CN102325858B - Low-pressure fischer-tropsch process - Google Patents

Abstract

A Fischer-Tropsch process for producing diesel fuel or diesel blendstock with a high cetane number at a concentration of 65-90% by weight at a pressure below 200 psia using rhenium and/or or cobalt catalysts with ruthenium promoters. The catalyst is a cobalt catalyst having crystallites with an average diameter greater than 16 nanometers and the resulting hydrocarbon product contains less than 10% by weight of wax (> _C23 ) after rough flashing.

Description

Low pressure Fischer-Tropsch

technical field

The present invention generally relates to a low pressure Fischer-Tropsch process for the conversion of carbon monoxide and hydrogen into diesel fuel or diesel blendstock.

Background technique

The Fischer-Tropsch (FT) process for converting carbon monoxide and hydrogen into liquid power fuels and/or waxes has been known since the 1920s.

During World War II, the gasification of coal was used in Germany to supply a 1:1 ratio of hydrogen to carbon monoxide for conversion into fuel hydrocarbons to make synthetic diesel. Due to trade sanctions and lack of natural gas, South Africa has further exploited coal via the gasification of syngas and using fixed-bed iron Fischer-Tropsch catalysts. Iron catalysts are extremely active for the water gas shift reaction, which removes gas components from insufficient amounts of hydrogen, thus moving closer to the optimal _H2 /CO ratio of about 2.0. When developing large natural gas supplies, steam reformers and autothermal reformers are used to generate syngas feedstock for slurry bed FT reactors using cobalt or iron catalysts.

In Gas-To-Liquids (GTL) plants, a compromise must be made between liquid product yields and plant operating and capital costs. For example, if there is an electricity market, a steam reformer design can be chosen because this technology produces a lot of waste heat: the flue gas heat can be converted into electricity using an "economizer" and a steam turbine. If conservation of natural gas feedstock and low capital costs are paramount, then an autothermal or partial oxidation reformer using air is advantageous.

Another factor in selecting the optimum reformer type is the nature of the reformer hydrocarbon feed gas. This situation can be advantageous if the gas is rich in _CO2 , since the desired _H2 /CO ratio can then be achieved directly in the reformer gas without removing excess hydrogen and some of the _CO2 is Conversion to CO, thereby increasing the potential amount of liquid hydrocarbon products that may be produced. Additionally, reducing the amount of steam required reduces process energy requirements.

The market for Fischer-Tropsch (FT) processes is focused on large "world-scale" plants with natural gas feed rates greater than 200 million scfd due to considerable economies of scale. These plants operate at high pressure (approximately 450 psia) and use extensive recycle of off-gas in the FT reactor. For example, the Norsk Hydro plant design has a recycle ratio of approximately 3.0. The focus is on achieving maximum wax yield. In terms of product composition, these large plants seek the highest yield of Fischer-Tropsch synthetic waxes in order to minimize the formation of C1-C5 products. Thereafter, the wax is hydrocracked into mainly a diesel fraction and a naphtha fraction. Unfortunately, light hydrocarbons are also formed during this process. Reformers typically use some form of autothermal reforming using oxygen, which is cryogenically generated from air, an expensive process in terms of operating costs and capital costs. Economies of scale justify the use of high operating pressure, use of oxygen natural gas reforming, extensive recycle of tail gas to the FT reactor for increased syngas conversion and controlled heat removal and product wax hydrocracking. To date, no economical FT plant design has been developed for small plants with capacities less than 100 million scfd.

The catalytic hydrogenation of carbon monoxide to produce products ranging from methane to heavier hydrocarbons (up to C ₈₀ and higher) and oxygenated hydrocarbons is commonly known as Fischer-Tropsch synthesis. High molecular weight hydrocarbon products mainly comprise n-paraffins which cannot be used directly as motor fuels because they are incompatible with the low temperature properties of said motor fuels. After further hydrotreatment, the Fischer-Tropsch hydrocarbon product can be converted into higher value-added products, such as diesel fuel, jet fuel or kerosene. Therefore, it is desirable to directly maximize the production of high-value liquid hydrocarbons so that component separation or hydrocracking does not have to be performed.

Catalytically active group VIII, especially iron, cobalt and nickel, are used as Fischer-Tropsch catalysts; cobalt/ruthenium is one of the most common catalytic systems. In addition, catalysts typically contain supported metals and support metals and promoters, eg, rhenium.

Contents of the invention

According to one aspect of the present invention there is provided a Fischer-Tropsch (FT) process with a cobalt catalyst comprising crystallites, wherein the crystallites have an average diameter greater than 16 nanometers. The process produces a liquid hydrocarbon product containing less than 10% by weight wax (>C ₂₃ ) and greater than 65% by weight diesel (C ₉ -C ₂₃ ). The process may have an FT catalyst support for the cobalt catalyst, wherein the catalyst support is selected from the group of catalyst supports consisting of alumina, zirconia, titania and silica. The cobalt catalyst may have a catalyst loading greater than 10% by weight. The operating pressure for the Fischer-Tropsch process may be less than 200 psia. A promoter may be utilized in this process, in which case the promoter is selected from the group consisting of ruthenium, rhenium, rhodium, nickel, zirconium, titanium and mixtures thereof. Flash distillation can be performed on the process to reduce the naphtha fraction. The process may use an FT reactor that does not use tail gas recycle. The method can also use a reformer that uses air as the oxygen source. The reactor can be a fixed-bed FT reactor or a slurry bubble-bed FT reactor.

