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Substitute Natural Gas and Fischer-Tropsch Synthesis
CHAPTER
13
Substitute Natural Gas and Fischer-Tropsch Synthesis Contents Overview Substitute Natural Gas SNG Processes Competition Between Coal Gasification and Natural Gas SNG as a Carrier of Coal Energy SNG Versus Synthetic Liquid Hydrocarbons Fischer-Tropsch History Fischer-Tropsch Chemistry FT reactor design Refining Fischer-Tropsch Fluids Diesel Production Naphtha and Gasoline Lubricating Base Oils Outlook for FT Fluid Refining FT Economics References
373 373 375 377 378 378 379 379 381 385 385 386 387 388 389 390
OVERVIEW This chapter covers the direct conversion of syngas to hydrocarbons, either substitute natural gas, or higher hydrocarbons via Fischer-Tropsch synthesis. In the prior chapter, an indirect route via methanol and dimethyl ether synthesis was considered. Fischer-Tropsch synthesis produces a synthetic crude oil, and the unique properties of this oil require special considerations during refining to produce petroleum products.
SUBSTITUTE NATURAL GAS Substitute natural gas (SNG), also known as synthetic natural gas, is methane. Methane is formed during gasification, and Chapter 3 briefly described gasification processes designed to produce methane as their primary product.1,2 The usual approach, however, is the catalytic hydrogenation of carbon monoxide: COðgÞ þ 3H2 ðgÞ/CH4 ðgÞ þ H2 OðgÞ
DH rxn ¼ 206:2 kJ=gmole
Coal Gasification and Its Applications. ISBN B978-0-8155-2049-8.10013-0, doi:10.1016/B978-0-8155-2049-8.10013-0
R-13.1
Ó 2011 Elsevier Inc. All rights reserved.
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The syngas will generally contain carbon dioxide, and this, too, can be hydrogenated to produce methane: CO2 ðgÞ þ 4H2 ðgÞ/CH4 þ H2 OðgÞ DH rxn ¼ 165:0 kJ=gmole
R-13.2
SNG catalysts, usually nickel, also have significant water gas shift activity: COðgÞ þ H2 OðgÞ4CO2 ðgÞ þ H2 ðgÞ
DH rxn ¼ 41:21 kJ=gmole
R-13.3
(a)
Equilbirum Conversion of CO,
Figure 13.1a shows the equilibrium conversion of carbon monoxide to methane via R-13.1 as a function of temperature and pressure. Conversion falls with increasing temperature. Reaction R-13.1 shows that four moles of gas react to form two moles of gas, so increasing reaction pressure pushes the reaction to the right. In all of the equilibria calculated for Figure 13.1a, the conversion of carbon monoxide was 99.2% or higher.
100 99 98 97 96 95 94 93 92 91 90 260
5 MPa 2 MPa 1 MPa
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(b)
Equilibrium Conversion of CO,
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1 MPa
7 2 MPa
6 5
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4 3 2 1 0 260
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Temperature, C
Figure 13.1 (a) Equilibrium conversion of carbon monoxide to methane via R-13.1. Most of the drop in conversion with increasing temperature is due to conversion of carbon monoxide to carbon dioxide via the water gas shift reaction (R-13.3). A stoichiometric 3:1 hydrogen to carbon monoxide feed ratio was used in these calculations. (b) Equilibrium conversion of carbon monoxide to carbon dioxide via the water gas shift reaction (R-13.3) during the conversion of carbon monoxide to methane via R-13.1.
Substitute Natural Gas and Fischer-Tropsch Synthesis
The reason for the decline in conversion of carbon monoxide to methane with increasing temperature is the competing water gas shift reaction, R-13.3. Figure 13.1b shows how the equilibrium carbon dioxide concentration increases as a result of the water gas shift reaction with increasing reaction temperature. The conversion of hydrogen was as high as 99.6% at the minimum temperature, 260 C, and the maximum pressure, 5 MPa, and as low as 87.9% at the maximum temperature, 500 C, and the minimum pressure, 1 MPa. Figure 13.2 shows the equilibrium conversion of carbon dioxide to methane via R-13.2. Methane yields are a little lower for carbon dioxide compared to carbon monoxide at the same reaction conditions. For carbon dioxide, a lower yield of byproducts are obtained. Less than 0.9% of the carbon dioxide is converted to carbon monoxide via the reverse water gas shift reaction, R-13.3.
SNG PROCESSES The Great Plains Synfuels plant,3 in Beulah, North Dakota, started operation in 1984. Lignite from an adjacent mine is crushed and fed to Lurgi gasifiers. Since these gasifiers cannot process fine coal, fines are sent to an adjacent pulverized coal combustion power plant. Syngas is cleaned in a Rectisol unit, also a Lurgi process. Carbon dioxide from the Rectisol unit is pipelined to Weyburn, Saskatchewan, where it is used for enhanced oil recovery. Cleaned syngas from the Rectisol unit is sent to a Lurgi methanation unit, where it is converted to SNG. The catalyst in the methanation unit is gradually poisoned by sulfur, so the catalyst lifetime is dictated by operation of the Rectisol unit, which removes the sulfur.
