EA045925B1 - MICROSCALE PLANT FOR DIRECT PRODUCTION OF LIQUID FUELS FROM GASEOUS HYDROCARBON RESOURCES - Google Patents

Description

Technical field

The present invention relates to a transportable, microscale method for the direct, cost-effective production of high-quality liquid fuels from low-volume gas-phase hydrocarbon resources.

State of the art

Gas-to-liquids (GTL) technologies for converting natural gas into liquid fuels have been around for decades. The recent resurgence of interest in converting flare gas, biogas, natural gas and other low-volume gas-phase hydrocarbon resources into liquid fuels is providing significant advances in the rapidly growing field of GTL. Such advances are driven by the need to eliminate and simplify costly conventional processes typically used in existing medium- and large-scale GTL (gas-to-liquid) plants.

Medium- and large-scale plants typically convert approximately 25 (approximately 707,900 cubic meters) to 250 million standard cubic meters into fuel and other products. feet (approximately 7,079,000 cubic meters) per day and more than 250 million standard cubic meters. feet (approximately 7,079,000 cubic meters) per day of gaseous hydrocarbons, respectively. All such medium and large installations use four main processes (A. de Klerk, 2012). These processes include: 1) production of synthesis gas; 2) purification of synthesis gas; 3) catalytic conversion of synthesis gas into hydrocarbon products, with the main product being wax; and 4) converting wax into fuels through complex and expensive refining processes.

In addition to wax, these plants produce by-products consisting of tail gas, liquid hydrocarbons and catalyst reaction water. The composition and concentration of such by-products depend on the composition and purity of the synthesis gas; catalytic reactor designs; catalyst composition and catalyst operating conditions.

Synthesis gas can be produced from many types of carbon-containing resources, including natural gas, coal, biomass, or virtually any carbon-containing feedstock, using thermochemical conversion methods. Such thermochemical conversion processes are generally classified into processes that 1) use oxygen or air, or 2) processes that do not use oxygen or air.

Synthesis gas production processes that use oxygen or air are generally referred to as direct conversion, partial oxidation (POC) or autothermal reforming (ATR) processes. ROC is carried out using substoichiometric hydrocarbon/oxygen gas mixtures in reformers at temperatures ranging from 1500 to 2700°F (about 815 to 1480°C). Praxair, Shell, ConocoPhillips and other companies have developed systems to convert gaseous hydrocarbon resources into synthesis gas using OOX. Each of these systems uses an oxygen supply, which requires the delivery of compressed oxygen to a plant that uses one of the methods described above. For example, the method proposed by Praxair uses a hot oxygen burner, which is non-catalytic and converts natural gas (or other hydrocarbons) and oxygen into synthesis gas, as described in US patent 8,727,767 (5/2014).

The conversion of solid and liquid phase carbon-containing raw materials using steam in the absence of oxygen or air is usually called indirect thermochemical conversion. Steam methane reforming (SMR) is a well-developed process for converting gas-phase hydrocarbons into synthesis gas. Because methane is difficult to steam reform effectively into synthesis gas at temperatures below about 2200°F (about 1200°C), catalysts are typically used to reduce the reforming temperature to about 1600 to 1700°F (about 870 to 930°C). . This process is called catalytic steam reforming and is very effective in reforming other gas-phase hydrocarbons such as C2-C16 hydrocarbons and C1- _C16 hydroxyalkanes and _C3 - _C16 ketones (Sa et al., 2010).

In table Table 1 lists some potential catalyst contaminants found in synthesis gas and the maximum recommended levels for such contaminants. There are numerous methods in the art today for removing hydrogen sulfide, sulfur dioxide, ammonia, hydrogen cyanide, nitrogen oxides, hydrogen chloride and particulates contained in synthesis gas.

- 1 045925

Harmful carboxylic acids can be formed when oxygen reacts with free radical molecules during the catalytic conversion of synthesis gas using CO and H2. When carboxylic acids are formed, they will be approximately distributed between the liquid fuel, catalyst reaction water and wax, as indicated in the table. 2. When such carboxylic acids are present in fuels, the fuel can corrode metal surfaces and the shelf life of the fuel is significantly reduced. Therefore, these acids must be removed (if present) from the fuel before distribution, storage and use, a process that is complex and expensive.

At the same time, when these carboxylic acids are present in the catalyst reaction water, they must be removed before the water can be recycled and used for production processes. In addition to the problem of metal surface corrosion, such acids can damage the catalysts typically used in catalytic steam reforming processes.

Table 2

Relative distribution of carboxylic acids (if formed) in catalyst reaction water, liquid fuels and wax

Carboxylic acid Boiling point (°C) Relative distribution (mole percent) Water Liquid fuels Wax Methane (ant) 101 100 0 0 Ethanic (acetic) 118 100 0 0 Propane 141 75 25 0 Butane 164 thirty 70 0 Pentanic 187 10 85 5 Hexane 205 5 80 15 Octane 239 <1 75 25

In the current art, there are many methods for purifying synthesis gas before catalytically converting the synthesis gas into hydrocarbon products. Thermochemical conversion of gas-phase hydrocarbons produces much lower concentrations of pollutants in the synthesis gas than the conversion of solid carbon-containing materials such as biomass, coal, municipal solid waste and other solids. Sulfur compounds are the most common contaminants in many gas-phase hydrocarbon resources. Such contaminants can be easily removed using a variety of solid phase binders such as iron oxide or zinc oxide.

There are two main approaches for the catalytic conversion of synthesis gas into fuels: 1) catalytic conversion of synthesis gas into intermediate products (mainly wax), followed by expensive wax enrichment and refining processes such as hydrocracking; 2) direct catalytic conversion of synthesis gas into fuels, which produces a minimal amount of wax [US patents 8394862 (8/2013) and 9090831 (7/2015)].

In all modern medium and large scale GTL plants, synthesis gas is converted to wax as the main product. Then to obtain fuels and other products from wax

- 2 045925 use refining/enrichment methods. Because such refining processes are complex and expensive, fuel production costs can increase by approximately 40% or more relative to direct production methods. Medium- and large-scale plant designs that incorporate traditional Fischer-Tropsch processes and utilize wax hydrocracking and other expensive upgrading processes are not economically feasible for distributed plants that process relatively small volumes of gas-phase hydrocarbons.

Micro-GTL plants include processes capable of converting approximately 0.05 to 2.5 million standard cubic meters. ft/d (about 1,415 to about 70,800 cu m/d) of gaseous hydrocarbons in about 5 to 250 bbl/d (about 0.79 to about 40 cu m/d) liquid fuels. GTL plants, which convert approximately 2.5 to 50 million standard cubic meters. ft/d (about 70,800 to about 1,415,800 cu m/d) of gas-phase hydrocarbons in about 250 to 5,000 bbl/d (about 40 to about 800 cu m/d) fuel, commonly referred to as small GTL units.

There are several types of catalytic reactors that have been introduced on an industrial scale to catalytically convert synthesis gas into hydrocarbon products. Multitubular fixed bed catalytic reactors consist of many small diameter tubes that are used to house the catalyst. These tubes are placed inside the reactor vessel, in which water circulates to remove the exothermic heat generated when the synthesis gas is converted into hydrocarbon products. The use of catalysts that form heavy waxes can result in coating of the catalyst, resulting in reduced catalyst activity. Such reactors operate in a multi-pass mode, removing products after each pass and recycling unreacted synthesis gas back to the catalytic reactors. Typically, two to three passes through these reactors convert approximately 90 vol.% CO into hydrocarbon products.

Slurry reactors use finely divided catalysts suspended in a liquid medium. Heat is removed using internal cooling coils. Synthetic gas is bubbled through a liquid medium, which also ensures mixing of the reactor contents. The small size of the catalyst particles improves heat mass transfer to the liquid medium. Waxy products must be separated from the catalyst particles.

Microchannel reactors consist of reactor cores containing thousands of thin process channels filled with very small particle size catalysts. Such reactor cores alternate with channels ranging in size from 0.1 to 10 mm containing a water coolant. Because the catalyst particles and channels are small, heat can be dissipated more quickly than with traditional 25 to 40 mm tubular reactors.

Many catalysts and catalytic processes have been developed and implemented to convert synthesis gas into wax. Such catalysts are usually called Fischer-Tropsch (F-T) catalysts (Jahangiri et al., 2014). The proposed methods describe catalysts that produce high molecular weight hydrocarbon reaction products (eg, waxes) that require further processing, including hydrotreating and other enrichment processes, to produce diesel fuel or a diesel fuel blending component.

In general, using a two- or three-phase separator, the product stream from the catalytic reactors is separated into the following fractions: tail gas; condensed liquid hydrocarbons, catalyst reaction water and waxes. The tail gas fraction typically consists of _H2 , CO, CO2 and C1-C5 hydrocarbons and oxygenated organic compounds; the condensed fraction contains C ₅ -C ₂₄ hydrocarbons and oxygenated organic compounds; the wax fraction contains C ₂ 3-C1 ₀₀₊ hydrocarbons; and the catalyst reaction water fraction consists of water containing up to about 5.0 vol.% dissolved oxygenated organic compounds.

Since the synthesis gas catalytic conversion efficiency is typically approximately 90% or higher when using tail gas recirculation tubular reactors, some H2 and CO will remain in the tail gas. In addition, the tail gas contains some CH4, which is formed as a result of the catalytic reaction. The composition of the tail gas depends on the type of thermochemical process, the catalyst used and the operating conditions. Such tail gas can be recycled back to the thermochemical conversion system to produce additional synthesis gas and/or it can be used as burner fuel. Almost all of the described catalytic processes have been used to convert synthesis gas, mainly into wax.

Brief description of the invention

The advantages of such a microscale GTL method are summarized below. Only three main modular processes are necessary, including: 1) synthesis gas purification unit and synthesis gas generator unit; 2) catalytic reactors and product separation/collection unit; and 3) a utility package, including recirculation pumps, process control systems, and utility networks. Such a microscale GTL plant can be economically

- 3 045925 quickly and efficiently convert gas-phase hydrocarbons in quantities of approximately 0.05 to 2.5 million standard cubic meters. feet (from about 1,415 to about 70,800 cubic meters/day) per day and even smaller volumes of gas ranging from 0.00 to 0.05 million standard cubic meters. feet (0.00 to approximately 1415 cubic meters) per day, and higher gas volumes ranging from approximately 2.5 million standard cubic meters. feet (approximately 70,800 cubic meters) per day up to 10 million standard cubic meters. feet (approximately 283,200 cubic meters) per day, directly into high-quality liquid fuels.