According to another aspect of the present invention there is provided an FT process which operates at less than 200 psia, which uses an air autothermal reformer, and which has a CO conversion of at least 65%, And a cobalt catalyst was used to provide greater than 60 wt% diesel yield in a single pass FT reactor. The catalyst has a metallic cobalt loading of greater than 5% by weight and a rhenium loading of less than 2% by weight on a catalyst support material selected from the group of catalyst support materials comprising: alumina, zirconia and silica. The cobalt catalyst is in the form of crystallites, wherein the crystallites have an average diameter greater than 16 nanometers. The FT catalyst support material may consist of alumina. This process can have a feed gas where hydrogen is removed from the feed gas using selective membranes or molecular sieves. Alternatively, the cobalt catalyst loading can be greater than 6% by weight and the operating pressure can be less than 100 psia. The reactor may further have a promoter, wherein the promoter comprises a promoter selected from the group of promoters consisting of ruthenium and rhenium or mixtures thereof.

According to yet another aspect of the present invention there is provided an FT process which operates at less than 200 psia, which uses an oxygen autothermal reformer, and which has a CO conversion of at least 65%, And a cobalt catalyst is used to provide greater than 60 wt% diesel yield in the FT reactor. The catalyst has a metallic cobalt loading of greater than 5% by weight and a rhenium loading of less than 2% by weight on a catalyst support selected from the group of catalyst supports consisting of alumina materials, zirconia materials, and bismuth Silicon oxide material. The cobalt catalyst is in the form of crystallites having an average diameter greater than 16 nanometers. The FT catalyst support may consist of alumina. The method may comprise off-gas from the reformer, wherein part of the off-gas is recycled to the reformer. The method may also include a feed gas, wherein a selective membrane or molecular sieve is used to remove hydrogen from the gas. Alternatively, the cobalt catalyst loading can be greater than 6% by weight and the operating pressure can be less than 100 psia. The reactor may further have a promoter, wherein the promoter comprises a promoter selected from the group of promoters consisting of ruthenium or rhenium or mixtures thereof.

According to yet another aspect of the present invention there is provided an FT process operating at less than 200 psia, the FT process employs an oxygen steam reformer, and the FT process has a CO conversion of at least 65%, and Diesel yields of greater than 60% by weight are provided in an FT reactor using a cobalt catalyst having a loading of metallic cobalt of greater than 5% by weight and a loading of rhenium of less than 2% by weight on a catalyst support selected from A catalyst support group consisting of: alumina material, zirconia material or silica material or mixtures thereof. The cobalt catalyst is in the form of crystallites, wherein the crystallites have an average diameter greater than 16 nanometers. The FT catalyst support may consist of alumina. The process may include a feed gas from which hydrogen is removed using a selective membrane or molecular sieve. The method may further include off-gas from the reformer, wherein some or all of the off-gas is combusted to provide heat to the reformer. Alternatively, the cobalt catalyst loading can be greater than 6% by weight and the operating pressure can be less than 100 psia. The reactor may further have a promoter, wherein the promoter comprises a promoter selected from the group of promoters consisting of ruthenium or rhenium or mixtures thereof.

According to yet another aspect of the invention there is provided an FT process which operates at less than 200 psia, which uses an air or oxygen partial oxidation reformer, and which has a CO conversion of greater than 65% rate, and using a cobalt catalyst having a metallic cobalt loading of greater than 5 wt % and a rhenium loading of less than 2 wt % on the FT catalyst support to provide a diesel yield of greater than 60 wt % in an FT reactor, the The FT catalyst support is selected from the group of catalyst supports comprising: alumina material, zirconia material and silica material. The cobalt catalyst is in the form of crystallites, and the crystallites have an average diameter greater than 16 nanometers. The FT catalyst support may consist of alumina. The process may include a feed gas from which hydrogen is removed using a selective membrane or molecular sieve. The method may further include off-gas from the reformer, wherein some or all of the off-gas is combusted to provide heat to the reformer. Alternatively, the cobalt catalyst loading can be greater than 6% by weight and the operating pressure can be less than 100 psia. The reactor may further have a promoter, wherein the promoter comprises a promoter selected from the group of promoters consisting of ruthenium or rhenium or mixtures thereof.

Description of drawings

Fig. 1 is the process flow chart of the specific embodiment of the present invention;

Figure 2 is a flow chart for the flash separation of the naphtha fraction and the diesel hydrocarbon fraction as a subsequent step of the Fischer-Tropsch process;

Figure 3 is a graph showing the C5+ carbon number distribution of catalysts (trilobes) of Example 3 at 190°C;

Figure 4 is a graph showing the effect of pressure on catalyst performance of Example 4;

Figure 5 is a graph of the C5+ carbon number distribution of the catalyst of Example 7 at 190°C, 70 psia;

Figure 6 is a graph of the C5+ carbon number distribution of the catalyst (LD-5) of Example 8a at 200°C, 70 psia;

Figure 7 is a graph of the C5+ carbon number distribution of the catalyst (F-220) of Example 9 at 190°C, 70 psia;

Fig. 8 is the graph that uses the C5+ carbon number distribution of the catalyst of the example 10 of ruthenium promoter;

Figure 9 is a graph of the C5+ carbon number distribution of the catalyst of Example 11 (Aerolyst 3038 silica);

Figure 10 is a graph showing the relationship between cobalt catalyst crystallite size and wax content of C5+FT products; and

FIG. 11 is a graph showing a comparison of the carbon distribution of the catalyst used in Example 9 with the traditional Anderson-Shultz-Flory distribution.

^* In all figures of the graph showing carbon numbers, naphtha is represented by large squares, diesel by diamonds, and light wax by small squares.

Detailed ways

introduce

In the Fischer-Tropsch process, various parameters such as the size and shape of cobalt crystallites affect the activity of supported cobalt catalysts. The size of the metal crystallites controls the number of active sites available for reduction (dispersion) and the degree of reduction.

