperatures.

Gas composition greatly affects the reaction products. In tri-reforming, many reactions are occurring at one time and finding the correct ratios of reactants is not trivial. During tri-reforming reactions, it was found that conversion of 0 ₂ was the highest of all oxidants, the 0 ₂ completely being consumed. Oxygen seems to have a high affinity for active sites on the catalyst and tends to react quickly. Remaining active sites or those where 0 ₂ had already disassociated are available for the other reactants. H ₂ 0 and C0 ₂ compete for the same active sites. Therefore, experiments were performed to understand how altering these two reactant concentrations influenced product ratios. Table 2 helps explain these effects and shows that increasing the H ₂ 0 ratio in the feed will increase the H ₂ :CO ratio. However, there is a point at which higher H ₂ 0 ratios led to a decrease in C0 ₂ conversion. One of the goals in tri-reforming is to maintain high C0 ₂ conversions while still producing desired H ₂ :CO ratios. High C0 ₂ conversion made the process more environmentally friendly and improved efficiency in FTS for liquid hydrocarbons. The results in Table 2 suggest that the adsorption of H ₂ 0 blocks the C0 ₂ adsorption sites leading to higher H ₂ :CO ratios and inhibition of C0 ₂ reforming. Lower than expected H ₂ 0 concentrations in the feed gas were found to produce high concentrations of H ₂ without greatly sacrificing C0 ₂ conversion. From Table 2, a CH :C0 ₂ :H ₂ 0:0 ₂ ratio of 1 :0.7:0.23:0.2 produced desired H ₂ :CO ratios above 2. This result demonstrates that optimum syngas composition for FTS applications can be achieved while maintaining high C0 ₂ conversions at lower H ₂ 0 ratios utilizing CH ₄ :C0 ₂ feed ratios resembling typical LFG compositions. At these conditions, the catalyst still showed a high resistance to coke formation on the catalyst surface. Table 2: Comparison of reaction results at T = 800°C with

Ce ₀ . ₆ Zro.402-8Ni8Mg (wet impreg.) for a variety of gas feed ratios.

An added benefit of the current results for FTS are the low steam amounts because steam is reported to deactivate catalysts. At the reaction temperature of 800°C and the composition ratios mentioned above, tri-reforming over Ceo _.6 Zr ₀ .4- 8Ni8Mg produced an upgraded syngas with desired H ₂ :CO ratios for FT applications that achieved C0 ₂ conversions above 76% and maintained resistance to coke formation at lower steam ratios. Negligible levels of coke were detected in TPO experiments and catalyst activity remained high at the above-reaction conditions. The ability to maintain high levels of C0 ₂ conversion without deactivation becomes a highly attractive option since C0 ₂ in FT feedstock syngas increases the H ₂ demand and H ₂ :CO ratios higher than 2 will be needed to produce low concentrations of olefins and oxygenates in the FT synthesized product.

In an effort to determine the effect of the GHSV on the product composition and insight into which reactions are occurring, the amount of catalyst was increased. The increase in catalyst amount forced reactant gas residence times to be longer (Table 3). The amount of catalyst used in the experiments ranged from 2.5-2.9 times (186-218 mg) the amount used in previous studies (i.e., 75 mg). A feed gas CH ₄ :C0 ₂ :H ₂ 0:0 ₂ ratio of 1 :0.7:0.5:0.2 was fed to the reactor. GHSV was calculated to be approximately 21000 hf ¹ and 25000 hr ¹ when 218 mg and 186 mg catalyst, respectively, were used to perform the reaction. Whereas CH ₄ conversions remained relatively unchanged, C0 ₂ conversions were slightly lower and H ₂ :CO ratios were significantly reduced compared to the previous studies in which the GHSV-61000 hr ^"1 . It is proposed that as the feed gas initially reacts and creates higher H ₂ concentration, the reverse WGS reaction becomes more favorable further down the catalyst bed. This could be an indication of steam reforming reactions (Eqs. 2 and 3) approaching equilibrium. This suggests that there may be an advantage to using higher GHSV to maintain higher H ₂ production.

Table 3: GHSV comparison at T = 800°C with Ce ₀ . ₆ Zro. ₄ 0 ₂ -8Ni8Mg

(wet impreg.) for a gas feed ratio of CH4:C0 ₂ :H ₂ 0:0 ₂ = 1 :0.7:0.5:0.2.