Equlibrium Conversion of CO2,
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Figure 13.2 Equilibrium conversion of carbon dioxide to methane via R-13.2. The feed for these calculations was a stoichiometric 4:1 hydrogen to carbon dioxide ratio.
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H2, CO, CO2
reactor 1
reactor 2
reactor 3
recycle gas
steam cooler
water
steam water
boiler feed water preheater cooler methane
recycle compressor water
Figure 13.3 Haldor Topsoe TREMPÔ (Topsoe Recycle Energy-efficient Methanation Process) for substitute natural gas.
Figure 13.3 shows a methanation unit design by Haldor Topsoe, called TREMPÔ (Topsoe Recycle Energy-efficient Methanation Process).4 A nearly stoichiometric blend of hydrogen, carbon monoxide, and carbon dioxide is mixed with recycle gas and fed to the first reactor. The effluent from this reactor is cooled, and then split into a recycle stream, and the feed is then sent to the second reactor. The recycle stream is further cooled in two exchangers. This gas is then fed to the recycle compressor. The recycle gas dilutes the feed gas, and reduces the temperature rise across the first reactor. A similar recycle gas cooling scheme is used in the Lurgi methanation process,5 in which the gas enters the first reactor at 300 C and leaves the first reactor at 450 C. The Lurgi methanator operates at about 1.8 MPa. Unlike ammonia or methanol synthesis, there is no separation of unreacted feed gas and product. With SNG, this would require an expensive cryogenic distillation process. Instead, the TREMPÔ SNG process is designed to achieve high single pass conversion. This is done by a series of three reactors, each reactor operating at a lower temperature than the previous reactor. As can be seen in Figures 13.1a,b, and 13.2, lower operating temperatures favor the formation of methane, and reduce the formation of carbon dioxide byproduct. Table 13.1 shows typical gas compositions and heating values for SNG produced by the TREMPÔ process.4
Substitute Natural Gas and Fischer-Tropsch Synthesis
Table 13.1 Typical gas composition and heating value for substitute natural gas produce by the TREMPÔ process.4 Component Mole %
Methane Carbon dioxide Hydrogen Carbon monoxide Nitrogen þ argon Higher heating value, KJ/Nm3
94e98 0.2e2 0.5e2 <100 ppm 2e3 37,380e38,370
COMPETITION BETWEEN COAL GASIFICATION AND NATURAL GAS Coal gasification and natural gas have a long history of commercial competition. The first coal gasifiers produced town gas for household energy use. These gasifiers were shut down when natural gas and electric power became widely available. The first ammonia and methanol plants were designed to use syngas from coal gasifiers. Now, most ammonia and methanol is made from natural gas-derived syngas. In the late 1960s through the early 1980s, there was widespread concern that North America was about to exhaust its natural gas reserves. This led to the construction of the Great Plains Synfuels plant. Natural gas reserves were not near exhaustion, and in the mid 1980s the price of natural gas fell. The Great Plains Synfuels plant was uneconomic, and the US government intervened to keep it operating. In the early 2000s rapidly rising natural gas prices had a number of companies considering ammonia and methanol production from coal, but natural gas prices fell again. When this book was written, the outlook was for abundant natural gas supplies and little likelihood of large price increases. Two developments are responsible for this situation. First, consider that natural gas markets traditionally are regional, rather than global. Regional producers are connected to regional consumers via a pipeline network. The USA imported most of its crude oil in 2009, but the USA produced 92% of the natural gas that it consumed that year. The imported gas was almost entirely from Canada, and shipped to the USA via pipeline. A similar regional market exists in Europe, with producers in the North Sea, Russia, and northern Africa. In some regions, such as east Asia, the potential consumption of natural gas is greater than the regional supply. In other regions, such as the Persian Gulf, regional natural gas reserves are large compared to regional consumption. One of the two major developments that affected the natural gas outlook recently is liquefied natural gas (LNG), which is converting the natural gas market into a global market. In regions with an abundance of natural gas, such as the Persian Gulf, natural gas is liquefied and loaded onto transoceanic tankers. It is then shipped to regions where local demand exceeds supply, such as east Asia, where it is vaporized and put into the
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local pipeline network. No great technological breakthrough was necessary to develop this market. Instead, sustained price differentials between supplying regions and consuming regions made LNG profitable. The second major development is a combination of two technical advances. The first is directional drilling, which allows the driller to follow the length of natural gas-bearing reservoir rock, rather than merely drilling a vertical hole through it. The second technical development is hydraulic fracturing, which fractures reservoir rock around the well bore. These two developments allow drillers to create large drainage areas in each well. Now, natural gas production from low permeability rocks, especially shale, is economically feasible. Huge volumes of shale in North America have been recently re-classified as natural gas reserves. Shale gas production, now common in North America, is expected to spread to other continents. So where does coal gasification fit in? Well, for many applications, such as ammonia and methanol production, the abundance of natural gas put gasification plans on hold. Uses of natural gas are expanding, such as an increase in natural gas for electric power production, as opposed to coal combustion, to reduce greenhouse gas production, and the potential use of natural gas as a transportation fuel to replace scarce petroleum. Of course, natural gas will not last forever, and SNG is a valuable back-up supply of methane when natural gas supplies diminish.