Under certain pretreatment and activation conditions, strong interactions between cobalt metal and oxide supports form unfavorable cobalt-support structures (e.g., cobalt-aluminum spinel), which may require high reduction temperatures . High reduction temperature may lead to sintering of cobalt crystallites and formation of large cobalt metal clusters. Not only temperature treatment, but also cobalt metal precursors and metal loadings as well as metal promoters also affect the size of cobalt crystallites. Low cobalt metal loadings may result in high metal dispersion and small crystallites, but enhance metal-support interactions leading to poor reducibility and low catalyst activity.

The hydrogenation of carbon monoxide using cobalt-supported catalysts is proportional to the amount of exposed cobalt atoms. Therefore, logically, increasing the dispersion of cobalt metal on the surface of the oxide support will increase the catalyst activity and C5+ selectivity. However, small cobalt crystallites strongly interact with the oxide support, forming an irreducible cobalt-support system. A strong correlation between cobalt metal crystallites and reducibility can affect catalyst activity and may result in undesirable products. Under typical Fischer-Tropsch reaction conditions, the cobalt crystallite size range (9-200 nm) and dispersion range (11-0.5%) have a minor impact on C5+ selectivity. However, smaller cobalt crystallites suffer severe deactivation. In fact, Barbie et al. 2000 studied the correlation between deactivation rate and cobalt crystallite size and observed a peak at 5.5 nm.

Embodiments of the invention

Embodiments of the invention described herein relate to low pressure Fischer-Tropsch processes and catalysts that produce high diesel fraction yields. Process pressures can be below 200psia. The catalyst is cobalt deposited on gamma alumina at greater than 5% by weight, optionally with 0.01-2% by weight rhenium or ruthenium, and has crystallites with an average diameter greater than 16 nanometers. Such catalysts have been found to be very effective in the conversion of syngas to diesel at low pressures at high yields, resulting in diesels (C9-C23) containing less than 10 wt% waxes (>C23) and greater than 65 wt% liquid hydrocarbon products. Embodiments of the invention are particularly well suited for converting low pressure gases containing low molecular weight hydrocarbons to FT liquids. Examples of applications are landfill gas, oilfield solution gas, and low pressure gas from depressurized gas fields. In all these cases, multiple stages of gas compression and air compression would be required in a conventional FT plant. The high efficiency of the FT catalyst of the present invention enables high CO conversion and produces a product stream containing up to 90+ wt% diesel in a single pass. The use of air in the natural gas reformer provides a syngas containing about 50% nitrogen, which facilitates the removal of heat in the FT reactor as sensible heat and improves gas velocity and heat transfer efficiency so that tail gas recycle is not required . The naphtha can be partially separated from the hydrocarbon products by low-cost flash distillation to produce a purer diesel product. This move is also used to provide some product cooling. Liquid hydrocarbon products are excellent for blending with petrodiesel to increase cetane number and reduce sulfur content.

Embodiments of the present invention can be applied not only to world-scale gas-to-liquid plants, but also to small FT plants using less than 100 million scfd. Embodiments of the invention, when applied to small FT plants, seek to make the best economic sense focusing on simplicity and minimizing investment costs, possibly at the expense of efficiency. The following is a comparison of the existing FT technology and the embodiment of the present invention when applied to a small FT factory:

Existing FT technology Embodiments of the present invention

In order to operate an FT process at high conversion using oxygen reformer syngas, the process must recycle tail gas at high ratios—3.0 or greater on a fresh feed gas basis. A secondary benefit is that the fresh gas is diluted in carbon monoxide, which reduces the rate required to remove heat from the FT reactor, reduces localized heating and improves product composition. However, exhaust recycling is an energy and capital intensive activity. Separating oxygen from air is also an energy and capital intensive activity.

The approach employed in the process of the present invention is to use air in the reformer, which produces a syngas containing approximately 50% nitrogen as an inert diluent, thereby eliminating the need for tail gas recycle to moderate the heat removal requirements of the FT reactor. Other approaches using air-blown syngas in the FT process have achieved the desired high CO conversions by using multiple FT reactors in succession, which required high capital costs and complex operations. The inventive process achieves high CO conversion in a simple single pass and achieves high diesel cuts through the use of specific catalysts as described more explicitly below.

The catalyst in one embodiment uses an alumina support with a high cobalt concentration, along with low levels of rhenium to facilitate catalyst reduction. High cobalt concentration increases catalyst activity, thereby enabling high per-pass synthesis gas conversion. The catalyst is made to have a relatively large average cobalt crystallite size, and this gives selectivity to a product that is essentially diesel.

Anderson-Shultz-Florey theory predicts that FT hydrocarbons cover a very wide range of carbon numbers (from 1 to 60), and the most ideal product is diesel fuel (C ₉ -C ₂₃ , defined by Chevron). In order to reduce the "loss" of CO to produce _C1 - _C5 hydrocarbons, the common approach is to strive for the production of mainly wax in the FT reactor, and then, in the separation operation, to try to hydrocrack the wax to mainly diesel and naphtha. Unexpectedly, the method and catalyst of the present examples produce diesel in high yields (up to 90% by weight) directly in the FT reactor, thereby avoiding the need for expensive and complex hydrocracking facilities.

Because oxygen purification, high pressure compression, tail gas recycle and hydrocracking are eliminated, the inventive process can be economically applied in much smaller plants than hitherto considered possible for FT technology.

