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bioweb.sungrant.org » Technical » Bioproducts » Bioproducts from Syngas » Fischer-Tropsch Synthesis
| Products from Syngas—Fischer-Tropsch Synthesis Products |
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In its simplest form, syngas (also called producer gas, town gas, blue water gas, and synthesis gas among others) is composed of carbon monoxide (CO) and hydrogen (H2), which provide the building blocks to produce a number of organic compounds. During the 1800’s, coal gasification was used for lighting and heating. The production of fuels and chemicals from syngas began in the early 20th century, and today, commercial production of methanol, ammonia, and hydrogen are mostly produced from syngas, although numerous other compounds can also be produced.
In principle, syngas can be produced from any hydrocarbon feedstock, including natural gas, naphtha, residual oil, petroleum coke, coal, and biomass. Under today’s conditions, the least expensive feedstock is natural gas. Using numerous synthesis pathways, a large number of organic compounds can be produced from syngas (figure 1).
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| Fischer-Tropsch synthesis reactions |
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Fischer-Tropsch (FT) products are produced using four main steps--syngas generation, gas purification, FT synthesis, and product upgrading (figure 2).
In the FTS process, autothermal reforming alone or in combination with steam reforming, are genereally recommended for the production of syngas from natural gas due to the resulting H2/CO ratio and a favorable economy of scale for air separation units (Abbott, 2002; Rostrup-Nielsen, 2000; Vosloo, 2001; Wilhelm, 2001). For coal, the syngas is produce via high temperature gasification in the presence of oxygen and steam. FT synthesis temperatures are usually kept below 400°C to minimize methane (CH4) production. FT reactors are operated at pressures ranging from 10-40 bar (145–580 psi). About 20% of the carbon contained in natural gas is converted to CO2 during reforming, while about 50% of the carbon in coal is converted due to the lower hydrogen content of coal (Dry, 2002). Coal gasification produces a syngas with a H2:CO ratio of about 0.67, providing a theoretical maximum conversion rate of CO to FT products of 33% (with no water gas shift reactions) (Raje, 1997). The theoretical maximum conversion of natural gas (methane) to paraffins is 78% on a how heat value (LHV) basis.
The mix of products depends on reactor temperature and pressure, the feed gas composition, and the types of catalysts and promoters used. Depending on the types and quantities of FT products desired, either low (200–240°C) or high temperature (300–350°C) synthesis is used with either a cobalt or iron catalyst respectively. Low temperature synthesis yields high molecular mass linear waxes while high temperatures produce gasoline and low molecular weight olefins. Production of gasoline products is highest under conditions of high temperatures using an iron catalyst in a fixed fluid bed reactor. The theoretical maximum is 48% conversion of the syngas. Production of diesel fractions is maximized in a slurry reactor using low temperatures and a cobalt catalyst with maximum production closer to 40%.
Upgrading involves use of a combination of hydrotreating, hydrocracking, and hydroisomerization in addition to product separation.
FT synthesis is a polymerization reaction consisting of (1) absorpotion of the reactant (CO) on the catalyst surface, (2) initiation of the polymer chain caused by the dissociation of the CO from the catalyst surface and followed by hydrogenation, (3) chain growth resulting from the insertion of additional CO molecules followed by hydrogenation, (4) termination of the polymer chain, and (5) desorption of the FT product from the catalyst surface. Chemisorbed methyl compounds are formed by dissociation of absorbed CO molecules and stepwise addition of hydrogen atoms. These methyl compounds can further hydrogenate to form methane or act as initiators for chain growth. Chain growth occurs via sequential addition of CH2 groups while the growing alkyl chain remains chemisorbed to the metal surface at the terminal methylene group. Chain termination can occur at any time during the chain growth process to yield either an α-olefin or an n-paraffin once the product desorbs. The general FT synthesis reaction is shown in equation 1 (Haid, 2000).
(eq. 1) CO + 2H2 à --CH2-- + H2O ΔHr (227°C; -165kJ/mol)
The water-gas shift (WGS) reaction is a secondary reaction that readily occurs when iron (Fe) catalysts are used (equation 2).
(eq. 2) CO + H2O à H2 + CO2
The net FT synthesis reaction when using Fe catalysts combines the general and WGS reactions and is shown in equation 3.
(eq. 3) 2CO + H2 à --CH2-- + CO2
The required H2 to CO ratio for a cobalt catalyzed FT reaction is 2.15 and is 1.7 for a Fe catalyzed reaction due to the combined FT and WGS reactions (Dry, 2002).