By decreasing residence time, the ability to limit reactions that consume H ₂ may be possible. However, even at the lower GHSV conditions, H ₂ :CO ratios were maintained between 1.55-1.66. Therefore, if H ₂ supplementation is needed for FT processing of the tri-reformed gas, the amount of H ₂ needed to be added to the tri- reforming process will be significantly lower than compared to more traditional reforming processes. These other reforming processes will also be significantly more expensive as higher amounts of steam will be needed and/or coking reactions will limit catalyst lifetime.

Post-reaction characterizations were performed on the Ce ₀ . ₆ Zr _0. -8Ni8Mg loaded by Wl. After 4 hours of reaction at 800°C and a feed gas CH ₄ :C0 ₂ :H ₂ 0:0 ₂ ratio of 1 :0.7:0.5:0.2, the catalyst surface area decreased from 34.5 to 22.1 m ² /g. This change is attributed to using a higher reaction temperature than the final calcining temperature used to synthesize the catalyst. Comparing initial and final conversions, catalyst performance appeared to be minimally influenced by the change in surface area (Tables 2 and 3). The formation of coke did not appear to play a role in the change in catalyst surface area. This finding is supported by the negligible amounts of coke detected using post-reaction TPO experiments. No evidence of crystalline carbon is present in the XRD pattern of the post-reaction sample (Fig. 8).

The post-reaction sample shows Ni in the reduced form, which is expected due to the high production of H ₂ during the reforming reaction. Peaks characteristic of reduced Ni show higher intensity while the characteristic peaks for (Ni,Mg)0 decreased in the post-reaction sample, indicating that Ni species in the (Ni, Mg)0 solid solution are reducible under reaction conditions for those catalysts prepared by Wl. This was an excellent result because the deactivation of Ni-reforming catalysts has been attributed to the inability to reduce Ni from an inactive oxide phase to a reduced Ni phase. Post- reaction samples showed the same (Ce, Zr)0 ₂ pattern as the pre-reaction sample, indicating that the cubic fluorite phase is stable under the reaction conditions employed.

In view of the above discussion, the catalyst used in the tri-reformer 18 of Fig. 1 comprises a mixture of nickel (Ni), magnesium (Mg), cerium (Ce), and zirconium (Zr). In some embodiments, the catalyst comprises Ce _{X) Zr _(1-X) -yNizMg. In such a case, x is the molar amount of Ce and (1 -x) is the molar amount of Zr in the support, while y and z indicate the mass % loading of Ni and Mg, respectively. The value of x is 0 < x > 1. In some embodiments, y and z are integers and x is less than one but greater than zero. As indicated above, in one example, x = 0.6, y = 8, and z = 8, in which case the catalyst is Ceo. ₆ Zr ₀ .4-8Ni8Mg.

With further reference to Fig. 1 , the tri-reformer 18 alters the ratios of the various components of the LFG to one in which FTS can be performed to produce liquid fuel. More particularly, the tri-reformer 18 produces synthesis gas that has a H ₂ to CO ratio of approximately 2:1 , meaning that the synthesis gas contains twice as much H ₂ than CO. In addition to H ₂ and CO, the synthesis gas may contain C0 ₂ and water vapor. In some embodiments, the tri-reformer 18 is configured as a packed-bed reactor and the LFG is flowed through the catalyst at an elevated temperature in the range of approximately 600 to 800°C. Although that temperature can be maintained by further combustion of the LFG, additional energy can be input into the tri-reformer 18, as indicated in Fig. 1 , to ensure the desired temperature is maintained. As is also shown in Fig. 1 , water can be provided to the tri-reformer 18 to assist in the reaction.

As noted above, the output from the tri-reformer 18 is synthesis gas having a H ₂ :CO ratio of approximately 2:1. Because the temperature of that synthesis gas is higher than is needed for FTS, the gas can be cooled using a syngas heat recovery unit 20, which lowers the temperature of the gas to approximately 200°C to 220°C. In some embodiments, the heat recovery unit 20 can comprise a heat exchanger and the extracted heat energy can be used for other purposes, such as heating the tri-reformer 18.

Once the synthesis gas is at the desired temperature, it is provided to the second reactor of the system 10, the FTS reformer 22. The FTS reformer comprises a further catalyst that converts the synthesis gas into liquid fuel. In some embodiments, the catalyst is a cobalt-silica catalyst. In testing, silica supported cobalt eggshell was used as the active catalyst material for the production of liquid hydrocarbon from the resultant syngas. The choice of this eggshell catalyst was based on the desire to increase the selectivity towards middle distillate products. Silica gel support was selected mainly due to its inertness, high surface area, and versatile nature (hydrophobic/hydrophilic).