Fig. 1 shows the process flow chart of the FT method of the embodiment of the present invention, and wherein letter A-K represents following content:

A Raw hydrocarbon-containing gas

B hydrocarbon gas conditioning equipment

C Reformer

D water

E Oxidizing gas

F cooler

G separator

H Hydrogen removal (optional)

I Fischer-Tropsch Reactor

J back pressure controller

K Product cooling and recovery (2-option)

The letter A denotes unprocessed hydrocarbon-containing treated feed gas. This hydrocarbon-containing process feed gas can come from various sources: for example, from natural gas fields, landfill facilities (biogas), petroleum processing facilities (dissolved gas), and the like. The pressure of the gas used in the process of the invention can vary widely, from atmospheric pressure to 200 psia or higher. Single or two stages of compression may be required, depending on source pressure and desired process operating pressure. For example, with landfill gas, the pressure is usually close to atmospheric pressure and a blower is used to deliver the gas to the combustion facility. The dissolved gas (usually flash) must also be compressed to process operating pressure. There are also many previously developed and later stage natural gas fields with pressures too low for acceptance into pipelines that could produce potential feedstocks for the process of the present invention. Other sources of natural gas that may or may not be distressed (failure to pipeline) may already be at or above the desired process operating pressure and are also candidates. Another candidate is natural gas with an inert gas content such as nitrogen that is too high to meet pipeline specifications.

The letter B indicates hydrocarbon gas conditioning equipment. The gas may need to be cleaned to remove components that would damage the reformer or FT catalyst. Examples of such components that will damage reformer or FT catalysts are mercury, hydrogen sulfide, silicones and organic chlorides. Organic chlorides such as those found in landfill gases produce hydrochloric acid in the reformer, which can cause severe corrosion. The silicone forms a continuous silica that coats the catalyst, blocking the pores. Hydrogen sulfide is a powerful FT catalyst poison and is typically removed to 1.0 ppm or less. Some gases from low sulfur fields may not require any conditioning (purification).

The hydrocarbon concentration in the raw gas affects the economics of processing because less hydrocarbon product is formed from the same volume of feed gas. However, the process may operate at methane concentrations of 50% or less, for example, using landfill gas. There may even be reasons for financially infringing local practices: for example, to meet greenhouse gas management or corporate emissions standards. The process can be operated with a feed gas containing only methane-based hydrocarbons or natural gas liquids by applying known reformer technology. Carbon dioxide is advantageously present in the feed gas.

The letter C designates a reformer, which can be of several types, depending on the composition of the feed gas. Significant benefits of low pressure reformer operation are the lower rate of the Brouard reaction and the reduction of metal dust.

Partial oxidation reformers typically operate at very high pressures (ie, 450 psia or greater) and are therefore not optimal for low pressure FT processes. Partial oxidation reformers are energy inefficient and can easily produce soot, however, the partial oxidation reformer does not require water and produces syngas with a _H2 /CO ratio close to 2.0, the syngas Gas is optimal for FT catalysts. Partial oxidation reformers may be used in the process of the invention.

Steam reformers are expensive to invest in and flue gas heat recovery is required in large plants to maximize efficiency. Because the syngas contains relatively low levels of inert gases such as nitrogen, temperature control in the FT reactor can be difficult without tail gas recycle to the FT reactor. However, the low level of inert gas allows some tail gas to be recycled to the reformer side leg to supplement the natural gas feed, or the low level of inert gas allows some tail gas to be recycled to the side shell to provide heat. It should be kept in mind that in any case the FT tail gas must be combusted before it is vented and this energy can be used to generate electricity or better yet be used to provide reformer heat which would otherwise be provided by burning natural gas. For small FT plants, a steam reformer is a viable option. Steam reformers may be used in the process of the invention.

An efficient process with relatively low capital costs, autothermal reforming uses moderate temperatures and moderate steam concentrations to produce soot-free gas with _H2 /CO of about 2.5 using low _CO2 natural gas feedstocks Syngas, _H2 /CO about 2.5 is closer to the desired ratio than _H2 /CO produced by steam reforming. However, for most natural gas feedstocks, some hydrogen still needs to be removed. If the feed gas contains greater than about 33% _CO2 (which is typically the case with landfill gas feedstocks), then a _H2 /CO ratio of 2.0 can be achieved without any recycle flow, and water usage can also be reduced. This is the most desirable type of reformer for the low pressure FT process of the present invention.

Letter D designates optional water injected as steam into the reformer. All reformer technologies except partial oxidation reformers require steam injection.

The letter E represents an oxidizing gas, and the oxidizing gas can be air, oxygen or oxygen-enriched air.

The letter F designates a cooler for reducing the reformer outlet temperature from greater than 700°C to close to ambient temperature. Cooling is preferably performed in a single stage, although cooling may be performed in several stages. Cooling can be achieved using shell and tube or plate and frame heat exchangers, and the recovered energy can be used to preheat the reformer feed gas, as is well known in the industry. Another way to cool the reformer off-gas is by injecting water directly into the stream or by passing the stream through water in a vessel.

The letter G designates a separator that is used to separate the reformer syngas from condensed water in order to minimize the amount of water entering downstream equipment.

The letter H indicates optional hydrogen removal equipment, such as the Prism ^™ hydrogen selective membrane sold by Air Products, or the Cynara membrane from Natco.

Certain reformer processes produce syngas with excess hydrogen, some of which must be removed for optimal FT reactor performance. The ideal _H2 /CO ratio is 2.0-2.1, while raw syngas may have a ratio of 3.0 or higher. High hydrogen concentrations cause greater CO loss, producing methane instead of the desired motor fuel or motor fuel precursors such as naphtha.

The letter I designates a typical FT reactor of the fixed bed or slurry bubble type, and either can be used. However, fixed beds are preferred in small plants because of their simplicity of operation and ease of scale-up.