Irrespective of operating conditions, FT synthesis always produces a range of olefins, paraffins, and oxygenated compounds (alcohols, aldehydes, acids, and ketones) as shown in equations 4 through 7.
(eq. 4) CO + 3H2 à CH4 + H2O (Methanation)
(eq. 5) nCO + (2n+1)H2 à CnH2n+2 + nH2O (Paraffins)
(eq. 6) nCO + 2nH2 à CnH2n + nH2O (Olefins)
(eq. 7) nCO + 2nH2 à CnH2n+1OH + (n-1)H2O (Alcohols)
Regardless of the product type, FT products are predominantly linear with high olefinicity. The paraffin-to-olefin ratio is lower than predicted based on thermodynamics, and the olefins that are formed are predominantly terminal (alpha). Significant quantities of monomethyl chain branches form with branching decreasing as chain length increases. The Boudouard reaction (equation 8) is an important competing reaction.
(eq. 8) 2CO à Cs + CO2
FT synthesis is kinetically controlled and the intrinsic kinetics is a stepwise chain growth (e.g., a polymerization of CH2 groups on a catalyst surface). The polymerization rates, and therefore the kinetics, are independent of the products formed which are determined by the ability of the catalyst to catalyze chain propagation versus chain termination reactions. The probabilities of chain growth and chain termination are independent of chain length, and thus the efficiency of producing various hydrocarbons can be predicted based on simple statistical distributions based on the chain growth probability and carbon number. The chain polymerization kinetics model (the Anderson-Shulz-Flory model) is represented by equation 9 with the predicted distributions for several chemicals of commercial interest
shown graphically in figure 3.
(eq. 9) Wn = n(1-α)2αn-1
Wn is the weight percent of a product containing n carbon atoms and α is the chain growth probability.
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| Fischer-Tropsch synthesis catalysts |
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Catalysts play a pivotal role in syngas conversion reactions. The basic concept of a catalytic reaction is that reactants adsorb onto the catalyst surface, rearrange, and combine into products that desorb from the surface. One of the fundamental functional differences between syngas synthesis catalysts is whether or not the adsorbed CO molecule dissociates on the catalyst surface. For FTS and higher alcohol synthesis, CO dissociation is a necessary reaction condition.
Group VIII transition metal oxides (e.g., Ru, Fe, Ni, Co, Rh, Pd, and Pt) are generally regarded as good CO hydrogenation catalysts (Adesina, 1996). Iron (Fe) and cobalt (Co) catalysts are most commonly used in FT synthesis.
Fe is very active; readily forms carbides, nitrides, and carbonitrides with metallic character that also have FTS activity; and undergoes the water gas shift (WGS) reaction. However, Fe catalysts also have a stronger tendency to produce carbon that deposits on the surface and deactivates the catalyst than do other catalysts such as Ni or Co. The use of promoters and supports are essential for Fe catalytic activity. Potassium is often used as a promoter because it increases the adsorption of the CO to the catalyst surface by increasing the alkalinity of the surface and providing a source of electrons which are removed by the CO from the Fe. The addition of potassium oxide to Fe catalysts also decreases the hydrogenation of adsorbed carbon compounds which results in the production of higher molecular weight products that are more olefinic. However, the use of potassium promoters also increases WGS activity causing greater deposition of carbon on the catalyst surface and a faster rate of deactivation of the Fe catalyst. Copper has also been successfully used as a promoter for Fe catalysts. Copper promoters increase the rate of synthesis more than potassium, decrease the rate of the WGS reaction, and facilitate iron reduction. Cu promoters increase the average molecular weight of the compounds produced, although not as much as when potassium is used. The combination of light transition metal oxides (e.g., MnO) with Fe increases the selectivity of light olefins in FTS. Fe/Mn/K catalysts have shown selectivity for C2-C4 olefins as high as 85-90%.
Co catalysts tend to have longer lifetimes than Fe catalysts and do not have water gas shift activity, which leads to improved carbon conversion to products because CO2 is not formed. In FTS processes, Co catalysts yield mainly straight chain hydrocarbons rather than oxygenated hydrocarbons that occur with Fe catalysts. Co catalysts are 230 times more expensive than Fe but are still used because they demonstrate activity at lower synthesis pressures so the higher catalyst cost can be offset by lower operating costs. Co catalysts are not sensitive to the addition of promoters. Early work demonstrated that the addition of ThO2 improved wax production at atmospheric pressure but had little effect at higher pressures. The addition of noble metals to Co catalysts increases FTS activity but not selectivity. Potassium promoters are not effective with Co catalysts.