SNG AS A CARRIER OF COAL ENERGY Currently, most coal is burned to produce electric power. Coal is hauled over land in trains, and a large fraction of a power plant’s delivered coal cost is transportation cost. Natural gas, on the other hand, may be transported long distances at relatively low cost in pipelines. Compared to coal, natural gas combustion produces much less carbon dioxide per unit of energy. Rather than shipping coal in a train, coal could be converted to SNG in a plant near the coal mine. By-product carbon dioxide could be locally sequestered. The SNG could then be shipped via pipeline to power plants and other consumers that might otherwise burn coal. This approach will lower transportation costs and greenhouse gas emissions.
SNG VERSUS SYNTHETIC LIQUID HYDROCARBONS Traditionally, liquid petroleum products were more expensive than natural gas on an equivalent energy basis. This is why burning oil to produce heat or electricity is generally limited to regions without natural gas supplies. Lately, the spread between oil and natural gas prices has widened considerably. If current projections about crude oil supply and demand come true, that gap will widen further.
Substitute Natural Gas and Fischer-Tropsch Synthesis
As shown in the previous chapter, liquid hydrocarbons can be made from syngas via methanol and dimethyl ether. This chapter discusses a second, better-known approach which is the Fischer-Tropsch synthesis. If a decision between converting syngas to SNG or liquid fuels needs to be made, one should remember that the liquids are worth much more.
FISCHER-TROPSCH HISTORY During the 1920s, Franz Fischer and Hans Tropsch, working at the Kaiser Wilhem Institute in Berlin, developed a process for the direct conversion of syngas to hydrocarbons. This is known today as the Fischer-Tropsch (FT) process. During World War II, Germany, cut off from crude oil suppliers, produced synthetic oil using the FT process and the Bergius process, which is a direct coal hydrogenation process (see Chapter 2). Today, the Fischer Tropsch Archive5 maintains an electronic collection of early FT process documents. Sasol, using Lurgi technology, started production of FT fluids in 1955 from locally mined coal in Sasolburg, South Africa.6 A second FT plant in Secunda, South Africa, started operations in 1980. Much of the FT process technology literature is based on Sasol’s South African experience. Shell started their SMDS (Shell middle distillate synthesis) plant in Bintulu, Malaysia in 1993. This plant uses natural gas as the feedstock, and the FT fluids are hydrocracked. The Oryx GTL (gas to liquids) plant, a joint venture between Qatar Petroleum and Sasol, started operations in 2006. The Pearl GTL plant, a joint venture between Qatar Petroleum and Shell, is expected to be the world’s largest gas to liquids plant when it reaches its full 140,000 bbl/day liquids production in 2011.
FISCHER-TROPSCH CHEMISTRY Fischer-Tropsch synthesis is a polymerization reaction. An alkane chain, absorbed on a catalyst site ()), reacts with carbon monoxide and hydrogen to add a methyl unit (CH2): H-ðCH2 Þn- þCO þ 2H2 /H-ðCH2 Þn CH2- þ H2 O
R-13.4
The heat of reaction per methyl group, estimated from the gas phase heat of formation for dodecane (C12H26), is 155.5 kJ/mole. This heat of reaction is about 20% of the heat of combustion of the feed gases. Iron and cobalt catalysts are typically used, although other metals, such as ruthenium and nickel, are also active. Iron has significant water gas shift activity (R-13.3), and typical FT synthesis conditions favor the right side of the water gas shift reaction. Combining R-13.3 with R-13.4 for an iron catalyst gives: H-ðCH2 Þn- þ2CO þ H2 /H-ðCH2 Þn CH2- þ CO2
R-13.5
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The heat of reaction per methyl group, again based on dodecane, is 196.7 kJ/mole. This is somewhat higher than R-13.4 due to the exothermic contribution of the water gas shift reaction. This heat of reaction is about 24% of the heat of combustion of the feed gases. An examination of R-13.4 and R-13.5 shows that, for a cobalt catalyst, the ideal feedstock has a 1:2 carbon monoxide to hydrogen ratio, while, for an iron catalyst, the ideal feedstock has a 2:1 carbon monoxide to hydrogen ratio. With a typical syngas from a coal gasifier, a water gas shift reactor is required for a cobalt catalyst, but not required for an iron catalyst. The water gas shift reactor, could, however, shift part of the heat removal from an iron FT synthesis reactor to the water gas shift reactor, and, as we will see later, this could improve operations. The absorbed alkane chain may desorb, halting growth. The usual desorption product is a normal alkane, but alcohols and olefins also form. This means that the final product has a range of molecular weights, ranging from methane to very high molecular weight alkanes. The choice of catalyst and operating conditions determine the distribution of molecular weights. Anderson7 adapted a model developed for other types of polymerization to FT synthesis. The result is known as the Anderson-Schultz-Flory (ASF) distribution. A single adjustable parameter, a, describes the probability that an alkane absorbed on the catalyst, H-(CH2)n-), will add a methyl unit to form H-(CH2)nþ1-). Conversely, (1a) is the probability that the absorbed alkane will desorb, halting growth. The final weight fraction of each alkane compound (using the approximation that only alkanes are formed) is: Eqn. 13.1 W ¼ nan-1 ð1-aÞ2 n