The letter J designates a back pressure controller that sets the process pressure. The back pressure controller can be placed in other locations, depending on product recovery and possible fractional separation processes used.

The letter K indicates product cooling and recovery. Product cooling is typically achieved by exchanging heat for cold water and used to preheat hot water for use elsewhere in the FT plant. Separation is achieved in separator vessels designed for oil/water separation. However, a second option is to flash cool the FT reactor product before the aforementioned cooler-separator as shown in FIG. 2 . This serves two purposes - firstly to reduce the product temperature and secondly to enable partial separation of the naphtha component in the hydrocarbon product produced thereby enriching the remaining liquid in the diesel component.

Figure 2 illustrates a process diagram for the flash separation of naphtha and diesel hydrocarbons, where:

1 is a fixed bed Fischer-Tropsch reactor.

2 is a mixture of gas, water, naphtha, diesel oil and light wax at about 190-240°C and pressure greater than atmospheric pressure.

3 is a pressure discharge valve.

4 is stream 2 at 14.7 psia at a reduced temperature due to gas expansion.

5 is the flash tank container.

6 is the vapor phase consisting of stream 2 minus diesel and light wax.

7 is cooler.

8 is stream 6 with liquid phase naphtha and water.

9 is a container for holding naphtha and water.

10 is the waste off-gas stream mainly composed of inert gases and light hydrocarbons.

FT product 2 flows through pressure discharge valve 3 and into flash tank 5 . Inert gases and lower boiling hydrocarbons, water and naphtha exit the flash tank as vapor overhead and pass through cooler 7 . Diesel oil and light wax are collected in container 5. The water and naphtha condense in cooler 7 and are collected in container 9 . The remaining gas exits stream 10 overhead and is typically combusted (sometimes with energy recovery), or used to generate electricity.

example

Catalyst carrier used

Table 1. Physical properties of catalyst supports

Example 1

Catalytic syntheses are carried out by common building blocks as practiced by those skilled in the art. The catalyst support was alumina trilobe extrudate (hereinafter referred to as 'trilobe') obtained from Sasol Germany GmbH. The extrudate dimensions were 1.67 mm in diameter and 4.1 mm in length. The carrier was calcined at 500° C. for 24 hours in air. A solution mixture of cobalt nitrate and perrhenic acid was added to the support by incipient wetness to achieve 5% by weight cobalt metal and 0.5% by weight rhenium metal in the finished catalyst (Catalyst 1). The catalyst is oxidized in the following three steps:

Step 1: Heat the catalyst to 85°C and keep it warm for 6 hours;

Step 2: Raise the temperature to 100°C at 0.5°C per minute and keep it warm for 4 hours;

Step 3: The temperature was raised to 350° C. at 0.3° C. per minute and kept for 12 hours.

The drying rate of wet catalyst depends to some extent on the size of the catalyst particles. Smaller particles will dry faster than larger particles, and the size of the crystals formed within the pores can vary with the rate of crystallization. A volume of 29 cc of oxidation catalyst was placed in a 1/2 inch outside diameter (OD) tube having an outer annular space through which temperature controlled water flowed under pressure to remove the heat of reaction . In fact, the FT reactor is a shell and tube heat exchanger with catalyst placed in a side pipe. Both the inlet gas and water were at the target reaction temperature. Catalyst reduction is achieved by the following procedure:

Reducing gas flow rate (cc/min)/H ₂ in nitrogen (%)/temperature (°C)/time (hour):

1.386/70/200/4, warm-up phase

2.386/80/ to 325/4, slow heating stage

3.386/80/325/30, fixed temperature stage

During Fischer-Tropsch catalysis, the total gas flow to the FT reactor was a gas volumetric space velocity (Gaseous hourly space velocity; GHSV) of 1000 hr ⁻¹ . The gas composition is representative of the air autothermal reformer gas: 50% nitrogen, 33.3% H2 and 16.7% CO. Air drying of the catalyst is used to reduce methane production. This was accomplished by maintaining the reactor temperature at 170°C for the first 24 hours. Presumably, this process produces carbonylation of the cobalt surface and enhanced FT activity. CO conversion and liquid production were measured at various temperatures between 190°C and 220°C.

Example 2

The catalyst (Catalyst 2) used in this example was the same as that used in Example 1, except that the cobalt metal loading was 10% by weight.

Example 3

The catalyst (Catalyst 3) used in this example was the same as that used in Example 1 except that the cobalt metal loading was 15% by weight.

Example 4

The catalyst (Catalyst 4) used in this example was the same as that used in Example 1, except that the cobalt metal loading was 20% by weight.

Example 5

The catalyst (Catalyst 5) used in this example was the same as that used in Example 1, except that the cobalt metal loading was 26% by weight.

Example 6

The catalyst (Catalyst 6) used in this example was the same as that used in Example 1 except that the cobalt metal loading was 35% by weight.

Example 7

The catalyst (Catalyst 7) used in this example was the same as that used in Example 1, except that the alumina support was CSS-350 from Alcoa, and the cobalt loading was 20% by weight. This carrier is spherical with a diameter of 1/16 inch.

Examples 8a, 8b, 8c and 8d

The catalysts used in these examples (Catalysts 8a, 8b, 8c, and 8d) were identical to those used in Example 1 except for the following differences: except that the alumina support was LD-5 from Alcoa and the cobalt loading was 20 Catalyst 8a is identical to Catalyst 1 except for weight %. This support was spherical with a mean particle distribution of 1963 microns. Example 8a used the as-is particle size mixture. Some of the primary particles were ground to a smaller sieve size: Catalysts 8b, 8c and 8d were made into particles having diameters of 214 microns, 359 microns and 718 microns, respectively. The cobalt loading in Examples 8b, 8c and 8d was consistent with catalyst 8a.