The catalyst, along with conductive inert particles, was placed in a fixed bed reactor for the conversion of syngas. The bench scale reactor comprised a cylindrical tube having 0.75 inch OD (1.905 cm) and 17 inch (43.18 cm) length. The Co/Si0 ₂ eggshell catalyst was first reduced in pure hydrogen at 673 K (400 °C). After reduction for 16 hours, the reactor temperature was reduced to 453 K (180 °C) and syngas mixed with hydrogen (to get the appropriate 2:1 ratio of H ₂ to CO) was delivered to the fixed bed reactor at a rate of 0.7 N L/min. The choice of flow rate was based on recommended values of space velocity in which the favorable range (for CO conversion) is from 2-10 L/g/h. The weight hourly space velocity in this process was 2.0 L/g(reactor contents)/hr. Maximum conversions have been earlier reported at this space velocity. After adjusting the flow rate, temperature was gradually raised to 473K (200°C) to carry out the Fischer Tropsch reaction (Pressure = 2 MPa). The temperature was then raised to 493K (220°C). Based on the fact that a temperature of 493 K (220°C) will result in heavier chain growth for an eggshell catalyst and less methane, the operation was continued at this temperature.

Precise control of the catalyst bed temperature during the startup (pore filling time) of FTS is essential to avoid thermal runaway. To overcome this limitation, inert materials, such as silicon carbide, having high thermal conductivity were added to the fixed bed. Active catalyst and SiC were effectively mixed at a ratio of 1 :3 within the reactor.

Table 4 summarizes results at the end of five-day operation of the fixed bed reactor with biomass derived syngas. As expected, the eggshell morphology resulted in high selectivity of middle distillates. In previous work by the inventors on pure gases, it was identified that a temperature of 483 K (210°C), results in significant production of lighter hydrocarbons. The current operation at 487 K (214°C) reduced the fraction of lighter hydrocarbons (C _1-4 ) produced when compared with the earlier work. The formation of C0 ₂ is still high, however some of the previous research work on biomass has reported this number even at lower conversions with minimal C0 ₂ in the feed. The CO conversion was lower than pure surrogates reported earlier, due to the presence of inert component (C0 ₂ / N ₂ /hydrocarbons). The kinetic equations provided by other researchers suggest that the rate is dependent on temperature and partial pressure of H ₂ and CO. For a same total pressure, the partial pressure of reactive components decreases in the presence of inert components. However, higher conversion (75 vs. 60%) has been considered in modeling because of the effective removal of inert C0 ₂ and the absence of N ₂ .

Table 4: Eggshell Catalyst performance with biomass derived syngas under FTS conditions i.e. 503 K and 2.0 MPa

b) Space time yield of hydrocarbon with carbon number greater than 5

Due to the optimization of eggshell design and tight control of reaction parameters, the C ₅ + selectivity was high, as shown in Table 4. Fig. 9 represents GC distribution of liquid hydrocarbons using HP-5 column. Analysis by mass spectrometer (Agilent 5975C) showed the presence of alcohols and olefins in addition to the expected paraffinic hydrocarbons. Hence, oxygenates are effectively produced in the FTS process with a cobalt catalyst. The presence of isomers is also visible between the bands of paraffin. These isomers enhance the octane/cetane value of the fuel. As shown in Fig. 9, the cobalt catalyst showed excellent reproducibility over the duration of test run.

With reference back again to Fig. 1 , it is noted that the use of only two reactors (i.e., the tri-reformer 18 and the FTS reformer 22) is unique because existing technologies typically require three separate reactors, including a WGS reactor. In the system 10, however, there is no WGS shift reactor. Therefore, the system 10 simplifies the process and is less costly to construct. The ability to use two reactors instead of three in large part is the result of the conditions within the tri-reformer 18 and the nature of the catalyst, which is specifically suited for a mixture of methane and carbon dioxide found in the LFG. The unique combination of the conditions and catalyst used in the reactor enable the production of synthesis gas in the desired hydrogen to carbon monoxide ratio. The liquid fuel produced by the FTS reformer 22 is delivered to a liquid fuel heat recovery unit 24 in which the fuel is cooled. In some embodiments, the heat recovery unit 24 can also comprise a heat exchanger to achieve this cooling.

The liquid fuel produced by the FTS reformer 22 may contain different types of fuels, such as diesel fuel and jet fuel. In such a case, the fuels can be separated using a liquid fuel separation unit 26. As is shown in Fig. 1 , outputs from the separation unit 26 can include water, which can be delivered to the tri-reformer 18 as steam, and fuel gas (e.g., CH ₄ ) that can be provided to the flare unit 16 and a fuel combustion unit 28, which can be used to provide heat energy to the tri-reformer. If further types of fuel, such as gasoline, are desired, a liquid fuel refining unit 30 can be used to produce that other fuel. For example, if gasoline is desired, the diesel fuel can be cracked to produce the gasoline.