Other catalysts can be used. Ni is basically a methanation catalyst and does not have the broad selectivity of other FT catalysts. Ru displays very high activity and selectivity for larger molecular weight products at low temperatures, but is also much more expensive (3x105 times) than Fe which is the least expensive FTS catalyst.
FTS catalysts can lose activity as a result of conversion of the active metal site to an inactive oxide site, sintering, loss of active area by carbon deposition, and chemical poisoning. Some of these mechanisms are unavoidable and others can be prevented or minimized by insuring that the impurity levels in the incident syngas are acceptable for the given process.
The deposition of carbon (coke) on the catalyst is the most important mode of catalyst deactivation. It is affected by the addition of promoters, reaction temperature, and reaction pressure. It is generally unavoidable during FT synthesis, and thus the process must be operated in a way that balances the tradeoffs between lower output due to coke deposition and catalyst regeneration and replacement costs. Because of its high activity, coke deposition rates are typically higher for Fe catalysts than Co catalysts, and consequently, Co catalysts have longer lifetimes.
Catalytic active sites can be deactivated by impurities in the syngas. Sulfur is major catalyst poison which is present in both natural gas and coal and which is converted to H2S and organic sulfur compounds during steam reforming or gasification. Sulfur compounds rapidly deactivate both Fe and Co catalysts, presumably by forming surface metal sulfides that do not have FTS activity. Co catalysts are more sensitive to sulfur poisoning than Fe catalysts. Other impurities in syngas that can poison FTS catalysts include halides and nitrogen compounds (NH3, NOx and HCN) (Turk, 2001). Water oxidizes both Fe and Co catalysts, with a higher oxidation rate for Fe. Water also inhibits Fe activity due to the WGS activity of Fe catalysts (Espinoza, 2000). Table 1 summarizes syngas impurities and tolerance levels.
Commercial processes are available to clean syngas to meet impurity tolerances. The Rectisol process uses chilled methanol to scrub the raw syngas and to remove NH3, H2S, tars, and CO2. Other chemical absorption processes include potassium carbonate or alkanolamine (MEA –monoethanolamine or DEA – diethanolamine) for wet scrubbing. Fixed bed reactors containing ZnO are also used for sulfur polishing. The level of gas cleaning required must balance the costs of gas cleaning with the reduced lifetime of the catalyst, with the gas cleaning costs a function of the scale (size) of the FTS operation. Given the relative cost of Co versus Fe, more efficient sulfur removal should be expected for FTS with Co catalysts.
The focus of catalyst development is on improved catalyst lifetimes, activity, and selectivity which are affected by the use of chemical and structural promoters, catalyst preparation and formulation, pretreatment and reduction, selective poisoning, and shape selectivity with zeolites. Catalyst preparation impacts the performance of catalysts. Fe catalysts can be prepared by precipitation onto catalyst supports such as SiO2 or Al2O3 or as fused iron where formulations are prepared in molten iron, cooled and crushed. The role of supports in Co catalysts is also important. Since Co is more expensive than Fe, precipitating the ideal concentration of metal onto a support can help reduce catalyst costs while maximizing activity and durability. |
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| Fischer-Tropsch synthesis reactors |
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The highly exothermic FTS reactions generate a large amount of heat and reactor design and process development have focused on heat removal and temperature control. Insufficient heat removal leads to localized overheating which causes high carbon deposition on the catalysts and subsequent deactivation of the catalyst. Methane formation also dominates at higher temperatures at the expense of desired FTS products. For large-scale commercial FTS reactors heat removal and temperature control are the most important design features to obtain optimum product selectivity and long catalyst lifetimes. Commercial FTS reactors consist of 4 basic designs (figure 4).
The fixed bed tubular reactor was one of the earliest FTS designs. After many years of development, Ruhrchemie and Lurgi refined the design into the ARGE high capacity FT reactor that has been used for many years. The reactor contains 2,000 tubes filled with Fe catalyst immersed in boiling water for heat removal. The water bath temperature is maintained in the reactor by controlling the pressure. Syngas is introduced into the top of the reactor, flows through the tubes, and the products exit at the bottom of the reactor. Wax accounts for 50% of the products. Conversion efficiencies are on the order of 70%. The reactor is operated at 20-30 bar at an operating temperature of 220-260ºC. Additional temperature control is obtained by using high gas velocities and gas recycling. Recycled gas typically is 2.5 times the amount of fresh syngas introduced into the reactor. Catalyst lifetimes are around 70-100 days (Wender, 1996) and catalysts removal can be quite difficult.