Example 9

The catalyst (Catalyst 9) used in this example was the same as that used in Example 1, except that the alumina support was F-220 from Alcoa, and the cobalt loading was 20% by weight. F-220 is a spherical support with a mesh size distribution of 7/14.

Example 10

The catalyst (Catalyst 10) used in this example was the same as Catalyst 4, except that the promoter was ruthenium instead of rhenium.

Example 11

The catalyst (Catalyst 11) used in this example was the same as Catalyst 3, except that Aerolyst 3038 silica catalyst support from Degussa was used instead of alumina.

Example 12

The catalyst (Catalyst 12) used in this example was identical to Catalyst 8d, with the same catalyst support, particle size and catalyst loading, except that the retention time of the oxidation process during the catalytic synthesis was doubled. That is, for the 3 oxidation steps described for Catalyst 1, the temperature holding times amounted to 12 hours, 8 hours and 24 hours, respectively. The intent of the slower catalyst oxidation rate for the small catalyst 12 particles was to achieve a larger cobalt crystallite size (21.07 nm). The methods used herein to control drying rate and catalyst cobalt crystallite size are not meant to exclude any other method of achieving larger crystallite size. For example, the relative humidity or pressure of the drying chamber can be varied to control the catalyst drying rate, and thus the cobalt crystallite size.

Catalyst Characterization

A Chembet 3000 (Quantachrome Instruments) TPR/TPD analyzer was used to analyze the average crystallite size (d(CoO), degree of dispersion (D%), and degree of reduction ( _DOR ) of the above catalysts. catalyst, and the cobalt dispersion was calculated assuming that one hydrogen molecule covers two cobalt surface atoms. Oxygen chemisorption was measured at 380°C using a series of ( _O2 /He) pulses across the catalyst after reduction of the catalyst at 325°C .Determine the molar amount of oxygen uptake, and calculate the degree of reduction, assuming that all cobalt metal is reoxidized to Co ₃ O ₄ .Cobalt crystallite size is calculated according to the following formula:

d(CoO)=(96/D%)DOR

D%: Dispersion

FT Catalyst Evaluation

(i) Effect of cobalt loading

Examples 1-6 were used to test the effect of Co loading on catalyst performance, and the results are shown in Table 2.

Table 2. Effect of Catalyst Loading on Performance for Examples 1-6 (trilobes) at 70 psia

The tests for each of Examples 1-6 were conducted at various temperatures, and the temperature that produced the greatest amount of hydrocarbon product is listed. Clearly, 5% cobalt is not sufficient to provide useful quantities of liquid hydrocarbons; the optimal concentration is 20 wt% Co, which yields 1.03 ml/h of liquid hydrocarbons. At cobalt loadings of 10 wt% cobalt or higher, the concentration of diesel range hydrocarbons in the hydrocarbon product ranged from 75.3-92.5%. The highest diesel production rate (0.78 ml/h) was achieved at 70 psia using the trefoil support with 20% cobalt.

The performance data of Catalyst 1 at 202.5°C are shown in Table 8. The level of wax (C > 23) on C5+ fluids was only 6.8%, and the diesel fraction was 73.5% (C9-C23). It was found that for all catalysts tested with an average crystallite diameter greater than 16 nm, the C5+ wax was less than 10% by weight, enabling the product to be used directly as a diesel blendstock.

FIG. 3 illustrates the carbon number distribution of Catalyst 3 (trilobe) in Example 3 at 190°C. Gets a very narrow spread without heavy waxes. Diesel is 90.8%, naphtha is 6.1%, and light wax is 3.1%. The cetane number is very high at 88. In the chart of all carbon numbers, naphtha is represented by large squares, diesel by diamonds, and light waxes by small squares.

effects of stress

Catalyst 4 in Example 4 was run at various pressures at a temperature of 202.5°C in the standard test apparatus as described above. The results in Table 3 and Figure 4 show that the productivity of the catalysts used to produce liquid hydrocarbons is significant at low pressures as low as 70 psia, with the best results obtained at pressures between 70 psia and 175 psia. A pressure of 70-450 psia is preferred, and a pressure of from 70 psia to 175 psia is most preferred. The diesel fraction over that pressure range is suitably constant at 70.8-73.5% by weight. As shown in Table 8, Catalyst 4 with 20% cobalt had an average crystallite size of 22.26 nm and a C5+ wax fraction of 6.8 wt%, enabling the product to be used as a diesel blendstock.

Table 3. Effect of pressure on catalyst performance (catalyst 4, 202.5°C)

Catalyst 7

As seen in Table 4, the maximum diesel production rate was achieved at 215°C and 70 psia. Compared to Catalyst 4, Catalyst 7 produced a lower diesel production rate but a higher diesel fraction at the optimum temperature of Catalyst 7 (215°C). Figure 5 illustrates the narrow range of carbon numbers in the liquid product at 190°C with 89.6% in the diesel range. The cetane number is 81. However, as shown in Table 8, the crystallite size was 18.26 nm and the wax fraction was 7.2%, enabling the product to be used as a diesel blendstock.

Table 4. Performance of Catalyst 7 (CSS-350) at various temperatures

Catalysts 8a, 8b, 8c and 8d

The test results are shown in Table 5. Catalysts 8b, 8c and 8d are shown to have a higher Co metal dispersion than catalyst 8a. Catalysts containing average Co crystallite sizes below 16 nm produced a high wax fraction of 17.6-19.3 wt% in the FT product, while catalysts 8a and 12 containing Co crystallites ^larger than 16 nm ^produced C5+ liquids of A lower wax fraction of 6.6% and 7.8% by weight enables the product to be used as a diesel blendstock. Notably, Catalyst 8a and Catalyst 12 produced similar low wax fractions, although Catalyst 8a and Catalyst 12 had very different particle sizes. This suggests that crystallite size, rather than particle size, controls the variation at low wax concentrations.