Fig. 10 illustrates a second embodiment of a fuel production system 50. The system 50 is similar in many ways to the system 10 described in relation to Fig. 1 , but utilizes solar energy to convert LFG into liquid fuel. Like the system 10, the system 50 includes a tri-reformer 52 that produces synthesis gas having a hydrogen H ₂ :CO ratio of approximately 2: 1 , an FTS reformer 54 that converts the synthesis gas into liquid fuel, and no further reactor, such as a WGS reactor. The liquid fuel provided by the system 50 can be separated by a liquid fuel separation unit 56 and refined by a liquid fuel refining unit 58. As is further shown in Fig. 9, H ₂ -rich gas from the FTS reformer 54 can be provided back to the tri-reformer 52, as can steam and crude synthesis gas from other sources described below. In addition to the synthesis gas from the tri- reformer 52, 0 ₂ and/or air can be input into the FTS reformer 54, as can water from the liquid fuel separation unit 56.

Instead of using combustion to provide the heat needed for the reaction in the tri-reformer 52, the system 50 utilizes solar energy generated using a solar collector. In the illustrated embodiment, the solar collector includes heliostats 60 that focus the sun's energy on heating media within a central receiver 62. The heated media can then be stored in a heat storage unit 64 and, when needed, can be provided to the tri- reformer 52. Optionally, an auxiliary heater 66 can be used to heat the media within the storage unit 64. In such a case, the heater 66 can be driven with the exhausted heating media from the tri-reformer 52 and/or electricity from a source described below.

In some embodiments, the system 50 can also generate liquid fuel from biomass from the landfill. In such a case, the biomass can be input into a feedstock pretreatment unit 68 that pretreats the biomass by, for example, drying it and removing components that cannot be used in the fuel generation process (e.g., metal, glass, etc.). The treated biomass can be provided to a gasifier 70 that extracts crude synthesis gas from the biomass. This is accomplished by adding heated 0 ₂ , air, and steam to the gasifier 70 and heating the mixture. The energy needed to heat the mixture can, for example, be provided by a steam turbine 72 that operates using steam output from the FTS reformer 54. The electricity produced by the turbine 72 can also be provided to the auxiliary heater 66 described above. The heated 0 ₂ , air, and steam can be provided to the gasifier 70 from the FTS reformer 54. The crude synthesis gas that is output from the gasifier 70 can then be provided to the tri-reformer 52.

Example

A detailed gas analysis was first done to determine the exact composition of the LFG gas collected at Sarasota Landfill. Table 5 shows the results of the analysis. Filter beds composed of Sulfatreat, silica gel, and activated carbon were used to remove the sulfides, siloxane, and halide compounds respectively.

Table 5

Table 6 below outlines the conditions and results for the tri-reforming reaction and the product synthesis gas H2:CO ratios fed to the FTS using a iMg/Ceo. ₆ Zr _0.4 0 ₂ (r=0.75) reforming catalyst over the course of the four day LFG run. Catalyst Bed CH ₄ conv. C0 ₂ conv.

CH ₄ :H ₂ 0 GHSV f h ¹ ) H ₂ :CO temp. (°C) ( ) (%)

780-785 0.56 30,000 99 59-61 1.70-1.74

Table 6

There was no evidence of deactivation by the reforming catalysts as methane conversions remained above 99% for the entire duration of the experiment. This was confirmed by TPO experiments conducted on the used reforming catalyst showing that only 0.0062% of the carbon fed to the reformer ended up as coke. Additionally, throughout the entire experiment no detectable amounts of contaminants were present in the LFG feed gas after flowing through the purification filters

Table 7 gives the conditions and results of the FTS reaction run using a Co/Si02 eggshell catalyst over the course of the four day LFG run. Note - nitrogen gas from air addition makes up -30% by wt. of the product out of reactor unit causing gas yields to appear inflated and liquid hydrocarbon yields to be deflated.

Table 7:

It was found that the energy content of the fuel gas out of the process has more than enough energy to power the entire plant. This allows for a completely self- sufficient/sustainable process at the commercial level as set out in more detail below.