High temperature circulating fluidized bed reactors have been developed for gasoline and light olefin production. These reactors are known as Synthol reactors and operate at 350°C and 25 bar. The combined fresh and recycled gas feed enters at the bottom of the reactor and entrains catalyst that is flowing down the standpipe and through the slide valve. The high gas velocity carries the entrained catalyst into the reaction zone where heat is removed through heat exchangers. Product gases and catalyst are then transported into a large diameter catalyst hopper where the catalyst settles out and the product gases exit through a cyclone. These Synthol reactors have been successfully used for many years, but are constrained due to the physical complexity of the reactors (circulation of large amounts of catalyst leads to erosion within the reactor), and capacity and throughput limitations (7,500 BPD) (Lutz, 2001).
The fixed fluidized bed Sasol Advanced Synthol reactor has replaced the circulating fluidized bed Synthol reactor. Gas is introduced through a distributor and bubbles up through the catalyst bed. Heat is removed by an internal heat exchanger immersed in the catalyst bed. Process conditions in the fixed fluidized bed reactors are similar to those established in the Synthol reactors. These new reactors have the same capacity, better thermal efficiency with a less severe temperature gradient, and a lower pressure drop across the reactor, but are half the cost and size of the circulating reactors. Operating costs are also considerably lower, process flexibility (in terms of product distribution) is greater, and they offer the potential for scale-up to 20,000 BPD (Lutz, 2001).
The low temperature slurry reactor was initially developed in the 1950s by Kolbel (Dry, 1996, 2002). These 3-phase reactors consist of a solid catalyst suspended and dispersed in a high thermal capacity liquid (often the FT wax product). Syngas is bubbled through the liquid phase achieving contact with the catalyst while also keeping the catalyst particles dispersed. Slurry reactors are optimized at low temperatures for FT wax production with low methane production. Compared to the fluidized bed reactors, liquid slurry bed reactors have better temperature control, lower catalyst loading, and significantly lower catalyst attrition rates. The improved isothermal conditions in slurry bed reactors allows for higher average reactor temperatures leading to greater conversion of syngas to products. Compared with multitubular fixed bed reactors, slurry reactors have lower pressure differences across the reaction resulting in lower costs for gas compression, reduced catalyst consumption rates (four-fold reduction), and operates at higher temperatures and has higher conversion rates (Dry, 1982). Slurry bed reactors also cost less. A significant disadvantage of slurry reactors is that poisons in the syngas will affect all of the catalyst in the reactor, whereas in a fixed tube design, they will primarily affect only the catalyst near the gas inlet. The reliable separation of the catalyst from FT waxes posed a significant technical barrier that has been largely overcome and these reactors are now beginning to be used in commercial applications. Sasol’s slurry reactor process has a thermal efficiency of about 60% and a carbon conversion efficiency of about 75% (Lutz, 2001).
Commercial production of Fischer-Tropsch compounds from syngas.
The first FT plants (9 plants, 660,000 tonnes/yr) began operation in Germany in 1938, but were closed following World War II (Dry, 2002). Sasol established its first FT plant (Sasol I) in Sasolburg, South Africa in 1955 (capacity of 6 million tonnes FT products/yr from coal). Since then, numerous FT plants of varying reactor designs have been established in South Africa and produce more than 200 fuel and chemical products (e.g., gasoline, diesel, candle waxes, hard waxes, hydrocarbon lubricants, methane, phenol and cresol, tar and pitch, ammonia, detergents, many of which are exported) using coal as the feedstock (http://www.gasandoil.com/goc/company/cna02527.htm).
Sasol supplies about 41% of South Africa’s liquid transportation fuel requirements. Sasol plans to use the natural gas as a supplementary or replacement feedstock for the coal in many of their facilities (http://www.oil-barrel.com/archives/features_archive/2002/jan-2002/sasol310102.htm) and http://www.eia.doe.gov/emeu/cabs/safrica.html). The Mossgas plant in South Africa, using natural gas as the feedstock, produces 1 million tones FT products/yr including motor gasoline, distillates, kerosene, alcohols and LPG (Dry, 2002).