The catalytic reactor operates at moderate temperatures from 350 to 450°F (about 177 to about 232°C), preferably from 400 to 425°F (from about 204 to about 218°C) and hourly gas space velocities from 100 to 10,000, preferably from 1000 to 2500. Since such a reactor uses highly efficient steam cooling to quickly remove heat from the exothermic catalytic reaction, conversion efficiency of H2 and CO into products of greater than about 70 mol% can be achieved in a single pass through multiple catalytic reactors in series. preferably greater than about 80 mol%, or more preferably greater than 90 mol%. This reduces or eliminates costly recompression and recycling of catalyst tail gas back to the catalytic reactors. The catalyst tail gas is recycled back to the synthesis gas generation step to produce additional synthesis gas, used as burner fuel, used for power generation, or various combinations of these recycling methods.

Such a catalyst produces a very small amount of wax (<25 vol.%, preferably <5 vol.% and more preferably <2 vol.% with an average H ₂ /CO ratio of 2.06), with the resulting wax remaining in the catalytic reactor as a liquid at operating temperatures ranging from 400 to 425°F (about 204 to about 218°C) because it is a light wax consisting primarily of C ₂₂ -C ₃₅ hydrocarbons. This liquefied wax passes through the catalytic reactors and is either removed from the bottom of the reactors or separated in a drum without interfering with the flow of synthesis gas. Because very little wax is produced, complex and expensive refining processes are not required to convert wax into fuels.

In one case, the operating conditions of the catalytic reactor may be changed (which may include changing the H ₂ :CO ratio, gas hourly space velocity, pressure or temperature), and depending on the specific operating conditions, the resulting distribution of hydrocarbons in the fuel product is changed so that optimize fuel products for sale in specific markets. For example, reducing the H ₂ :CO ratio below 2.0:1.0 will shift the hydrocarbon distribution toward higher molecular weights, which may be desirable in some market sectors and applications. Conversely, the H ₂ :CO ratio can be increased to greater than 2.0:1.0 to shift the hydrocarbon distribution toward lower molecular weights.

Since the catalyst reaction water contains undetectable levels of corrosive and harmful carboxylic acids of less than about 100 ppm, or preferably less than 25 ppm, or more preferably less than 15 ppm, such catalyst water can generally be directly recycled and used in the steam reformer. As a result, the unit requires little or no water from external sources to operate. The high-quality fuel consists of _C5 - _C24 hydrocarbons that can be used directly in off-highway diesel engines, blended at approximately 20 vol.% with conventional petroleum diesel fuel, or lightly distilled into a variety of fuel products, depending on local market requirements.

Brief description of graphic materials

In fig. Figure 1 illustrates the basic processes for a microscale GTL system.

In fig. 2 shows details of a catalytic reactor 1 containing a direct fuel production catalyst.

In fig. Figure 3 illustrates the types of fuel products obtained from the distillation of liquid fuels obtained directly from a microscale GTL system.

Detailed Description of the Invention

The first aspect of the microscale process of the present invention is the introduction of a direct fuel production catalyst 109b designed to directly produce high quality liquid fuels and catalyst reaction water containing undetectable or barely detectable harmful carboxylic acids. In this context, undetectable is defined as values that are at or below the detection limit, with the stated detection limit being the lowest concentration that can be reliably detected but that lies below a level that is quantifiable with acceptable accuracy. The limit of quantitation is the lowest concentration that can not only be detected, but also quantified with a certain degree of accuracy. The limit of quantitation is always higher than the detection limit, usually by a factor of three or four. For example

- 4 045925 measures, the detection limit of the GC/MS (gas chromatography/mass spectrometry) method used for the quantitative determination of oxygenated hydrocarbons listed in table. 4 was 25 ppm, and the limit of quantitation was from 75 to 100 ppm. Barely detectable is defined as values between the limit of detection and the limit of quantitation. Therefore, undetectable or barely detectable is defined as <25-100 ppm.

When the direct fuel catalyst described in US Pat. Nos. 8,394,862 (8/2013) and 9,090,831 (7/2015) is produced using a support whose surface pH is close to neutral (e.g., pH about 7.0), and when the oxygen concentration in synthesis gas is less than 5000 ppm, preferably less than 1000 ppm and more preferably less than 500 ppm, specific carboxylic acids are not detected in the catalyst reaction water and liquid fuel fractions at levels above about 100 ppm, preferably not above about 25 ppm and more preferably not above approximately 15 ppm.

Catalyst characteristics include the following.

The catalyst form is ideally an extrudate with a petal, corrugated, or paddle cross-section, but may also be a sphere, granule, powder, or other carrier form that provides effective performance. For example, the use of a petal design can significantly increase the area to volume ratio of the catalytic reactor, thereby improving the volumetric efficiency of the catalytic reactor system. Lobe designs also provide improved pressure drop, resulting in lower pressure differences upstream and downstream of the catalyst bed, especially when used in fixed-bed reactors.

The effective radius of a catalyst or support pellet refers to the maximum radius, which is the distance from the midpoint of the support to the surface of the support. For petal media, the effective tablet radius refers to the minimum distance between the midpoint and the outer surface of the tablet. In some cases, the effective tablet radius is approximately 600 microns or less. In other cases, the effective tablet radius is about 300 microns or less.

The catalyst tablet or carrier material may be porous. In some cases, the average diameter of the surface pores of the support material is greater than about 100 angstroms. In other cases, the specified diameter is more than 80 angstroms.

Any suitable material that can have a neutral surface (eg, pH about 7.0) can be used as a catalyst support material in the Fischer-Tropsch process. Non-limiting examples of supports include metal oxides such as alumina, silica, zirconia, magnesium oxide, or combinations of these metal oxides. Aluminum oxide is often used as a carrier.

Catalytically active metals that are included in or dispersed in the support material include substances that promote the production of diesel fuel in the catalytic reaction. For example, such metals include cobalt, iron, nickel, or any combination thereof. Various promoters can also be added to the support material. Examples of promoters include cerium, ruthenium, lanthanum, platinum, rhenium, gold, nickel and rhodium.

In some cases, the catalyst support has a crush strength of from about 1 lb/mm (about 0.45 kg/mm) to about 10 lb/mm (about 4.5 kg/mm) and a BET specific surface area greater than about 75 m ² /G. In other cases, the catalyst has a crush strength of from about 2 lb/mm (about 0.91 kg/mm) to about 5 lb/mm (about 2.3 kg/mm) and a BET specific surface area of greater than about 100 m ² / G. In still other cases, the catalyst has a crush strength of from about 3 lb/mm (about 1.4 kg/mm) to 4 lb/mm (about 1.8 kg/mm). In still other cases, the catalyst has a BET specific surface area greater than about 125 m ² /g or 150 m ² /g.

The distribution of active metal on the support is typically from about 2% to about 10%. Often the active metal distribution is from about 3% to about 5% (eg, about 4%). The active metal distribution represents the fraction of atoms on the surface of the catalyst that are exposed, which is expressed as follows: D=NS/NT, where D represents the dispersion, NS represents the number of surface atoms, and NT represents the total number of atoms of the material. Dispersion increases with decreasing crystallite size.

In some cases, the supported catalyst contains cobalt, iron, nickel, or combinations thereof, supported in an amount of from about 5 wt.% to 30 wt.%, on gamma alumina having a neutral surface, more typically supported in an amount of about 20 wt. .%, per gamma-alumina having a neutral surface, based on the total weight of the supported catalyst. In addition, in these cases, the supported catalyst composition includes selected combinations of one or more promoters consisting of ruthenium, palladium, platinum, gold, nickel, rhenium, and combinations thereof, in an amount of from about 0.01 to 20.0 wt.%, more typically from about 0.1 to 0.5 wt.%, at

- 5 045925 promoter. Methods for preparing the catalyst include impregnation and other preparation methods commonly used in industry and described in the art.

Liquid fuels produced directly by use of the catalyst of the present invention have very little or no corrosive properties and undergo little or no oxidation or degradation during storage and can be stored for years without change. In addition, the catalyst reaction water can typically be directly recycled 112 to the synthesis gas generator 103.

When using the catalyst of the present invention to convert synthesis gas directly into fuels, the catalyst reaction water contains very little or undetectable levels of carboxylic acids. In table 4 summarizes the content of hydroxyalkanes (eg, alcohols) and carboxylic acids in the catalyst reaction water obtained from the catalysis of synthesis gas that was produced by steam reforming of natural gas, natural gas liquids and glycerol using the catalyst of the present invention.

Although the total concentration of hydroxyalkanes was found to be 12831 ppm, 16560 ppm and 18877 ppm, respectively, in the case of synthesis gas obtained from these three feedstocks, acetic acid, propionic acid and malonic acid were not detected in the catalyst reaction water samples ( each content was less than about 25 ppm).

Since hydroxyalkanes and carboxylic acids will partition between the catalyst reaction water and the fuel, the possible presence of carboxylic acids in liquid fuels can be easily assessed using the ASTM D130 copper strip corrosion test. In the presence of carboxylic acids, the surface of the copper strip changes color, the change of which is characterized by indicators from 1a to 4c, with the designation 1a indicating very little or no corrosion, while at 4c the fuel causes corrosion of the copper strip with the appearance of a dark brown/black color, indicating that the fuel in question contains unacceptable levels of carboxylic acids (ASTM International, 2012). Fuels produced directly from these feedstocks were found to provide a test result consistent with 1a, which confirmed that carboxylic acids were undetectable.

Table 4

Concentration of oxygenated organic compounds in catalyst reaction water obtained from synthesis gas obtained from various gas-phase hydrocarbons by using a direct fuel catalyst

Oxygenated organic compound Gas-phase hydrocarbon resource Natural gas Evaporated gas condensate liquids Evaporated glycerin Concentration (ppm) in catalyst reaction water ¹ Methanol 4470 4980 6177 Ethanol 4890 5040 6529 1-Propanol 1970 1930 2209 1 -Butanol 1980 2530 1888 1-Pentanol 1080 1380 1342 1-Hexanol 310 290 333 1-Geptanol 111 60 67 1-Octanol <25 <25 122 1 -Nonanol <25 <25 <25 Acetic acid <25 <25 <25 Propionic acid <25 <25 <25 Malonic acid <25 <25 <25 Total 12,831 16,560 18,877

¹ⁿ Non-detectable is defined as values that are at or below the detection limit of 25 ppm, with the detection limit being the lowest concentration that can be reliably detected but which is below a level that can be quantified with acceptable accuracy.