Variations on the ASF distribution equations are used to improve the fit to experimental measurements of molecular weight distributions. For example, olefins appear to reabsorb and re-initiate growth, so high molecular weight alkanes are a little more abundant than predicted by Eqn. 13.1. Except for a relatively small quantity of hydrocarbon solvents, petroleum products are not sold as pure compounds. Instead, petroleum products are sold as hydrocarbon mixtures that are primarily distinguished by their boiling point ranges. There can be considerable overlap between adjacent products. For example, some of the higher molecular weight compounds in gasoline can also be some of the lower molecular compounds in diesel. Still, it is useful to approximate the FT fluid distillation product yield using molecular weight ranges. The approximate distillation yield for an FT fluid versus the chain growth probability factor, a, is shown in Figure 13.4. The catalyst, reactor design, and operating conditions are all optimized to maximize the value of the hydrocarbon products. In a typical petroleum refinery, the refinery may be operated to maximize the yield of naphtha, which is further processed to make
Substitute Natural Gas and Fischer-Tropsch Synthesis
100
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Figure13.4 Product distillation yield for a Fischer-Tropsch liquid as a function of a, the chain growth probability factor. Carbon numbers are used to estimate distillation boiling ranges. Yield versus carbon number was estimated using Eqn. 13.1.
gasoline. FT naphthas, on the other hand, tend to have very low octane numbers, while FT diesels tend to have very high cetane numbers. Lamprecht8 reports a 62 cetane number for an FT diesel, versus a 40 minimum cetane specification9 for US diesel. Both the excellent FT diesel cetane number and the poor FT naphtha octane number are for the same reason, a high normal alkane content. Therefore, FT plants tend to maximize diesel production. The least valuable fraction is the fuel gas. In a typical oil refinery, this gas is burned to provide heat for refinery operations. The next least valuable fraction is LPG. This may be recovered and sold, but the price is substantially less than for gasoline, jet fuel, or naphtha. For example, in 2009, the average US price for propane was $1.236/ gallon, versus $1.888, $1.704, and $1.833 per pre-tax gallon for gasoline, jet fuel, and diesel, respectively.10 Many refineries have dimerization processes that convert LPG to naphtha. Demand for paraffin wax is small, so the usual approach is to hydrocrack FT wax to produce lower molecular weight hydrocarbons. To maximize production of diesel, the FT reactor is designed to operate with a high chain growth probability factor, a, and then the wax is hydrocracked.
FT REACTOR DESIGN Lu11 reviewed a number of FT experimental studies; and developed approximate correlations between the chain growth probability factor, a, and reactor temperature and pressure. The most prominent trend is that a tends to decline with increasing
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temperature. This means that the operating temperature is a compromise between good reaction rates, they increase with increasing temperature; and high molecular weight distributions, which increase with decreasing temperature. A further challenge is that FT synthesis is highly exothermic, so the reactor must be designed to remove a large quantity of heat to maintain the reactor temperature within an optimum operating temperature range. Figure 13.5 shows the Sasol Synthol circulating fluid bed reactor. Fresh and recycled gas are mixed with a hot catalyst stream and fed to a vessel containing cooling bundles. The catalyst/gas mixture is then fed to a disengagement vessel where product gas is taken off the top, through cyclones to remove fine catalyst solids. The catalyst flows through a standpipe on the bottom of the disengagement vessel. The first Sasol Advanced Synthol reactor12 started production in Sasolburg in 1989. This reactor, shown Figure 13.6, is a bubbling fluidized bed reactor with heat transfer tubes in the catalyst bed. Compared to the original Synthol reactor, the Advanced Synthol reactor is more compact for a similar production capacity. The Synthol reactors operate at a fairly high temperature, 330 to 350 C. Sie and Krishna13 analyzed fluidized bed FT reactors, such as the Synthol reactors. They noted that because of the small catalyst diameter, 100 mm, there are no intraparticle diffusional
catalyst + gas product gas cyclones
cooling bundles catalyst
Fresh gas + recycle
catalyst + gas
Figure 13.5 The Sasol Synthol circulating fluid bed reactor.