Shell (using the Shell Middle Distillate Synthesis process) produces 500,000 tonnes of FT products/yr (12,000 BPD/yr) using natural gas and a cobalt catalyst at its facility in Bintuli, Malaysia (Senden, et al, 1992). The plant was built in 1993 and produces mostly automotive fuels, specialty chemicals, and waxes. The conversion rate is about 63% (Senden, 1992). Syntroleum operates a 10,000 BPD plant in Australia using natural gas to produce liquid fuels and specialty products (Haid, 2000).
Most of the capital cost for FTS plants is for the generation of the syngas (Dry, 2002; Senden, 1992; Vosloo, 2001) and the cost of producing syngas from methane is 30% lower than that from coal and is a more efficient process (Dry, 2002). The capital cost of the Shell plant in Malaysia was about $660 million (Senden, 1992). The capital cost of replacing Sasol’s 16 circulating fluidized bed reactors with 8 fixed fluidized bed reactors was $225 million with an operating cost reduction projected to be $1/BBL for the same production capacity (Chang, 2000). Studies by Sasol indicate that a two train GTL plant composed of slurry phase reactors producing 30,000 BPD of liquid transport fuels can be constructed at a capital cost of about $25,000/BPD capacity including utilities, offsite facilities, and infrastructure (Lutz, 2001; Vosloo, 2001). Sasol estimated that a plant producing 10,000 BBL transportation fuel/day (425,000 ton/yr) using slurry phase reactors and natural gas (100 MMSCFD/day) required a capital investment of around $300 million ($30,000/daily BBL) (Jager, 1998).
Feedstocks represent a significant component of FTS product costs, and these prices fluctuate significantly. Recent natural gas prices are shown in table 2.
At a natural gas price of $0.50/MMBtu (inexpensive), the feedstock accounts for $5/BBL of the product price. Operating costs (fixed and variable) are estimated to contribute an additional $5/BBL to the cost of production (Gradassi, 1998; Jager, 1998; Lutz, 2001; Vosloo, 2001).
Assuming a 50,000 BPD plant, an internal rate of return on the capital investment of 15%, and a natural gas cost of $0.50/MMBtu (inexpensive), the estimated cost of producing FT liquids is $26/BBL ($0.83/gallon). Under these assumptions, the natural gas to FT liquids process can be competitive for the production of transportation fuels with a crude oil price of $18/BBL (Gradassi, 1998). Gray and Tomlinson (1997) estimate that for a 50,000 BPD facility producing FT liquids, and a natural gas cost of $0.5/MMBtu (inexpensive), the cost of producing FT liquids is $24/BBL ($0.76/gallon) and is $52/BBL ($1.7/gallon) if the price of the natural gas is $4/MMBtu. They estimated the cost of production to be $46/BBL ($1.5/gallon) if Illionois #6 coal was used as the feedstock. For a once-through plant that co-produces electricity (which is sold at $0.05/kWh), the cost of producing the FT liquid is reduced to $35/BBL ($1.1/gallon). Three studies have examined the production of FT liquids from biomass feedstocks and are summarized in table 3.
Compared to conventional fuels, FT fuels contain no sulfur and low aromatics. These properties along with a high cetane number result in superior combustion characteristics (Jager, 1998; Alleman, 2003). Tests performed on heavy duty trucks showed decreased vehicle emissions of HC, CO, NOx, and PM when using a FT fuel (Haid, 2000; Lutz, 2001). FT diesel has been tested in a variety of light- and heavy-duty vehicles and engines. FT diesel fuel properties and emission information are summarized in Alleman (2003). Overall, FT diesel showed a reduction in regulated as well as some unregulated emissions compared to conventional diesel. Life cycle assessments (LCA) have been performed on a variety of transportation fuels including FT diesel and gasoline (General Motors, 2001; Marano, 2001; and MIT, 2000). Most studies examined only greenhouse gas emissions and energy consumption, but the General Motors study also examined five criteria pollutants (VOCs, CO, NOx, PM10, and SOx). Emission and energy consumption estimates vary based on feedstock type and procurement, technology, and vehicle assumptions. In general, FT liquids produced from fossil fuel feedstocks offer little energy and greenhouse gas emission benefits compared to traditional transportation fuels. Use of biomass feedstocks has the potential to alter this situation, and because of the improved combustion characteristics of the FT liquids, a complete LCA including criteria pollutants will mostly likely show overall benefits of FT liquids compared to conventional transportation fuels. |
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