The second aspect of the microscale method according to the present invention is

- 6 045925 direct recycling of catalyst reaction water, since such water contains undetectable or barely detectable harmful carboxylic acids. If the syngas production processes do not require much steam, such as a reciprocating engine syngas reformer, such catalyst reaction water can be used to recover additional oil from a nearby oil well.

When oil is present in underground rock formations such as sandstone, carbonate or shale, such oil can generally be developed by drilling a wellbore into the oil-bearing formation, allowing existing pressure gradients to force the oil up the wellbore. This method is known as the primary oil recovery method. If and when pressure gradients are not sufficient to produce oil at the required rate, an enhanced recovery method is typically used to recover additional oil. This method is known as the secondary oil recovery method.

Even after secondary oil recovery using water injection, large quantities of the original oil may remain in the reservoir. The proportion of unrecoverable hydrocarbon is usually highest in heavy oils, shale oils and complex geological formations. In some large oil fields, more than a billion barrels (approximately 0.16 billion cubic meters) of oil can remain after normal water injection.

Tertiary oil recovery methods then come into focus. It is estimated that modern tertiary oil recovery techniques can remove an additional 5 to 20 percent of the oil remaining in the reservoir. Developing effective tertiary oil recovery strategies to more efficiently recover oil promises to have a significant economic impact. Current tertiary oil recovery methods are effective but expensive because many oil-producing areas have limited water supplies.

Hydroxyalkanes (alcohols) consisting of one to four carbon atoms dissolved in water have been found to be ideal for tertiary oil recovery [US Patent 7,559,372 (7/2009)]. However, adding mixed alcohols to water for tertiary oil recovery is a very expensive procedure. Because such catalyst reaction water obtained from the catalyst described herein contains up to about 2.0 vol.% C ₁ -C ₅ alcohols, it is ideal for direct use in tertiary oil recovery.

Microscale GTL systems typically refer to plants that convert less than approximately 2.5 million standard cubic meters. feet (approximately 70,800 cubic meters) per day (MMScf/day) of gaseous hydrocarbons into liquid hydrocarbon products. Medium and large scale GTL systems cannot be economically scaled to process less than approximately 2.5 million standard cubic meters. feet (approximately 70,800 cubic meters) per day of gas-phase hydrocarbons without eliminating the main typical processes and simplifying other processes.

In fig. 1 illustrates aspects of typical processes for a microscale GTL process in accordance with the present invention. The method includes: a series of catalytic reactors 108 (FIG. 2) that eliminate or reduce the need for tail gas recompression and recirculation; a direct fuel catalyst 109 providing undetectable levels of individual carboxylic acids in the liquid fuel and catalyst reaction water of less than or equal to about 100 ppm, or preferably less than 25 ppm, or more preferably less than 15 ppm; a process for directly recycling catalyst reaction water 112 for use as steam in the synthesis gas generator 103 and/or for use in tertiary oil recovery 119; introducing liquid fuel directly into crude oil pipeline 117; direct on-site fuel use for off-road diesel engines used in generators, compressors, pumps, tractors, etc. 118; or transporting liquid fuels to another location to produce high quality liquid fuel products 116 (FIG. 3) that meet or exceed ASTM and/or other equivalent fuel standards.

In fig. 3 shows details of the catalytic reactor used in the typical process 108 (FIG. 2), which contains a series of four reactors 202-205 designed to efficiently convert purified and compressed synthesis gas 201 into high-quality fuel 207 in a single-pass manner.

The method of the present invention results in the elimination of several unnecessary typical processes and the simplification of other processes, which provides significant reductions in capital, operating costs, system portability and ease of operation for micro-scale direct production of liquid fuels from gas-phase hydrocarbon resources.

The following description is provided to enable any person skilled in the art to make and use the present invention, while the description is presented in the context of the particular application and its requirements. Various modifications to the described embodiments will be apparent to those skilled in the art, and the general principles defined herein may be applied to other embodiments and applications.

- 7 045925 without deviating from the essence and scope of the present invention. Thus, the present invention is not limited to the embodiment shown, but is intended to be within the broadest scope consistent with the principles and features described herein.

The term contains and its grammatical equivalents are used herein to denote the optional presence of other components, ingredients, steps, etc. For example, the resulting product containing components A, B and C may contain only components A, B and C, or may contain not only components A, B and C, but also one or more other components.

In fig. 1 shows the main typical processes for one of the embodiments of the present invention. Gas-phase hydrocarbon resources and/or vaporized liquid hydrocarbon resources are introduced into system 101. These resources may be natural gas; biogas; associated gas; gas-phase hydrocarbons (for example, C ₂ -C ₄ ); a mixture of a wide fraction of hydrocarbons or a mixture of gas condensate liquids (NGL); individual components extracted from natural gas streams, such as ethane, propane, butane or other components, natural gas condensates (C ₅₊ ); or other similar gases or liquids (such as naphtha or condensate) that can readily evaporate to form gas. Any harmful contaminants, such as sulfur compounds, are removed 102 from the gas-phase hydrocarbons before being fed into the synthesis gas generator 103.

The first synthesis gas generator 103 uses steam to facilitate the conversion of hydrocarbons into synthesis gas. This steam is produced primarily through direct recirculation and the use of catalyst reaction water 112, and in some cases make-up water may be required to control the steam to carbon ratio to eliminate the formation of elemental carbon. Synthesis gas generator 103 utilizes a catalyst that efficiently converts methane and other hydrocarbons into synthesis gas at operating temperatures below about 1700°F (about 930°C). Syngas post-treatment unit 105 may be used to remove any contaminants that may still be present in the synthesis gas. Since the H ₂ /CO ratio may become too high during the catalytic steam reforming process, membrane system 106 is used to bring the ratio to about 2.2. Compression 107 may be required to increase the synthesis gas pressure to about 350 psi (about 2413 kPa).

An alternative synthesis gas generator 103 may be used that uses a modified piston engine to produce synthesis gas [US Patent Application Publication 2014/0144397 (5/2014)]. A membrane separator (Lin et al., 2013) or a vapor pressure swing adsorption (VPS A) system 104 [Praxair, 2016] is used to enrich the air with O2 from approximately 30% to 93%, respectively, for introduction with gas-phase hydrocarbons into the generator 103 synthesis gas with piston engine.

Contaminants contained in the synthesis gas that may be generated by the reciprocating engine reformer synthesis gas generator 103, such as ammonia, HCN, and particulate matter, are removed by methods commonly used in the art 105. Because the H ₂ /CO ratio is in an almost ideal range of 1.6 to 2.0, no adjustment of this ratio 106 is required. Compression 107 may be required to increase the synthesis gas pressure to approximately 350 psi (approximately 2413 kPa).

In fig. 2 shows design details of a catalytic reactor 200 containing a series of multiple reactors 202-205 for efficient, single-pass direct conversion of synthesis gas into high quality fuel (4 reactors in series are shown in FIG. 2). This configuration eliminates the need for recompression and tail gas recirculation, which are typically used in the current art. A series of multiple horizontal or vertical catalytic reactors or reactors positioned at an angle of 200 relative to a horizontal and vertical orientation are designed to efficiently convert into products about 70 mol%, preferably 80 mol%, or more preferably 90 mol% of H2 and CO contained in synthesis gas, without re-compression and recirculation of synthesis gas. In preferred embodiments, two to five reactors (e.g., 2 reactors, 3 reactors, 4 reactors, or 5 reactors) are used, or six to ten reactors are used (e.g., 6 reactors, 7 reactors, 8 reactors, 9 reactors) to achieve the desired results. or ten reactors).

Such multi-tubular catalytic reactors 202-205 effectively remove heat from the exothermic catalytic reaction. They operate at temperatures from about 350°F to 450°F (about 177°C to about 232°C), but preferably from 400 to 430°F (about 204°C to about 221°C); at a pressure of from about 250 to 450 psig (from about 1724 to about 3103 kPag); and gas space velocities from about 100 to 10,000 h ^-1 , but preferably from 500 to 3,000 h ^-1 .

Purified synthesis gas 201 (FIG. 2) is introduced into a catalytic reactor 1 202 operating at temperatures, pressures and space velocities necessary to convert approximately 35% of the CO and H2 to be converted into products. The remaining CO and H ₂ and products from the catalytic reactor 1 202 are introduced into an elongated catalytic reactor 2203 operating at temperatures, pressures and volume

- 8 045925 speeds required to convert approximately 50% of the remaining CO and H2 into additional products.

The remaining CO and H2 and products from catalytic reactor 2203 are introduced into an even longer catalytic reactor 3204, which is operated at the required temperatures, pressures and space velocities to convert approximately 65% of the remaining CO and H2 into additional products.

However, in some embodiments, these reactors may have different numbers of tubes, or in other embodiments, they are the same size and/or have the same number of tubes. The unique arrangement and design of these catalytic reactors provide conversion efficiencies of approximately 90% or greater for converting CO and H2 into liquid fuels, without the need to separate liquid fuels and catalyst reaction water from the catalyst tail gas, and to recompress and recycle the products typically used in modern field of technology.

Direct fuel catalyst 109 is prepared as described by Schuetzle in US Pat. No. 9,090,831 (7/2015), except that a catalyst support having nearly neutral surface properties (defined as a surface that is neither acidic nor basic) is used.

When using a catalyst support whose surface is neither acidic nor basic, undetectable harmful carboxylic acids are formed in the fuel and catalyst reaction water.

As a result, directly charged liquid fuels are non-corrosive or have very little or no corrosive properties, undergo very little or no oxidation or degradation during storage, and can be stored for years without change. In addition, in general, the catalyst reaction water can be directly recycled 112 to a synthesis gas generator such as a steam reformer 103 without any problem.

In some cases where the synthesis gas generator requires little or no steam, such as a reciprocating engine synthesis gas generator, catalyst reaction water can be injected into oil wells for secondary or tertiary oil recovery 119.