Substitute Natural Gas and Fischer-Tropsch Synthesis
product gas steam
water syngas
Figure 13.6 The Sasol Advanced Synthol reactor is a bubbling fluidized bed reactor with heat transfer tubes in the catalyst bed.
restrictions. The reactor temperature and chain growth factor, a, are chosen to prevent the condensation of liquid products in the catalyst bed. The Synthol reactor, shown in Figure 13.5, was one of two original reactor types used by Sasol. The other type is the ARGE reactor, shown in Figure 13.7. This reactor is built like a shell-and-tube heat exchanger, with catalyst packed in the tubes and coolant circulating through the shell side. Both gaseous and liquid hydrocarbons are formed, so syngas is fed to the top of the reactor and liquid wax trickles down through the packed
feed gas
steam
water product gas
wax
Figure 13.7 The Sasol ARGE reactor. Catalyst is packed in the tubes, and coolant circulates on the shell side.
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Table 13.2 Fischer-Tropsch product distribution for iron-based catalysts.14 Component, % Low temperature (220e250 C) High temperature (330e350 C)
CH4 C2-C4 olefins C2-C4 alkanes Naphtha Distillate Oils and waxes Oxygenates
4 4 4 18 19 48 3
7 24 6 36 12 9 6
catalyst tube. There are two product ports on the bottom of the reactor, one for liquid wax and the other for gas. The ARGE reactor is operated at about 220 to 250 C, a much lower temperature than the Synthol reactor. This results in a higher chain growth factor, a, leading to a heavier product, as shown14 in Table 13.2 for iron-based catalysts. Cobalt catalysts tend to give higher a values,15 which lead to higher molecular weight products. The Shell SMDS reactor in Bintulu, Malaysia uses a packed tube configuration similar to the ARGE reactor. More recently, slurry bubble reactors have been used for low temperature FT synthesis. The reaction media is a three phase churning mixture of solid catalyst, liquid wax product, and feed and product gas. Heat exchange tubes in the mixture remove the heat of reaction. Catalyst particle diameters14 are between 20 and 200 mm. The lower diameter limit is set by the difficulty of removing smaller solids from the product wax, and the upper diameter limit is set by the need to keep the catalyst particle suspended in the reaction media. There have been reports of operational difficulties with slurry bubble reactors. Figure 13.8 shows a microchannel reactor.16 This reactor is created by etching parallel grooves in flat metal plates. The plates are stacked to form flow channels bounded by
Figure 13.8 A pilot scale microchannel reactor for Fischer-Tropsch synthesis built by Velocys, Inc.
Substitute Natural Gas and Fischer-Tropsch Synthesis
adjacent plates. Alternating layers of flow channels are filled with catalyst. The channels between the catalyst channels contain a heat transfer fluid, flowing at right angles with respect to the reactant flow. Conceptually, this reactor is much like the packed tube ARGE reactor. The microchannel reactor, however, has a much higher heat transfer area/catalyst volume ratio and a much smaller flow channel cross section. Both channels allow for a much better control of the reaction temperature range.
REFINING FISCHER-TROPSCH FLUIDS FT fluid can be thought of as a synthetic crude oil. As such, it is an unusual crude oil that requires special refining considerations. A typical crude oil contains normal alkanes, branched and cyclic alkanes, aromatics, sulfur compounds, and nitrogen compounds. Unless the crude oil is produced using thermal techniques, crude oil contains essentially no olefins. Oxygen compounds are present in crude oil, but usually at low levels. FT fluid, on the other hand, consists primarily of normal alkanes, with lesser but significant levels of olefins and alcohols. FT fluid contains very few branched alkanes, and essentially no cyclic compounds, sulfur or nitrogen compounds. We17,18 modeled FT refineries using a pure component approach and the Aspen Plus process simulation software. Three scenarios were considered in this study; a refining scheme that maximized diesel production, another, very similar scenario that maximized jet fuel production, and a third refining scheme to produce lubricating base oils.