Table 5. Performance of Catalysts 8a-8d and Catalyst 12 at 70 psia

Catalyst 9

Catalyst 9 was tested at 70 psia. As shown in Table 6 and Figure 7, the 190°C hydrocarbon product contained 99.1% "naphtha plus diesel" diesel itself was 93.6%. Very little light wax is present. The cetane number is 81. As shown in Table 8, the crystallite size was 22.22 nm and the wax fraction was 2.3%, enabling the product to be used directly as diesel fuel.

Table 6. Performance of Catalyst 9 (F-220) at various temperatures (pressure 70 psia)

Catalyst 10

The data in Table 7 and Figure 8 demonstrate that the use of a ruthenium catalyst promoter in place of rhenium also provides a narrow hydrocarbon distribution with 74.42% in the diesel range with a total cetane number of 78. As shown in Table 8, the crystallite size was 20.89 nm and the wax fraction was 3.73%, enabling the product to be used as a diesel blendstock.

Table 7. Properties of catalyst 10 (ruthenium promoter, LD-5 alumina support)

Catalyst 11

For Catalyst 11, the hydrocarbon liquid production rate was 0.55 ml/h at 210°C. The carbon distribution curve shown in Figure 9 exhibits a narrow distribution with a high diesel fraction. As shown in Table 8, the crystallite size was 33.1 nm and the wax fraction was 5.2%, enabling the product to be used as a diesel blendstock, perhaps after flashing off the naphtha fraction.

Table 8. Summary of the Effect of Cobalt Crystallite Size on C5+ Wax Concentration

Catalysts 1 to 12 in this disclosure (except catalysts 8b, 8c and 8d) show that when the FT catalyst has cobalt crystallites larger than 16 nm, hydrocarbons with low wax content (<10 wt%) are obtained (mainly at Diesel range) as shown in Figure 10 (large squares are not part of this example). In the case of small catalyst particles (eg, Catalyst 12), it is necessary to control the rate of crystallization to obtain the desired crystallite size.

Figure 11 compares this result with the expected value from the Anderson-Shultz-Flory (A-S-F) carbon number distribution based on chain growth. The A-S-F distribution provided a diesel fraction of only 50% by weight, whereas the inventive examples provided a diesel fraction of >65% by weight.

The liquid hydrocarbon product of the catalyst of the present invention is more valuable than the products of the broad A-S-F type, because the liquid hydrocarbon product of the catalyst of the present invention can be directly used as a raw material for diesel blending without hydrocracking to increase the cetane number and reduce petrochemical Diesel sulfur content. Because the inventive method can be a simple one-way process, the inventive method can require low capital costs.

While this disclosure describes and illustrates preferred embodiments of the invention, it should be understood that the invention is not limited to these particular embodiments. Many variations and modifications will now occur to those skilled in the art. For a full definition of the invention and its intended scope, reference is made to the Summary of the Invention and the appended claims, which are read and considered in conjunction with the disclosure and drawings herein.

Claims (39)