Liquid fuels are separated using a three-phase separator 110 into tail gas 111 (C1- _C4 hydrocarbons, oxygenated hydrocarbons, _CO2 and unreacted H2 and CO), catalyst reaction water 112 and liquid fuels 113 (mainly consisting of _C5 - _C24 hydrocarbons and oxygenated organic compounds). In addition, a small amount of wax (mainly consisting of C ₂₄ -C ₄₀ hydrocarbons) is formed. According to one embodiment of the present invention, less than 25 wt.% of wax by weight is formed relative to the total production output and; in a preferred embodiment, less than 5 wt.% wax is formed, and in a more preferred embodiment, less than 2 wt.% wax is formed.

In some embodiments of the present invention, to optimize economics for application in a particular market, the operating conditions of the catalytic reactor may be varied to change the final product line. Operating conditions in the catalytic reactor that affect the product line include the H ₂ :CO ratio, gas hourly space velocity, pressure and/or temperature. For example, decreasing the H ₂ :CO ratio below 2.0:1 will shift output toward heavier hydrocarbon distribution, which may be desirable in some market sectors and applications. In some embodiments of the present invention, tail gas 111 may be recycled to thermochemical synthesis gas generator 103, where said tail gas may be converted into additional synthesis gas or used as burner fuel.

An optional aspect of the present invention is to directly recycle the catalyst reaction water 112 to the steam reformer or, when using a synthesis gas generator requiring little or no steam, the catalyst reaction water can be used directly for tertiary oil recovery. This innovation is made possible because the catalyst reaction water contains undetectable or barely detectable levels of harmful carboxylic acids.

Fuel oil 113 can be used directly and locally in non-road engines used in diesel generators, tractors, compressors, water pumps, agricultural machinery, construction equipment, etc.

Liquid fuel 113 may be collected and transported by road and/or rail to a centralized location where it is distilled 115 to produce the high quality fuel products 116 shown in FIG. 3, for distribution in local fuel markets.

Possible products obtained by distillation of liquid fuel according to FIG. 3 include: reformulated gasoline blending components (approximately C5-C8 hydrocarbons and oxygenated organic compounds) 303; diesel fuel No. 1 (kerosene) (approximately C ₅ -Ci ₆ hydrocarbons and oxygenated organic compounds) 304; diesel fuel No. 2 305 (approx.

- 9 045925 specifically C ₉ -C ₂₀ hydrocarbons and oxygenated organic compounds); diesel fuel No. 3 306 (approximately _C16 -C23 hydrocarbons and oxygenated organic compounds); and a small fraction of 307 wax (C22+ hydrocarbons and oxygenated organic compounds). A small amount of gases (C ₂ -C ₅ ) 302 is obtained, as well as a small residue (mainly oxidized hydrocarbons) 308.

Alternative or complementary processes can be used to further distill liquid hydrocarbons to separate high-value alpha-olefins and hydroxyalkanes from liquid fuels. An alternative embodiment involves directly introducing liquid fuels into the combined crude oil pipeline 117 at the wellhead, wherein the fuels are mixed with the crude oil for transportation to a refinery and/or chemical processing plant. Because liquid fuels have a much lower density and viscosity than crude oil, they serve to improve the flow of oil through pipelines.

The methanol catalysts can be used in the catalytic reactor in conjunction with the catalytic reactor 108 to produce intermediate methanol feedstocks that can be transported to a refinery and/or chemical plant for further processing into fuels and/or chemical compounds. Such a plant can also be integrated with an ammonia plant, whereby excess hydrogen from the GTL plant is used as feedstock in the ammonia plant, thereby optimizing the economics of production of said product.

In some cases, to prevent coking and other undesirable reactions in some synthesis gas generators 103, the water to carbon ratio of the feedstock is adjusted in the range of 1.5-3.0/1.0 and preferably 2.0-3.0/1.0 to prevent coking and other undesirable reforming reactions. Although some make-up water may be necessary when running the integrated process described in FIG. 1, typically, catalyst reaction water will be sufficient to maintain an efficient catalytic steam reforming process after startup without the need for makeup water.

In some cases, the methods of the present invention include the following processes.

A suitable hydrocarbon resource and/or vaporized liquid hydrocarbon resource is introduced into a synthesis gas generator containing a catalyst. In some cases, the feed material is a natural gas stream or a separate component thereof. Undesirable contaminants are removed from the gas-phase hydrocarbons before introducing the synthesis gas into the generator. One type of synthesis gas generator uses steam to facilitate the conversion of hydrocarbons into synthesis gas. Another type of synthesis gas generator uses a modified piston engine to produce synthesis gas.

Where steam assists in the conversion of hydrocarbons, it is generally produced by direct recycling of water produced in the direct production process of the present invention. Typically, direct recirculation results in more than 50 weight percent steam. Often, direct recirculation results in at least 60 weight percent, at least 70 weight percent, at least 80 weight percent, or at least 90 weight percent steam.

The catalyst used in the synthesis gas generator converts hydrocarbons into synthesis gas at temperatures below about 1700°F (about 930°C). An optional post-treatment unit is used to remove any contaminants that may be present in the synthesis gas. If the H2/CO ratio in the synthesis gas is too high, a membrane separation system can be used to control the ratio. In some cases, compression is used to increase the pressure of the synthesis gas, such as pressures ranging from about 300 psi (about 2068 kPa) to about 500 psi (about 3447 kPa), from about 300 psi (about 2068 kPa) to about 450 psi ( about 3103 kPa), from about 300 psi (about 2068 kPa) to about 400 psi (about 2758 kPa), from about 325 psi (about 2241 kPa) to about 475 psi (about 3275). The oxygen concentration in the synthesis gas is typically less than 1000 ppm. Often the oxygen concentration is less than 750, 500 or 250 ppm.

The resulting synthesis gas is introduced into multi-tubular catalytic reactors, typically 2, 3, 4 or 5 reactors connected in series, arranged either horizontally, vertically, or at an angle between horizontal and vertical orientations. Catalytic reactors typically operate under the following gas temperature, pressure, and space velocity conditions: from about 400 to about 430°F (about 204 to about 221°C), for example, from about 400 to about 415°F (from about 204°C to about 213°C) or from about 415 to about 430°F (from about 204°C to about 221°C); from about 250 to about 450 psig (from about 1724 to about 3103 kPag), for example from about 250 to about 300 psig (from about 1724 to about 2068 kPag), from about 300 psig to about 350 psig (from from about 2068 to about 2413 kPag), from about 350 psig to about 400 psig (from about 2413 to about 2758 kPag), from at

- 10 045925 approximately 400 psig to approximately 450 psig (from approximately 2758 to approximately 3103 kPa g); at an hourly space velocity of from about 500 to about 3000, for example from about 500 to about 750, from about 750 to about 1000, from about 1250 to about 1500, from about 1500 to about 1750, from about 1750 to about 2000, from about 2000 to about 2250, from about 2250 to about 2500, from about 2500 to about 2750, from about 2750 to about 3000.

Each catalytic reactor contains a catalyst, typically a supported catalyst. The catalyst support has a number of physical and chemical properties, such as material type, shape, pellet radius, average pore diameter, crush strength, pore volume, and surface area. The listed properties typically fall within a certain range of magnitude or distinctiveness as follows: materials including metal oxides (e.g., alumina, silica, zirconia, magnesium, or combinations thereof), zeolites, and carbon nanotubes having approximately neutral surface pH (eg, a pH in the range of 6.5 to 7.5, a pH in the range of 6.75 to 7.25, and preferably a pH of about 7.0); shapes, including petal (for example, three, four or five petals of equal or unequal length), corrugated or lobed cross-sections, spherical shape, granule and powder; tablet radius less than about 600 μm, less than 500 μm, less than 400 μm, or less than 300 μm; an average pore diameter greater than about 75 angstroms, greater than about 100 angstroms, greater than 110 angstroms, or greater than 120 angstroms; crushing strength from about 1 lb/mm (about 0.45 kg/mm) to about 10 lb/in (about 4.5 kg/mm), for example, from about 1 lb/mm (about 0.45 kg/mm ) to about 7.5 lb/mm (about 3.4 kg/mm), from about 1.5 lb/mm (about 0.68 kg/mm) to about 6.0 lb/mm (about 2.7 kg) /mm), from about 2.0 lb/mm (about 0.91 kg/mm) to about 5.5 lb/mm (about 2.5 kg/mm), from about 2.5 lb/mm (about 1 .1 kg/mm) to about 5.0 lb/mm (about 2.3 kg/mm) or about 3.0 lb/mm (about 1.4 kg/mm) to about 4.5 lb/mm ( approximately 2.0 kg/mm); BET specific surface area greater than about 75 m ² /g, greater than about 100 m ² /g, greater than 125 m ² /g, or greater than 150 m ² /g.

The supported catalyst has a number of physical and chemical properties, such as active metal type, active metal dispersion, weight percent of supported catalyst, promoter type, and promoter weight percent. Such properties typically fall within a certain range of magnitude or signature as follows: active metal, including cobalt, iron, nickel, or combinations thereof; dispersion from about 2 percent to about 10 percent, for example, from about 2 percent to about 4 percent, from about 4 percent to about 6 percent, from about 6 percent to about 8 percent, from about 8 percent to about 10 percent; a weight percent of supported catalyst from about 5 weight percent to about 30 weight percent, from 10 weight percent to about 25 weight percent, from 15 weight percent to about 25 weight percent, or from 17.5 weight percent to about 22.5 weight percent; a promoter, including ruthenium, palladium, platinum, gold, nickel, rhenium, iridium, silver, osmium, or combinations thereof; promoter weight percent from about 0.01 weight percent to about 2.0 weight percent, for example from about 0.05 weight percent to about 1.75 weight percent, from about 0.075 weight percent to about 1.50 weight percent, from about 0.1 weight percent to about 1.25 weight percent, about 0.1 weight percent to about 1.0 weight percent, 0.1 weight percent to about 0.75 weight percent, or about 0.1 weight percent to approximately 0.5 mass percent.

In the presence of three multi-tubular catalytic reactors connected in series, purified synthesis gas is introduced into the first catalytic reactor. It operates at temperatures, pressures and volumetric velocities at which more than 25% of CO and H ₂ are converted into products. In some cases, more than 30% is converted; in other cases, approximately 35% is converted. The remaining CO and H ₂ and products from reactor 1 are introduced into reactor 2, which is operated under conditions such that more than 35% of the feed material is converted into products. In some cases, more than 40%, more than 45%, or approximately 50% are converted. The remaining CO, H ₂ and products are introduced into reactor 3, operating under conditions under which more than 50% of the supplied material is converted. In some cases, more than 55%, more than 60%, or approximately 65% are converted. Series-connected multi-tubular catalytic reactors provide at least 80% efficiency in converting CO and H ₂ into liquid fuels (eg, C ₅ -C ₂₄ N-alkanes). In some cases, such reactors provide at least 85% conversion efficiency, at least 90% conversion efficiency, at least 92.5% conversion efficiency, or at least 95% conversion efficiency.