Diesel Production FT diesel fuel, due to its high alkane content, has very high cetane numbers. Blends of FT and conventional diesel are sold as premium products in Europe. Normal alkanes, however, have high melting points compared to branched or cyclic hydrocarbons. Normal hexadecane, for example, has a melting point of þ18 C, versus 23 C for its branched isomer, 3-methyl-pentadecane. An FT diesel produced simply by distilling an FT fluid would have poor low temperature properties. This may be one of the reasons why FT diesel is sold in Europe as a blend with conventional diesel, rather than as a pure product. Most FT diesel, however, is not simply an FT fluid distillation product, but a blend of this distillation product and products of hydrocracked FT wax. Sie19 studied the hydrocracking of normal alkanes, and presented evidence for a cyclopropane intermediate mechanism. As a result of this mechanism, propane is the smallest fragment produced by a hydrocracking reaction. This reaction converts a normal paraffin molecule into two smaller molecules. One of these is also a normal alkane, while the other is a branched alkane. This shows a close relationship between hydrocracking and isomerization. As an approximation, any of the carbonecarbon bonds within the molecule,
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except for the two closest to the ends, have an equal probability of fracturing. Sie’s data actually show that the bonds towards the center of the molecule are slightly more likely to fracture than bonds near the end of the molecule. Leckel and co-workers20-23 published several papers describing Sasol FT wax hydrocracking pilot plant experiments. The FT wax is hydrocracked and then distilled. The distillation bottoms, which are unconverted wax, are recycled to extinction. Hydrocracking severity (conversion) can be controlled using a number of process parameters, including catalyst acidity, reactor temperature, reactor pressure, and reactor space velocity. Leckel showed that low temperature properties of diesel improve with increased hydrocracker severity, which is expected since hydrocracking creates branched alkanes. Increased hydrocracking also resulted in lower diesel yields, since a larger fraction of the hydrocracked product was naphtha or lighter products. Thus, there is a close relationship between diesel yield and low temperature properties. In our diesel model, we distilled an FT fluid produced by a cobalt catalyst with a high chain growth factor, a. Both the wax and the diesel fractions were hydrocracked. The diesel fraction did not need to be hydrocracked to meet the boiling point specifications. Rather, this was done to increase the branched alkane/normal alkane ratio in the final diesel product. Sie’s rules were used to model the hydrocracker; and the extent of hydrocracking was arbitrarily adjusted so that our simulation results fit Leckel’s experimental results. Although we were able to estimate the diesel composition, we did not find a reliable means of predicting low temperature properties based on that composition. In the USA, diesel cloud points are specified9 according to the time of year and the local climate. The cloud point is the temperature at which wax crystals become visible as the sample is cooled. Below the cloud point, these wax crystals plug fuel filters. Krishna et al.24 offer an approximate method of predicting pour points based on composition. Pour points are typically a few degrees lower than cloud points. Leckel’s results show that FT diesel can meet low temperature specifications, but we are not currently able to reliably predict diesel yields without experimental data.
Naphtha and Gasoline Naphtha processing proved to be challenging, due to its low octane numbers for FT naphtha. Brent et al.25 investigated FT naphtha upgrading techniques. They approached two naphtha reforming process providers, and were told by both that FT naphtha was not a good reforming feedstock. Naphtha reforming increases octane numbers by increasing aromatic content, primarily through dehydrogenation of cycloalkanes. Instead, they chose to sell FT naphtha as an olefin production feedstock. FT naphtha, compared to other naphthas, is an excellent olefin production feedstock because of its high alkane content. There are two problems with this approach. The first is that this would put FT
Substitute Natural Gas and Fischer-Tropsch Synthesis
naphtha in direct competition with natural gas liquids, which sell at substantially lower prices than gasoline. The second is that olefin plants designed to use naphtha feed are primarily located in countries that do not have significant natural gas production. In our model, we sought to convert FT naphtha to gasoline. To do this, we used two processes. The first is the TIP (total isomerization process) process,26 in which the octane number is increased by converting normal alkanes into branched alkanes. The isomerization reactor effluent is fed to a mole sieve bed, where normal alkanes are selectively adsorbed and recycled to the isomerization reactor. We used Sie’s hydrocracking rules to model the isomerization reactor, although catalysts with better isomerization/hydrocracking selectivity are available. The second process is an alkylation process that converts C4 compounds to isooctane. The problem with this refining approach was that octane improvement was accompanied by molecular weight reduction due to hydrocracking in the isomerization reactor, and we could not simultaneously meet both octane specifications and boiling point specifications. A more selective isomerization catalyst would improve the performance of this refining approach. Just as steam reforming of natural gas is a common method of producing syngas, LPG and naphtha can also be steam reformed to produce syngas.27 If a suitable outlet for FT naphtha or LPG is not found, then these streams can be recycled, thereby increasing the fractional yield of diesel or jet fuel. Our jet fuel model was very much like our diesel model, with the boiling point range adjusted to meet jet fuel standards. When jet fuel production is maximized, the diesel yield is insignificant.