1. A Fischer-Tropsch process for producing liquid hydrocarbons comprising diesel fuel or diesel blending feedstock, said process producing a liquid hydrocarbon product containing less than 10% by weight wax and greater than 65% diesel, said wax being >C ₂₃ wax, the diesel oil is C ₉ -C ₂₃ diesel oil, and the Fischer-Tropsch method comprises: operate at pressures below 200 psia; and A cobalt catalyst is used comprising a Fischer-Tropsch catalyst support having cobalt metal crystallites thereon having an average diameter greater than 16 nanometers. 2. The method of claim 1, wherein the Fischer-Tropsch catalyst support is a catalyst support selected from the group of catalyst supports consisting of alumina, zirconia, titania, silica and mixtures thereof. 3. The method of claim 2, wherein the alumina is gamma alumina. 4. The method of claim 1, wherein the cobalt catalyst has a metallic cobalt loading, and wherein the metallic cobalt loading is at least 15% by weight. 5. The process of claim 1, wherein the conversion of CO in the feed gas is at least 60%. 6. The method according to any one of claims 1 to 5, wherein a promoter is used in the method and is selected from the group of promoters consisting of: ruthenium, rhenium, rhodium, nickel , zirconium and titanium and mixtures thereof. 7. The method of any one of claims 1 to 5, wherein flash distillation is performed to reduce the light hydrocarbon fraction, which has a lower boiling point than diesel. 8. The method of claim 1, wherein the method uses a Fischer-Tropsch reactor that does not use tail gas recycle. 9. A method according to any one of claims 1 to 5 or 8, wherein the method uses a reformer which uses air as a source of oxygen. 10. The method according to any one of claims 1 to 5 or 8, wherein the Fischer-Tropsch reactor used in the Fischer-Tropsch process is a fixed-bed Fischer-Tropsch reactor or a slurry bubbling bed reactor -Tropsch reactor. 11. A Fischer-Tropsch process operating at less than 200 psia, using an air autothermal reformer, and having a CO conversion of at least 60% and provide a diesel yield greater than 65% by weight in a single-pass Fischer-Tropsch reactor, the Fischer-Tropsch process comprising the steps of: Using a cobalt catalyst having a metallic cobalt loading of at least 15% by weight and a rhenium loading of less than 2% by weight, the cobalt catalyst has a catalyst support material selected from the group of catalyst support materials consisting of: alumina, zirconia , silica, and mixtures thereof, and the cobalt catalyst has cobalt metal crystallites thereon having an average diameter greater than 16 nanometers. 12. The method of claim 11, wherein the Fischer-Tropsch catalyst support material consists of gamma alumina. 13. The process of claim 11 having a Fischer-Tropsch feed gas wherein selective membranes or molecular sieves are used to remove hydrogen from the Fischer-Tropsch feed gas. 14. The process of claim 11, wherein the operating pressure is at least 40 psia and the temperature in the Fischer-Tropsch reactor is at least 190°C. 15. The method of claim 11, wherein the operating pressure is less than 100 psia. 16. The Fischer-Tropsch process of claim 11, the cobalt catalyst further having a promoter, wherein the promoter comprises a promoter selected from the group of promoters consisting of ruthenium, rhenium, and mixtures thereof. 17. A Fischer-Tropsch process having a CO conversion of at least 60% and which provides a diesel yield in a Fischer-Tropsch reactor greater than 65% by weight, said The Fischer-Tropsch method consists of the following steps: Operates at pressures less than 200 psia; using an oxygen autothermal reformer; and Using a cobalt catalyst having at least 15% by weight metallic cobalt loading and less than 2% by weight rhenium loading on a Fischer-Tropsch catalyst support material selected from catalyst support materials consisting of The group: alumina, zirconia, silica, and mixtures thereof, wherein the cobalt catalyst is in the form of cobalt metal crystallites having an average diameter greater than 16 nanometers. 18. The method of claim 17, wherein the Fischer-Tropsch catalyst support consists of alumina. 19. The process of claim 17 having off-gas from a Fischer-Tropsch reformer, wherein the off-gas is partially recycled to the reformer. 20. The process of claim 17, further having a Fischer-Tropsch reactor feed gas, wherein selective membranes or molecular sieves are used to remove hydrogen from the feed gas. 21. The process of claim 17, wherein the operating pressure is at least 40 psia and the temperature in the Fischer-Tropsch reactor is at least 190°C. 22. The method of claim 17, wherein the operating pressure is no greater than 100 psia. 23. The Fischer-Tropsch process of claim 17, said reactor further having a promoter, wherein said promoter comprises a promoter selected from the group of promoters consisting of ruthenium, rhenium, and mixtures thereof. 24. A Fischer-Tropsch process for a Fischer-Tropsch reactor, said Fischer-Tropsch process comprising the steps of: Operates at pressures less than 200 psia; Using an oxygen steam reformer; having a CO conversion of at least 60% and providing a diesel yield of greater than 65% by weight; and Using a cobalt metal catalyst having at least 15% by weight metallic cobalt loading and less than 2% by weight rhenium loading on a Fischer-Tropsch catalyst support material selected from the group consisting of The group of catalyst support materials: alumina, zirconia, silica and mixtures thereof, wherein the catalyst support material is supported with cobalt metal crystallites, the crystallites having an average diameter greater than 16 nanometers. 25. The method of claim 24, wherein the Fischer-Tropsch catalyst support consists of gamma alumina. 26. The process of claim 24 further having a Fischer-Tropsch reactor feed gas, wherein selective membranes or molecular sieves are used to remove hydrogen from the feed gas. 27. A method as claimed in any one of claims 24 to 26 having off-gas from the reformer wherein some or all of the off-gas is combusted to feed the reformer Provide heat. 28. The method of any one of claims 24 to 26, wherein the operating pressure is at least 40 psia and the temperature is at least 190°C. 29. The method of any one of claims 24 to 26, wherein the operating pressure is less than 100 psia. 30. The Fischer-Tropsch process according to any one of claims 24 to 26, said reactor further having a promoter, wherein said promoter comprises a promoter selected from the group of promoters consisting of: ruthenium, Rhenium and its mixtures. 31. A Fischer-Tropsch process having a CO conversion greater than 60% and providing a diesel yield of greater than 65% by weight, the Fischer-Tropsch process comprising: Operate at less than 200psia; using air or oxygen partial oxidation reformers; and Using a Fischer-Tropsch reactor with a cobalt catalyst having a loading of metallic cobalt of greater than 15% by weight and a loading of rhenium of less than 2% by weight on a Fischer-Tropsch catalyst support material, the The Fischer-Tropsch catalyst support material is selected from the group of catalyst support materials consisting of alumina, zirconia and silica and mixtures thereof, wherein the cobalt catalyst is in the form of metal crystallites having a diameter greater than 16 nanometer average diameter. 32. The method of claim 31, wherein the Fischer-Tropsch catalyst support consists of alumina. 33. A process as claimed in claim 31 or 32 having a Fischer-Tropsch reactor feed gas wherein selective membranes or molecular sieves are used to remove hydrogen from the feed gas. 34. The method of claim 31, wherein the operating pressure is at least 40 psia and the temperature is at least 190°C. 35. The method of claim 31, wherein the operating pressure is less than 100 psia. 36. The Fischer-Tropsch process of claim 31 , said reactor further having a promoter, wherein said promoter comprises a promoter selected from the group of promoters consisting of ruthenium, rhenium, and mixtures thereof. 37. The process of any one of claims 1, 11, 17, 24, or 31, wherein the temperature in the Fischer-Tropsch reactor is at least 190°C. 38. The process of any one of claims 1, 11, 17, 24, or 31, wherein the temperature in the Fischer-Tropsch reactor is at least 190°C and the operating pressure is at least 40 psia, wherein A promoter is used in the process, the promoter being selected from the group of promoters consisting of ruthenium, rhenium, rhodium, nickel, zirconium, titanium or mixtures thereof; and wherein the CO conversion is greater than 65%. 39. The method of claim 38, wherein the CO conversion is greater than 65%.