This efficiency is achieved without recompression and/or recirculation of the catalyst tail gas.

The output (product) of series-connected multi-tubular catalytic reactors typically includes liquid fuel and water. Liquid fuels and water typically each contain less than 500 ppm of the sum of acetic acid, propionic acid and malonic acid (combined acids). In some cases, the liquid fuel and water each contain less than 250 ppm combined acids, less than 200 ppm combined acids, less than 150 ppm combined acids, less than 100 ppm combined acids, less than 50 ppm combined acids, or less than 25 ppm combined acids

- 11 045925 acids.

In other cases, the methods of the present invention include the following processes.

1) To facilitate the conversion of hydrocarbons into gas, a synthesis gas generator uses steam. Direct recycling of the water produced by the production process of the present invention results in greater than 60 weight percent, or 70 weight percent, or 80 weight percent steam. The oxygen concentration in the synthesis gas is less than 1000 ppm or less than 500 ppm. Three multi-tubular catalytic reactors are connected in series horizontally. Operating conditions: 400 to about 430°F (about 204 to about 221°C); from about 250 psig (about 1724 kPa g) to about 350 psig (about 2413 kPa g) or from about 350 psig (about 2413 kPa g) to about 450 psig (about 3103 kPa g); space velocity from about 500 to about 1750 or from about 1750 to about 3000. A supported catalyst, wherein said support has an approximately neutral pH and wherein the support has the following physical and chemical properties: alumina as a material; petal shape; tablet radius less than about 500 microns; the average pore diameter is more than 110 angstroms; crush strength from about 3 lb/mm (about 1.4 kg/mm) to about 4.5 lb/mm (about 2.0 kg/mm); specific surface area according to the BET method is more than 100 m ² /g or more than 125 m ² /g. The catalyst has the following physical and chemical properties: the active metal is cobalt; an active metal distribution of from about 2 percent to about 4 percent or from about 4 percent to about 6 percent; a weight percent of active metal on the support from about 15 weight percent to 25 weight percent; the promoter is ruthenium, rhenium, palladium or platinum; promoter weight percent from about 0.1 weight percent to about 0.5 weight percent. The first series-connected reactor makes it possible to convert more than 30% of CO and H2 into products; the second reactor makes it possible to convert more than 45% of the supplied material into products; the third reactor makes it possible to convert more than 60% of the supplied material into products. Three reactors connected in series provide at least 90% conversion efficiency, at least 92.5% conversion efficiency, or at least 95% conversion efficiency. This efficiency is achieved without recompression and/or recirculation of the catalyst tail gas.

The resulting liquid fuel and water each contain less than 100 ppm in the sum of acetic acid, propionic acid and malonic acid, less than 50 ppm in the sum, and preferably less than 25 ppm in the sum. The final water is directly recycled to provide steam for the synthesis gas generator.

2) To facilitate the conversion of hydrocarbons into synthesis gas, steam is used in the synthesis gas generator. Direct recycling of the water produced by the production process of the present invention results in greater than 60 weight percent, or 70 weight percent, or 80 weight percent steam. The oxygen concentration in the synthesis gas is less than 1000 ppm, less than 500 ppm, or preferably less than 250 ppm. Three multi-tubular catalytic reactors connected in series horizontally. Operating conditions: 400 to about 430°F (about 204 to about 221°C); from about 250 psig (about 1724 kPa g) to about 350 psig (about 2413 kPa g) or from about 350 psig (about 2413 kPa g) to about 450 psig (about 3103 kPa g); hourly gas space velocity from about 500 to about 1750 or from about 1750 to about 3000. A supported catalyst, wherein said support has an approximately neutral pH and wherein the support has the following physical and chemical properties: alumina as a material; petal shape; tablet radius less than about 500 microns; the average pore diameter is more than 110 angstroms; crush strength from about 3 lb/mm (about 1.4 kg/mm) to about 4.5 lb/mm (about 2.0 kg/mm); specific surface area according to the BET method is more than 100 m ² /g or more than 125 m ² /g. The catalyst has the following physical and chemical properties: the active metal is iron; an active metal distribution of from about 2 percent to about 4 percent or from about 4 percent to about 6 percent; a weight percent of active metal on the support from about 15 weight percent to 25 weight percent; the promoter is ruthenium, rhenium, palladium or platinum; promoter weight percent from about 0.1 weight percent to about 0.5 weight percent. The first series-connected reactor makes it possible to convert more than 30% CO and H ₂ into products; the second reactor makes it possible to convert more than 45% of the supplied material into products;

the third reactor makes it possible to convert more than 60% of the supplied material into products. Three reactors connected in series provide at least 90% conversion efficiency, at least 92.5% conversion efficiency, or at least 95% conversion efficiency. This efficiency is achieved without recompression and/or recirculation of the catalyst tail gas. The resulting liquid fuel and water each contain less than 100 ppm of the sum of acetic acid, propionic acid and malonic acid. The final water is directly recycled to provide steam for the synthesis gas generator.

3) To facilitate the conversion of hydrocarbons into gas, the synthesis gas generator uses steam. In re

- 12 045925 As a result of direct recycling of water obtained in the production process according to the present invention, more than 60 mass percent, or 70 mass percent, or 80 mass percent steam is obtained. The oxygen concentration in the synthesis gas is less than 1000 ppm or less than 500 ppm. Three multi-tubular catalytic reactors connected in series horizontally. Operating conditions: 400 to about 430°F (about 204 to about 221°C); from about 250 psig (about 1724 kPa g) to about 350 psig (about 2413 kPa g) or from about 350 psig (about 2413 kPa g) to about 450 psig (about 3103 kPa g), space velocity from about 500 to about 1750, or from about 1750 to about 3000. A supported catalyst, wherein said support has an approximately neutral pH and wherein the support has the following physical and chemical properties: alumina as a material; petal shape; tablet radius less than about 500 microns; the average pore diameter is more than 110 angstroms; crush strength from about 3 lb/mm (about 1.4 kg/mm) to about 4.5 lb/mm (about 2.0 kg/mm); specific surface area according to the BET method is more than 100 m ² /g or more than 125 m ² /g. The catalyst has the following physical and chemical properties: the active metal is nickel; an active metal distribution of from about 2 percent to about 4 percent or from about 4 percent to about 6 percent; a weight percent of active metal on the support from about 15 weight percent to 25 weight percent; the promoter is ruthenium, rhenium, palladium or platinum; promoter weight percent from about 0.1 weight percent to about 0.5 weight percent. The first series-connected reactor makes it possible to convert more than 30% of CO and H2 into products; the second reactor makes it possible to convert more than 45% of the supplied material into products; the third reactor makes it possible to convert more than 60% of the supplied material into products. Three reactors connected in series provide at least 90% conversion efficiency, at least 92.5% conversion efficiency, or at least 95% conversion efficiency. This efficiency is achieved without recompression and/or recirculation of the catalyst tail gas. The resulting liquid fuel and water each contain less than 100 ppm in the sum of acetic acid, propionic acid and malonic acid, less than 50 ppm in the sum, and preferably less than 25 ppm in the sum. The final water is directly recycled to provide steam for the synthesis gas generator.

4) Synthesis gas generator, which uses a modified piston engine. Direct recycling of the water produced by the production process of the present invention results in greater than about 70 weight percent, preferably about 80 weight percent, or more preferably 90 weight percent steam. The oxygen concentration in the synthesis gas is less than 1000 ppm, less than 500 ppm, or preferably less than 250 ppm. Three multi-tubular catalytic reactors connected in series horizontally. Operating conditions: 400 to about 430°F (about 204 to about 221°C); from about 250 psig (about 1724 kPa g) to about 350 psig (about 2413 kPa g) or from about 350 psig (about 2413 kPa g) to about 450 psig (about 3103 kPa g); space velocity (GHSV) from about 500 to about 1750 or from about 1750 to about 3000. A supported catalyst, wherein said support has an approximately neutral pH and wherein the support has the following physical and chemical properties: alumina as a material; petal shape; tablet radius less than about 500 microns; the average pore diameter is more than 110 angstroms; crush strength from about 3 lb/mm (about 1.4 kg/mm) to about 4.5 lb/mm (about 2.0 kg/mm); specific surface area according to the BET method is more than 100 m ² /g or more than 125 m ² /g. The catalyst has the following physical and chemical properties: the active metal is cobalt; an active metal distribution of from about 2 percent to about 4 percent or from about 4 percent to about 6 percent; a weight percent of active metal on the support from about 15 weight percent to 25 weight percent; the promoter is ruthenium, rhenium, palladium or platinum; promoter weight percent from about 0.1 weight percent to about 0.5 weight percent. The first series-connected reactor makes it possible to convert more than 30% CO and H ₂ into products; the second reactor makes it possible to convert more than 45% of the supplied material into products; the third reactor makes it possible to convert more than 60% of the supplied material into products. Three reactors connected in series provide at least 90% conversion efficiency, at least 92.5% conversion efficiency, or at least 95% conversion efficiency. This efficiency is achieved without recompression and/or recirculation of the catalyst tail gas. The resulting liquid fuel and water each contain less than about 100 ppm in the sum of acetic acid, propionic acid and malonic acid, less than 50 ppm in the sum, and preferably less than 25 ppm in the sum. The final water is used in the tertiary oil recovery process.