Lubricating Base Oils We also examined a refinery model to maximize the production of lubricating base oils. This is the largest volume non-fuel petroleum market. A typical lubricating oil contains about 90% base oil and about 10% of an additive package. Branched alkanes, primarily in the C20 to C40 range, offer the best combination of volatility, viscosity, and low temperature performance for base oils. Traditionally, base oils with high branched alkane content are made using a paraffinic crude oil. A vacuum distillate cut is processed to increase the alkane content, either by solvent extraction of aromatics, or by converting aromatics to alkanes in a catalytic hydrogenation reactor. The oil is then dewaxed (high melting point normal alkanes are removed) by mixing the oil with a blend of methyl ethyl ketone and toluene, chilling the oil to crystallize wax, and then filtering the wax/ oil/solvent slurry. The solvent dewaxing process is the most expensive step in the traditional base oil production process. The traditional base oil process produces a complex and variable mixture of hydrocarbons. With increasingly stringent lubricating oil standards, producers found it difficult to meet standards with these traditional base oils. This led to the production of
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synthetic oils, such as polyalpha olefins, in which the hydrocarbon molecular structure can be more closely controlled to optimize oil properties. The cost of synthetic base oils is much higher than that of traditional base oils. In the Chevron Isodewaxing process28 normal alkanes are isomerized to branched alkanes. The pore size of the catalyst allows methyl branches to form, but the pores are too small to readily allow the formation of the cyclopropane intermediate compound that leads to hydrocracking. As a result, this catalyst has excellent isomerization/ hydrocracking selectivity. The Isodewaxing process was initially marketed as an inexpensive alternative to solvent dewaxing, but the isomerized waxes have physical properties that are superior to traditional base oils. When the feedstock is a paraffin wax, the base oil product is nearly as good as polyalpha olefins. This process offers a low cost route to high quality lubricants. Partly in response to this process, the American Petroleum Institute developed a classification system29 for base oils. Traditional base oils are either Class I or Class II, depending on their sulfur content. The higher quality base oils produced by wax isomerization are Class III. Poly alpha olefins are Class IV, and other types of synthetic base oils are Class V. In our lubricating base oil refinery model, the FT wax is isomerized using the Isodewaxing process. Detailed technical information on this process in the open literature is scarce, so we modeled the isomerization reactor based on pure component yield data from Claude et al.30-31 for a different catalyst, Pt/H-ZSM-22. This catalyst has a pore geometry that is similar to the Chevron catalyst. The yield data were for normal alkane feeds between C10 and C24. Yield data were extrapolated to higher molecular weights for this simulation. Three different viscosity grade oils were made. The base oil refinery model is attractive, in part, because base oils are worth considerably more than fuels. In March, 2010, US Gulf Coast Group III base oils sold for about $3.50/gallon,32 about twice the price of diesel.
Outlook for FT Fluid Refining In our refinery models, we sought, to use conventional petroleum refining processes. The performance of the fuel refinery cases could be improved with the application of an isomerization catalyst with high isomerization/hydrocracking selectivity, such as the Chevron Isodewaxing catalyst used for base oil production. A combination of this catalyst and hydrocracking would give the refiner independent controls over molecular weight and low temperature properties, and this could greatly enhance diesel or jet fuel yields. It is not clear whether gasoline octane specifications can be met, but a higher octane is achievable if better isomerization selectivity is attained in the naphtha isomerization reactor. As a stand-alone process for liquid hydrocarbon production from syngas, FischerTropsch technology competes with methanol-based processes, such as the ExxonMobil
Substitute Natural Gas and Fischer-Tropsch Synthesis
MTG process discussed in the previous chapter. The MTG process is attractive because it produces specification grade gasoline in high yield with minimal processing. Fischer-Tropsch technology may be better suited as a companion technology for heavy oil refining. The trend in petroleum refining is away from light crude oils and towards heavy crudes. This produces much more aromatic products. High aromatic content in gasoline boosts octane numbers, but aromatic contents are limited due to toxicity concerns about aromatics, especially benzene. Aromatics tend to lower diesel cetane numbers. FT products are nearly free of aromatics, when blended with products from heavy crudes, the blends more closely resemble the mixture of alkanes and aromatics that are normally associated with refined products made from light crudes.
FT ECONOMICS Van Bibber et al.33 prepared a process and economic study of a Fischer-Tropsch plant in Healy, Alaska, based on Usebelli coal, a locally mined sub-bituminous coal. Their plant would produce 14,640 bbl/day of FT fluids, of which 58 vol.% would be diesel and 42 vol.% would be naphtha. Diesel prices were assumed to be the same as conventional diesel prices, rather than charge a premium for the higher cetane FT diesel. They also assumed that they could sell the FT naphtha for a $0.10 discount relative to gasoline, rather than dealing with naphtha upgrading issues. We34 modified this report for a Wyoming Powder River Basin (PRB) location. This is appropriate because PRB coal is very similar to Usebelli coal. To account for the higher elevation of the PRB site, another compressor was added to the air separation unit. Capital costs were updated to 2008. Locally, there is a demand for carbon dioxide for enhanced oil recovery, so we assumed that we could sell carbon dioxide for $1.00 per million standard cubic feet. This assumption had little impact on plant economics. With carbon capture and sequestration, the plant would cost about $1.9 billion. The FT fluid cost would have to be $89/bbl for a 10% internal rate of return. This is close to the pre-tax price of gasoline and diesel when this was written. Figure 13.9 is a tornado
Figure 13.9 Effect of a 25% change in assumed economic parameters on the percent internal rate of return for a Fischer-Tropsch coal-to-liquids plant with carbon capture and sequestration in the Powder River Basin of Wyoming. For the base case, $89/bbl is the assumed product price.