A detailed fuel analysis was performed on the liquid fuel produced from the LFG run. Figure 1 1 compares the carbon number distribution of the fuel produced directly from the process with that of commercial diesel. Table 8 compares the PIACO analysis results from the LFG run compared to commercial diesel. Table 9 gives the ASTM fuel analysis results per ASTM D975 for "Standard Specification for Diesel Fuel Oils." Hydrocarbon family Process of the invention Comme cal Diesel

(%wt) (%wt)

Paraffins 67.164 19.95

Isomers 28.243 31.6

Olefins 4.323 0.92

Aromatics 0.0 39.48

Cyclics 0.25 8.05

Table 8

Fuel Analysis, ASTM Standard Spec (No. 2 Diesel) Commercial Diesel TRIFTS LFG TRIFTS LFG (Dist 55C)

Specific Gravity, ASTM D4052 (g/cc) 0.8215 0.7386 0.7489

Cetane Index, ASTM D976 > 40 57.6 84.5 72.7

Cetane Index, ASTM D4737 > 40 59.7 92.3 83.4

Flash Point, ASTM D93 (°C) > 52 87 49 57

Cloud Point, ASTM D2500 (°C) -6 -6 -3

Pour Point, ASTM D97 (°C) -9 -9 -6

Distillation, ASTM D86 (°C)

IBP: 0.5wt% 203 143 142

10% 220 164 154

50% 269 234 216

90% 282-338 329 327 314

FBP: 99.5% 378 388 378

Net Heat Comb., ASTM D3338 (MJ/kg) 43.164 44.520 44.355

Table 9

A review of the results of the PIACO hydrocarbon family distribution analysis shows that the liquid fuel has an excellent distribution for the desired hydrocarbon types. The low aromatic content contributes to the exceptional net heat of combustion (ASTM D3338) of the liquid fuel vs. commercial diesel (Table 9). It is important to note that the liquid fuel distilled at 55°C passed all No.2 Diesel fuel specifications and that the boiling point distribution compared very well with typical commercial diesel as the final boiling points (99.5%) were both 378°C. These results show that the liquid fuel produced utilizing actual LFG will be an excellent candidate as a drop-in fuel that could be utilized by today's diesel engines with no modifications to the engine needed. In fact, the high cetane index fuel of the liquid fuel produced in the process will contribute to less knocking within the diesel engine and therefore have a much smoother combustions cycle compared to commercial diesel that could lead to extended engine lifetimes and therefore less operational costs to run the engine. Additionally the low aromatic content of the liquid fuel allows for a more complete combustion and therefore less soot formation that improves the characteristics, as they relate to environmental impact, of the exhaust compared to the combustion of commercial diesel.

In the method and system of the present invention the efficient use of energy and steam recycling plays a critical role in reducing CAPEX and OPEX costs. A major modification is in relation to utilizing the high-energy content fuel gas exiting the FTS reactor. Utilizing the bench scale data and ASPEN generated data, it was determined that a 1 ,500 scfm LFG commercial scale plant would produce a fuel gas out of the FTS having a net heat of combustion of 22.5 MMBTU/hr. Calculation of all the heating requirements of the reactors and equipment to determine a total energy requirement is 34.5 MMBTU/hr for the commercial scale plant. These results can be seen in Table 10. Therefore there is more than enough energy to meet all energy requirements of the full-scale plant by utilizing the fuel gas energy content. The largest energy consumer will be the single tri-reformer reactor that requires 14.4 MMBTU/hr of energy to perform the tri-reforming reaction at the desired temp of 800-850°C. Therefore a large share of the fuel gas will be used to heat the tri-reformer reactor in a direct fire furnace type application. Additional fuel gas can be used to generate any steam in the boiler to meet the steam requirements for the feed gas to the reformer and water/steam jacketed FTS reactor. After balancing all energy requirements with energy generated there is still a net positive energy content in the fuel gas to generate approximately 1.5 MW of electricity. This amount of electricity would be more than enough to power any and all auxiliary equipment the plant would need as well as a surplus to send back to the grid as an additional profit stream. This analysis shows how incorporating the C0 ₂ content of LFG and the hydrogen from the steam into the backbone of the diesel fuel produced, leads to remarkable efficiency and productivity gains as compared to traditional waste to energy projects.

Energy Requirements and Generation from 1500 scfm

LFG Plant

BTU/hr Required BTU/hr Produced

Reformer requires 14,472,000

FTS generates 9,761,143

Fuel Gas Energy Content 22,488,465

Boiler 1,820,786

LFG cooloer 160,414

Reformer HX 7,319,143

Syngas cooler 1,807,500

FTS cooler 3,451,114

FT reactor 9,761,143

Compressor 1 399112.0

Air Compressor 258223.3

Compressor 2 1287299.8

Compressor 3 1068006.3

TOTALS

Equip/RXN Required 34,485,599

Energy Produced ^r 39,568,751

Net Energy Produced 5,083,152

Additional Electric

1.50 MW Power That Can be

Table 10