5) Synthesis gas generator, which uses a modified piston engine. Direct recycling of the water produced in the production process of the present invention results in greater than 60 weight percent, or 70 weight percent, 80 weight percent steam, or preferably 90 weight percent steam. The oxygen concentration in the synthesis gas is less than 1000 ppm, less than 500 ppm, or preferably less than 250 ppm. Three multi-tubular catalytic converters

- 13 045925 ski reactors connected in series horizontally. Operating conditions: 400 to about 430°F (about 204 to about 221°C); from about 250 psig (about 1724 kPa g) to about 350 psig (about 2413 kPa g) or from about 350 psig (about 2413 kPa g) to about 450 psig (about 3103 kPa g); space velocity (GHSV) from about 500 to about 1750 or from about 1750 to about 3000. A supported catalyst, wherein said support has an approximately neutral pH and wherein the support has the following physical and chemical properties: alumina as a material; petal shape; tablet radius less than about 500 microns; the average pore diameter is more than 110 angstroms; crush strength from about 3 lb/mm (about 1.4 kg/mm) to about 4.5 lb/mm (about 2.0 kg/mm); specific surface area according to the BET method is more than 100 m ² /g or more than 125 m ² /g. The catalyst has the following physical and chemical properties: the active metal is iron; an active metal distribution of from about 2 percent to about 4 percent or from about 4 percent to about 6 percent; a weight percent of active metal on the support from about 15 weight percent to 25 weight percent; the promoter is ruthenium, rhenium, palladium or platinum; promoter weight percent from about 0.1 weight percent to about 0.5 weight percent. The first series-connected reactor makes it possible to convert more than 30% of CO and H2 into products; the second reactor makes it possible to convert more than 45% of the supplied material into products; the third reactor makes it possible to convert more than 60% of the supplied material into products. Three reactors connected in series provide at least 90% conversion efficiency, at least 92.5% conversion efficiency, or at least 95% conversion efficiency. This efficiency is achieved without recompression and/or recirculation of the catalyst tail gas. The resulting liquid fuel and water each contain less than 100 ppm in the sum of acetic acid, propionic acid and malonic acid, less than 50 ppm in the sum, and preferably less than 25 ppm in the sum. The final water is used in the tertiary oil recovery process.

6) Synthesis gas generator, which uses a modified piston engine. Direct recycling of the water produced by the production process of the present invention results in greater than 60 weight percent, or 70 weight percent, or 80 weight percent steam. The oxygen concentration in the synthesis gas is less than 1000 ppm, less than 500 ppm, or preferably less than 250 ppm. Three multi-tubular catalytic reactors connected in series horizontally. Operating conditions: 400 to about 430°F (about 204 to about 221°C); from about 250 psig (about 1724 kPa g) to about 350 psig (about 2413 kPa g) or from about 350 psig (about 2413 kPa g) to about 450 psig (about 3103 kPa g); space velocity (GHSV) from about 500 to about 1750 or from about 1750 to about 3000. A supported catalyst, wherein said support has an approximately neutral pH and wherein the support has the following physical and chemical properties: alumina as a material; petal shape; tablet radius less than about 500 microns; the average pore diameter is more than 110 angstroms; crush strength from about 3 lb/mm (about 1.4 kg/mm) to about 4.5 lb/mm (about 2.0 kg/mm); specific surface area according to the BET method is more than 100 m ² /g or more than 125 m ² /g. The catalyst has the following physical and chemical properties: the active metal is nickel; an active metal distribution of from about 2 percent to about 4 percent or from about 4 percent to about 6 percent; a weight percent of active metal on the support from about 15 weight percent to 25 weight percent; the promoter is ruthenium, rhenium, palladium or platinum; promoter weight percent from about 0.1 weight percent to about 0.5 weight percent. The first series-connected reactor makes it possible to convert more than 30% CO and H ₂ into products; the second reactor makes it possible to convert more than 45% of the supplied material into products; the third reactor makes it possible to convert more than 60% of the supplied material into products. Three reactors connected in series provide at least 90% conversion efficiency, at least 92.5% conversion efficiency, or at least 95% conversion efficiency. This efficiency is achieved without recompression and/or recirculation of the catalyst tail gas. The resulting liquid fuel and water each contain less than 100 ppm of the sum of acetic acid, propionic acid and malonic acid. The final water is used in the tertiary oil recovery process.

When applying Method 1, 2 or 3 described above in a micro-GTL plant, allowing the conversion of approximately 0.05 to 2.5 million standard cubic meters. ft/d (about 1,415 to about 70,800 cu m/d) of gas phase hydrocarbons in about 5 to 250 bbl/d (about 0.79 to about 40 cu m/d) of liquid fuels, and when placing the specified micro -GTL installations at a site to which water is required to be transported (i.e., transported to the site by land, air or water from a location at least 2.5 miles (approximately 4 km)) to provide steam to facilitate conversion hydrocarbons into gas, recycling water containing less than about 100 ppm of the sum of acetic acid, propionic acid and malonic acid can reduce the cost of producing liquid fuels by at least 1 percent. IN

- 14 045925 in some cases, this reduces the cost of obtaining liquid fuels by at least 2 percent, by at least 3 percent, by at least 4 percent, or by at least 5 percent.

When applying Method 1, 2 or 3 described above in a micro-GTL plant, allowing the conversion of approximately 0.05 to 2.5 million standard cubic meters. ft/d (about 1,415 to about 70,800 cu m/d) of gas phase hydrocarbons in about 5 to 250 bbl/d (about 0.79 to about 40 cu m/d) of liquid fuels, and when placing the specified micro -GTL installations in such a site where it is necessary to use demineralized water (for example, water from a salt water source where said water has been treated to remove all or a significant amount of salt) to provide steam to facilitate the conversion of hydrocarbons to gas, recycling water containing less approximately 100 ppm in the sum of acetic acid, propionic acid and malonic acid, reduces the cost of obtaining liquid fuels by at least 1 percent. In some cases, this reduces the cost of obtaining liquid fuels by at least 2 percent, by at least 3 percent, by at least 4 percent, or by at least 5 percent.

When applying Method 4, 5 or 6 described above in a micro-GTL plant, allowing the conversion of approximately 0.05 to 2.5 million standard cubic meters. ft/d (about 1,415 to about 70,800 cu m/d) of gas-phase hydrocarbons in about 5 to 250 bbl/d (about 0.79 to about 40 cu m/d) of liquid fuels, the amount of oil recovered by Tertiary method of oil recovery using final water, averages at least one barrel (approximately 0.16 cubic meters) per day over a period of time of at least 30 days. In some cases, this amount averages at least two barrels (about 0.32 cubic meters) per day, at least five barrels (about 0.80 cubic meters) per day, or at least 10 barrels (about 1 .6 cubic meters) per day for at least 30 days.

Liquid fuels produced according to the present invention can be stored in steel, plastic, or other types of tanks commonly used for storing fuels at an average temperature greater than or equal to 60°F (about 16°C) for at least six months without the formation of more than 25 g of deposits / cubic meter of fuel according to the EN 1575 I standard without the addition of a biocide or treatment to improve fuel stability. See Czamock el al. Diesel Fuel Degradation During Storage Proceedings, CHEMIK 2015, 69, 11, 771-778. In some cases, liquid fuel can be stored under the same conditions without the formation of more than 20 g of deposits/cube, m, without the formation of more than 15 g of deposits/cube, m, without the formation of more than 10 g of deposits/cube, m or without the formation of more than 5 g of deposits /cube, m without adding biocide or treatment to increase fuel stability.

The above description of embodiments of the present invention has been provided for purposes of illustration and description only. The foregoing description is not intended to be exhaustive or to limit the present invention to the forms described. Accordingly, many modifications and variations will be apparent to those skilled in the art. Moreover, the above description is not intended to limit the present invention. The scope of the present invention is defined by the appended claims.

Examples

Examples of processes are provided and a description is given of how such processes can best be integrated with embodiments of typical processes used in the microscale GTL method of the present invention.

Although there are many ways to produce synthesis gas from gas-phase hydrocarbon resources, catalytic steam reforming is a typical method for producing synthesis gas because methane can be converted to synthesis gas at temperatures below about 1700°F (about 930°C) and typically in the range from 1600 to 1650 °F (from about 870 to about 900 °C), with the efficiency of this methane conversion being more than 97%.

The average composition of gas-phase hydrocarbon resources from oil wells varies widely depending on the location of the well as shown in Table. 5. However, for the sake of simplicity, the average composition of gas-phase hydrocarbons from North Dakota will be used to demonstrate the two preferred reforming methods.

- 15 045925

Table 5

Average composition of gases produced from oil wells in North Dakota and New Mexico

Gas phase components Concentration (volume percent) North Dakota (average) New Mexico (range) Methane 59.3 74-95 Ethane 17.7 0-10 Propane 9.4 0-5 Butane 2.7 0-3 Pentane+ 0.92 0-0.5 Carbon dioxide 0.51 0-10 Oxygen Undefined 0-0.2 Nitrogen 7.1 0-3

Although catalytic steam reforming of methane according to the reaction in Equation 1 should ideally provide an H ₂ /CO ratio of 3.0/1.0, according to Equation 2, additional H ₂ is formed from a portion of the methane contained in the synthesis gas, resulting in of which the stoichiometry of the reaction is determined by the equation. 3.

CH ₄ + H ₂ O —► CO + ZH2 Lv. 1 (main reaction) |0.20CH ₄ + 0.3N ₂ O -> 0.1 CO ₂ + 0.ICO + 0.7H2 Eq. 2 (minor reaction)

1.2CH ₄ + 1.3HgO -► 1.1CO + 0.1 CO2 + 3.7Ng Eq. 3 (combined reaction)

As a result, the H ₂ /CO ratio produced in a catalytic steam methane reformer is typically greater than 3.0 (Norbeck et al., 2008). As shown in Equation 3, the H ₂ /CO ratio is 3.36. According to Equation 3, the required molar ratio of H2O to carbon should be at least 1.08, but preferably in the range of 1.5 to 3.0 to eliminate the possibility of formation of elemental carbon.

Catalytic steam reforming of ethane, the second most abundant gas-phase hydrocarbon contained in the North Dakota sample, produced synthesis gas with an _H2 /CO ratio of 2.50/1.0 as defined by Equation 4.

C ₂ H ₆ + 2 H ₂ O 2CO + 5 H ₂ Eq. 4

Catalytic steam reforming of propane, the third most abundant gas-phase hydrocarbon contained in the North Dakota sample, produced synthesis gas with an _H2 /CO ratio of 2.33/1.0 as defined by Equation 5.

S ₃ N ₈ + ZN ₂ O ZSO + 7N ₂ Lv. 5

In table Figure 6 shows the composition of the synthesis gas obtained in the synthesis gas generator by introducing the average composition of gas-phase hydrocarbons extracted from oil wells in North Dakota. The H ₂ /CO ratio in this case was 3.00.

Table 6

Average composition of synthesis gas produced from North Dakota flare gas using catalytic steam reforming at ~1625°F (885°C)

Synthesis gas components H ₂ CO CH ₄ C ₂ + CO ₂ n ₂ Total Volume percentage 70.4 23.5 2.0 0.3 2.4 1.7 100.0

In table 7 shows the average composition of synthesis gas obtained from flare gas produced in North Dakota using catalytic steam reforming at ~1625°F (885°C) after membrane H ₂ separation to bring the H 2 /CO ratio to about 2.20 /1.00.

Table 7

The average composition of synthesis gas obtained from flare gas produced in North Dakota using catalytic steam reforming at ~1625°F (885°C) and membrane _H2 separation to bring the H2/CO ratio to 2.20/1, 00

Synthesis gas components n ₂ CO CH ₄ C ₂ + CO ₂ n ₂ Total Volume percentage 63.4 28.8 2.5 0.3 2.9 2.1 100.0

- 16 045925

The second preferred synthesis gas generator 103 uses a piston diesel engine that has been converted to operate on gas phase hydrocarbons in a dual-fuel mode. When such an engine is operated under enriched conditions (eg, a [fuel/air] equivalence ratio greater than about 1.8), synthesis gas is produced from gas-phase hydrocarbons.

In table Figure 8 summarizes the effect of equivalence ratio on CH4 conversion, _H2 /CO ratio and particulate matter emissions. A VPSA (Vacuum Pressure Swing Adsorption) system was used to increase the oxygen concentration in the air from 21% to approximately 93%. This level of enrichment allows the said engine to operate at an equivalence ratio of 2.0 to 2.2 and significantly reduces the nitrogen concentration in the synthesis gas. At the specified equivalence ratio, methane was converted with an efficiency of approximately 91 to 93% into synthesis gas, the H ₂ /CO ratio of which ranged from 1.62 to 1.85.

When a small amount of H2 (approximately 5%) 107 was added to gas-phase hydrocarbons, the _H2 /CO ratio increased to 1.83 and 1.95 with an equivalence ratio of 2.0 and 2.2, respectively. Such an amount of H2 can be obtained by recycling the tail gas 111 to the synthesis gas generator 103.

Particulate matter emissions at equivalence factors of 2.0 to 2.2 were approximately 0.09 to 0.13 mg/L (90,000 to 130,000 μg/m3). Since the pollutant specification for catalysts in a catalytic reactor is 500 µg/m3 or less, such particulate emission levels must be significantly reduced.

Since diesel particulate eliminators are established technologies, adding soot eliminators to a reformer with such an engine can reduce particulate emissions to below 500 µg/m3. These soot eliminators can operate in a lead-lag mode, so that the flow of synthesis gas is not interrupted.

Table 8

Effect of Equivalence Ratio on CH4 Conversion, _H2 /CO Ratio and Particulate Emissions for a Reciprocating Engine Reformer

Equivalence coefficient (CH ₄ /O ₂ ) % CH ₄ conversion H ₂ /CO ratio Particulate matter emissions (mg/ ^m3 ) Without adding _H2 Adding 5% H ₂ Without adding _H2 Adding 5% H ₂ 2.0 91 1.62 1.83 0.14 0.09 2.2 93 1.85 1.95 0.18 0.13

In table 9 shows the average composition of synthesis gas produced by converting flare gas produced in North Dakota using a reciprocating engine reformer where the tail gas from the catalytic reactor is recycled back to the engine reformer. The H ₂ /CO ratio in the synthesis gas was approximately 1.95 with an equivalence ratio of 2.2. This ratio is within the acceptable range for a direct fuel catalyst.

Table 9

Average composition of synthesis gas obtained from converting flare gas from North Dakota into synthesis gas using a reciprocating engine reformer (equivalence ratio 2.2) with tail gas recirculation from the catalytic reactor

n ₂ CO dn ₄ CO ₂ N ₂ Total 45 23 19 4 9 100.0

Catalytic reactor 108 consisted of a number of reactors connected in series, designed to efficiently convert syngas directly into high quality fuels in a single pass. This configuration eliminated the need to recompress and recirculate the tail gas through the catalytic reactor, which is typically used in the current art. This example illustrates how these three reactors are capable of converting approximately 90 mol% of the H2 and CO contained in the synthesis gas into products without recompressing or recycling the synthesis gas. The composition of the synthesis gas (Table 10) obtained in the synthesis gas generator 103 by catalytic steam reforming was used as a material supplied to the catalytic reactor 108 when using the direct fuel production catalyst 109.

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Table 10

Conversion of synthesis gas and production of products in a catalytic reactor using a direct fuel production catalyst

Components Composition (volume percentage) Feed material from synthesis gas generator ¹ Output from 1 reactor Manufactured products from 2 reactors Manufactured products from reactor 3 Manufactured products from 4 reactors H ₂ 63.4 55.8 36.3 22.1 5.7 CO 28.8 25.8 17.6 10.9 3.5 CH ₄ 2.5 2.9 3.5 6.2 13.1 n ₂ /co 2.20 2.12 2.06 2.03 1.97 c ₂ -c ₄ hydrocarbons o, s 1.1 1.8 2.9 3.9 from ₂ 2.9 3.2 4.7 6.6 7.1 n ₂ 2.1 2.3 3.4 4.5 5.9 Intermediate quantity Non-condensable products 100.0 91.1 67.3 53.2 39.2 from ₅ to ₂₄ hydrocarbons - 8.9 32.5 46.4 59.8 C ₂₄₊ hydrocarbons - - 0.2 0.4 0.6 Intermediate quantity Condensable products 0.0 8.9 32.7 46.8 60.8 Total 100.0 100.0 100.0 100.0 100.0

In this example, 10.4%, 38.9%, 62.2% and 88.9% of the CO were collectively converted to products in the first, second, third and fourth reactors, respectively. The overall conversion rates of _H2 and CO were 90.0% and 88.9%, respectively. Condensable products (C ₅ -C ₂₄ ) and non-condensable tail gases accounted for 60.8 vol.% and 39.2 vol.% of the total volume of gas-phase components. In table Figure 11 shows the composition of the tail gas after removal of condensable products.

Table 11

Composition of tail gas after conversion of synthesis gas in catalytic reactors with a direct fuel production catalyst

H ₂ CO dn ₄ from ₂ to ₅ from ₂ n ₂ Total 14.5 8.9 33.4 9.9 18.1 15.1 100.0

In conclusion, this series of four horizontal or vertical catalytic reactors, or reactors angled relative to horizontal and vertical orientations, effectively converted approximately 90% of the _H2 and CO contained in the synthesis gas into products while maintaining constant temperature and pressure gases in each reactor. In addition, there was no need to remove products from each stage and recycle tail gas to the reactors.

The direct fuel catalyst was operated under conditions of temperature, pressure and space velocity at which approximately 90% of the H ₂ and CO contained in the synthesis gas was converted to products. Liquid fuels consisting of C ₅ -C ₂₄ hydrocarbons comprised more than about 60 vol.%, about 70 vol.%, or preferably 75 vol.% of the gas phase components leaving the catalytic reactor. Such gas-phase components were collected as liquid fuels in a product separation and collection unit (Fig. 1: 110). In fig. Figure 4 shows the typical distribution of C ₅ -C ₂₄ hydrocarbons when the H ₂ /CO ratio is

- 18 045925 in the synthesis gas from the catalytic steam reformer was approximately 2.20/1.0 at the inlet, with an average inlet to outlet ratio of 2.08/1.0.

The distribution shown in Fig. 4, shifted slightly to the left (distribution toward lighter hydrocarbons) as the H ₂ /CO ratio increased above 2.08, and shifted slightly to the right (distribution toward heavier hydrocarbons) as the H ₂ /CO ratio decreased below 2.08.

The composition of the fuel at an average H ₂ /CO ratio of 2.08 is given in table. 12. Normal alkanes (Nalkanes) ranging from C8 to _C24 were the predominant hydrocarbons with a concentration of 65.4 vol.%. Normal C5-C7 alkanes accounted for 17.5 vol.% of the total volume. Approximately 1.0 volume percent of the sample was paraffinic hydrocarbons (C24+). The remainder of the liquid fuel sample was normal 1-alkenes, normal 1-hydroxyalkanes, and isoalkanes.

In the same system and when using the same catalyst, when the H ₂ : CO ratio shifted from the input ratio to 1.7, the output products shifted towards minimizing the amount of C4-C7 fraction and maximizing the amount of diesel fuel (including heavy diesel fraction or light paraffin factions).

In the same system and using the same catalyst, the reactor pressure can be adjusted to pressure levels greater than 350 psig (approximately 2413 kPag) to shift production toward heavier products (e.g., toward distribution of higher molecular weight products). In the same system and using the same catalyst, the reactor temperature can be adjusted to less than 410°F (210°C) to shift production toward heavier products. In the same system and using the same catalyst, the gas hourly space velocity (GHSV) in the reactor can be changed to less than 2000 h ^-1 to shift production towards heavier products.

In fig. Figure 5 shows the effect of changing the H ₂ /CO ratio on the concentration of 1-alkenes in liquid fuel. The concentration of these alkenes varied from about 3.5 vol.% at an H ₂ /CO ratio of 2.2 to 15.0 vol.% at an H ₂ /CO ratio of 1.5.

Table 12

Composition of liquid fuel

Composition of liquid fuel Volume percentage n-Alkanes (C5-C7) 17.5 n-Alkanes (C8-C24) 65.4 n-Alkanes (C24+) 1.0 Normal 1-alkenes (alpha-olefins) (C5-C7) 1.39 Normal 1-alkenes (alpha-olefins) (C8-C24) 1.78 Normal 1-hydroxyalkanes (C5-C7) 0.76 Normal 1-hydroxyalkanes (C8-C24) 3.91 Isoalkanes (C5-C7) 2.60 Isoalkanes (C8-C24) 5.70 Total 100.0

When the product recovery and separation unit is operated at a temperature of 50 to 55°F (10 to about 13°C), the composition of the tail gases is similar to that shown in Table 1. 11. The tail gas contained CH4 (33.4%), which was mainly produced by the catalytic reaction of the synthesis gas. The catalytic reaction produced small amounts of C2-C5 hydrocarbons and, because they have a high vapor pressure, they did not condense as liquid fuels. There was also some unreacted H2 and CO present because the direct fuel reactor was operated under conditions where approximately 89% of the CO was converted to products.

The amount of nitrogen in the tail gas depended on the nitrogen concentration in the gas-phase hydrocarbon feedstock. When using a reformer with an engine to produce synthesis gas, the concentration of nitrogen in the tail gas can increase significantly.

In table 13 shows the composition of synthesis gas produced by catalytic steam reforming by recycling tail gas 111 and catalyst reaction water 112 to synthesis gas generator 103 using an average composition of flare gas produced in North Dakota.

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Table 13

Composition of synthesis gas from a synthesis gas generator using catalytic steam reforming and recycling of catalyst reaction water and tail gas

n ₂ CO dn ₄ from ₂ to ₅ from ₂ N ₂ Total 56.3 28.1 2.0 1.0 8.0 4.6 100.0

In table 14 shows the average total yield (in gallons) of fuel obtained from 30,000 standard cubic meters. ft (approximately 850 cubic meters) of methane converted into synthesis gas at an average H ₂ /CO ratio of 2.08 by applying the proposed method. This volume of methane was chosen because it contained approximately 1,000 pounds of carbon. Therefore, the average production of fuel obtained from 500,000 standard cubic meters. ft of methane/day (approximately 14,200 cubic meters/day) was approximately 2.31 gallons/500,000 standard cubic meters. ft (approximately 8.7 L/14,200 cu.m.) or approximately 55 barrels/500,000 standard cu.m. ft. (approximately 8.7 cubic meters/approximately 14,200 cubic meters).

Table 14

Total Yield (gallons/30,000 standard cubic feet (gallons/approx.

850 cc cubic meters) of fuel products obtained from methane

Test No. Gallons/30,000 standard cu. ft (gallons/approximately 850 cubic meters) 8b 138 (approximately 522 l) 16a 140 (approx. 530 l) 17g 133 (approximately 503 l) Average value 139 (approximately 526 l)

In table 12 shows the concentrations of oxygenated organic compounds in catalyst reaction water obtained by using synthesis gas produced from various feedstocks through the proposed microscale GTL process. Oxygenated organic compounds in the catalyst reaction water consist of C1-C ₇ hydroxyalkanes (alcohols) in the case of synthesis gas obtained from three types of gas-phase hydrocarbon feedstocks. No carboxylic acids (<25 ppm) were detected in such water samples below the GC/MS detection limit.

When recycling catalyst reaction water containing these hydroxyalkanes to the synthesis gas generator, catalytic steam reforming of alcohols led to a decrease in the H ₂ /CO ratio.

Equations 3, 4 and 5 illustrate the reaction products and stoichiometry of the products obtained from reforming methanol, ethanol and propanol, taken as examples.

CHzOH + H ₂ O CO + 2H ₂ + H ₂ O Eq. 3

CH ₃ CH ₂ OH + 2 H ₂ O -+ 2CO + 4H ₂ + H ₂ O Eq. 4

CH ₃ CH ₂ CH ₂ OH + ZN ₂ O ZSO + 6H ₂ + H ₂ O Eq. 5

When steam is not used in the synthesis gas process, catalytic reaction water is ideal for injecting water into oil wells to enhance oil recovery.

Links.

References related to the topics discussed herein are cited as US Patents; publication of US patent applications; foreign patents and articles in magazines and books.

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US Patents

Cited Patents date Authors 9138688В2 09/2015 Prakash et al. 9090831 B2 7/2015 Schuetzle et al. 9067806 B2 6/2015 Camelli et al. 8999164 B2 4/2015 Franzosi et al. 8795597 B2 8/2014 Greer 8727767 B2 5/2014 Watson et al. 8591737В2 11/2013 Kukkonen et al. 8535487В2 9/2013 Camelli et al. 8529865 B2 9/2013 Belt et al. 8394862 B2 8/2013 Schuetzle et al. 8293805 B2 10/2012 Khan et al. 8158029В2 4/2012 Ernst 8057578 B2 11/2011 Argawal et al. 8048178В2 11/2011 Smith et al. 8043571В2 10/2011 Dannoux et al. 7989510В2 8/2011 Locatelli et al. 7939953 B2 5/2011 Lomax et al. US Patents Cited Patents date Authors 7744829 B2 6/2010 Brophy et al. 7559372 B2 7/2009 Cobb 7470405 B2 11/2008 Knopf et al. 7404936 B2 7/2008 Mazanec et al. 7323497 B2 1/2008 Abbot 7318894 B2 1/2008 Juby et al. 7276105 B2 10/2007 Pruet et al. 7261751 B2 8/2007 Dutta et al. 7235172 B2 6/2007 Lawson et al. 7166219В2 1/2007 Kohler et al. 7153432В2 12/2006 Kohler et al. 7150831В2 12/2006 Kohler et al. 7147775 B2 12/2006 Kohler et al. 7108070В2 9/2006 Hall et al. 6942839 B2 9/2005 Huisman et al. 6887908 B2 5/2005 Pruet et al. 6744066 B2 6/2004 Wang et al. 6576196В1 5/2003 Akporiaye et al. 6533945В1 3/2003 Shah et al. 6262131В1 7/2001 Arcuri et al. 6225358 A 5/2001 Kennedy et al. 6156809 A 12/2000 Clark et al. 5620670 A 4/1997 Benham et al. 5053581 A 10/1991 Hildinger et al. 4499209 A 2/1985 Hoek et al.

Publications of US patent applications

Publication of patent applications date Authors 2015/0259609 Al 9/2015 Wang et al. 2014/0144397 Al 5/2014 Bromberg et al. 2014/102981 Al 4/2014 Miglio et al.

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2014/0140896 Al 5/2014 Moon et al. 2012/0071572 Al 3/2012 Voolapelli et al. 2010/0184874 Al 7/2010 Hoek et al. 2010/0000153 Al 1/2010 Kurkjian et al. 2007/095570 Al 8/2007 Tomlinson et al. 2005/113426 Al 11/2005 Clur et al. 2005/0106086 Al 5/2005 Tomlinson et al. 2004/0262199 12/2005 Roelofse et al. 2003/0225169 12/2003 Yetman

Other patent applications

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Articles in magazines and books.

Asadullah, M.: Biomass gasification gas cleaning for downstream applications: A comparative critical review. Renewable and Sustainable Energy Reviews 40, 118-131 (2014).

ASTM International, Standard test method for corrosiveness to Copper from petroleum products by Copper Strip Test, ASTM D130-12, Conshohocken, PA (2012).

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Hoekman, S.K. et al.: Characterization of trace contaminants in syngas from the thermochemical conversion of biomass. Biomass Conversion and Biorefmery 3, 113-126 (2013).

Jahnagiri, H., Bennett, J., Mahjoubi, P., Wilson, K., Gu, S.: A review of advanced catalyst development for Fischer-Tropsch synthesis of hydrocarbons from biomass derived syngas. Catalysis Science and Technology 4, 2210-2229 (2014).

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Claims (10)

Lin, H., Zhou, M., Ly, J., Vu, J., Wijmans, J.G., Merkel, T.C., Jin, J., Haldeman, A., Wagener, E.H., Rue, D.: Membrane-Based Oxygen-Enriched Combustion, Ind. Eng. Chern. Res., 52, 1082010834(2013). Lim, E. G. et al.: The engine reformer: syngas production in an engine for compact gas toliquid synthesis. Canadian Journal of Chemical Engineering 34 (2016). McKendry, P.: Energy production from biomass gasification technologies 83, 55-63 (2002). O'Rrien, R.J., Davis, B.H.: Impact of copper on an alkali promoted iron Fischer-Tropsch catalyst. Catalysis Letters 64 (2004). Sa, S., Silva, H., Brandao, L., Mendes, A.: Catalysts for methanol steam reforming. Applied Catalysis In Environmental 99, 43-57 (2010). Schuetzle, D. et al.: The effect of oxygen on the formation of syngas contaminants during the thermochemical conversion of biomass. International Journal of Energy and Environmental Engineering, Springer-Verlag GmbH, Berlin, Heidelberg, Online ISSN: 2251 -6832 and Print ISSN: 2008-9163, 1-13 (2015). Wang, X.ct al.: Dilution sampling and analysis of particulate matter in biomass-derived syngas. Frontiers of Environmental Science & Engineering 5, 320-330 (2011). Yaying, J.: Partial oxidation of methane with air or O2 and steam to synthesis gas over a Nibased catalyst. Journal of Natural Gas Chemistry 9, 291-303 (2000). CLAIM 1. Microscale installation for producing liquid fuels from gas-phase hydrocarbon resources, including: a) synthesis gas generator; b) a catalytic reactor system comprising two or more horizontal reactors, vertical reactors or reactors angled relative to horizontal and vertical orientations and connected in series, wherein said reactor contains a catalyst for directly converting synthesis gas into liquid fuels and catalyst reaction water, and wherein said catalyst contains a support having a neutral surface pH; c) fuel product separation and collection unit. 2. Microscale plant according to claim 1, characterized in that the synthesis gas generator is selected from a steam reformer or a piston engine synthesis gas generator. 3. Microscale plant according to claim 1, characterized in that the synthesis gas generator contains a catalyst for converting gas-phase hydrocarbon resources into synthesis gas. 4. Microscale installation according to claim 1, additionally including a synthesis gas purification unit and a membrane system. 5. Microscale plant according to claim 1, characterized in that the catalytic reactor is a multi-tubular fixed-bed catalytic reactor and is configured to remove heat from the exothermic catalytic reaction through steam cooling. 6. Microscale plant according to claim 1, characterized in that the catalytic reactor system contains from two to five reactors or from six to ten reactors. 7. The microscale plant of claim 1, further comprising a utility unit including recirculation pumps, process control systems, and utility networks. 8. Microscale plant according to claim 1, characterized in that said catalyst for direct conversion of synthesis gas into liquid fuels and catalyst reaction water has a shape selected from petal, corrugated, bladed, spherical, granular and powder. 9. The microscale plant of claim 1, wherein said catalyst for direct conversion of synthesis gas into liquid fuels and catalyst reaction water has an average surface pore diameter of the support material greater than about 100 angstroms, wherein the support is selected from alumina, silica, zirconium dioxide, magnesium oxide or a combination of these metal oxides. 10. The microscale plant of claim 1, wherein said catalyst for direct conversion of synthesis gas to liquid fuels and catalyst reaction water has a crush strength of from about 1 lb/mm (about 0.45 kg/mm) to about 10 lb/mm (about 4.5 kg/mm) and a BET specific surface area greater than about 125 or 150 m ² /g. -