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plot showing how a 25% change in assumed economic parameters affects the internal rate of return. Not surprisingly, the plant economics are sensitive to product price. The process economics are dominated by capital construction costs, so changes in capital costs and plant availability have a strong impact on profitability. Coal costs, on the other hand, have a minor effect. The average mine-mouth price of Wyoming coal in 2008 was $11.39/ton, so a 25% change in this price does not have a large impact on operating costs. The price of carbon dioxide has little impact on the overall economics.
REFERENCES 1. Nahas NC. Exxon catalytic coal gasification process, fundamentals to flowsheets. Fuel. 1983;62: 239-241. 2. Great Point Energy. Our technology.; 2008. 3. Panek JK, Grasser J. Practical experience gained during the first twenty years of operations of the great plains gasification plant and implications for future projects. U.S. Dept. of Fossil Energy, Office of Fossil Energy, ; 2006. 4. Haldor Topsoe. From solid fuels to substitute natural gas using TREMPÔ. . 5. Fischer-Tropsch Archive. www.fischer-tropsch.org. 6. Sasol. 50 Years of innovation. . 7. Anderson RB. The Fischer-Tropsch Synthesis. Academic Press; 1984. 8. Lamprecht D. Fischer-Tropsch fuel for use by the U.S. military as battlefield-use fuel of the future. Energy & Fuels. 2007;21:1448-1453. 9. American Society for Testing and Materials. Standard specification for diesel fuel oils. ASTM D 975e05. 10. U.S. Energy Information Administration. tonto.eia.doe.gov/dnav/pet/pet_pri_refoth_dcu_nus_a.htm. 11. Lu X. Computer aided process design for Fischer-Tropsch conversion of coal to liquid fuels. M.S. thesis, University of Wyoming; 2003. 12. Duvenhage DJ, Shingles T. Synthol reactor technology development. Catal. Today. 2002;71:301-305. 13. Sie ST, Krishna R. Fundamentals and selection of advanced Fischer-Tropsch reactors. Appl. Catal. A: Gen. 1999;186:55-70. 14. Jager B. Fischer-Tropsch reactors, Presented at the Am. Inst. of Chemical Engineers Nat. Meeting, New Orleans; March 31-April 4, 2003. 15. Bartholomew CH. Recent developments in Fischer-Tropsch catalysis. Catal. Lett. 1990;7:303-316. 16. Kilanowski DR. Microchannel Fischer-Tropsch approaches market entry phase. Presented at the 4th BTLtec, Graz, Austria, Sept. 24e25, 2009. . 17. Bell DA, Costopolous SE. A Fischer-Tropsch fluids refinery. Presented at the Am. Inst. of Chemical Engineers Nat. Meeting, Salt Lake City, Utah; Nov. 2007. 18. Harkins SEC. Fischer-Tropsch fluids refinery. M.S. thesis, University of Wyoming; 2010. 19. Sie ST. Acid-catalyzed cracking of paraffinic hydrocarbons. 3. evidence for the protonated cyclopropane mechanism from hydrocracking/hydroisomerization experiments. Ind. Eng. Chem. Res. 1993;32:402-408. 20. Leckel D. Hydrocracking of iron-catalyzed Fischer-Tropsch waxes. Energy & Fuels. 2005;19:17951803. 21. Leckel D, Liwanga-Ehumbu M. Diesel-selective hydrocracking of an iron-based Fischer-Tropsch wax fraction (C15-C45) using a MoO3-modified noble metal catalyst. Energy & Fuels. 2006;20:2330-2336. 22. Leckel D. Selectivity effect of oxygenates in hydrocracking of Fischer-Tropsch waxes. Energy & Fuels. 2007;21:662-667.
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23. Leckel D. Low-pressure hydrocracking of coal-derived Fischer-Tropsch waxes to diesel. Energy & Fuels. 2007;21:1425-1431. 24. Krishna R, Joshi GC, Purohit RC, Agarwal KM, Verma PS, Bhattacharjee S. Correlation of pour point of gas oil and vacuum gas oil fractions with compositional parameters. Energy & Fuels. 1989;3: 15-20. 25. Brent FD, Shah L, Berry E, Schrader CH, Anderson J, Erwin J, et al. Early entrance coproduction plant, task phase II, task 2.5: Fischer-Tropsch product upgrading. Texaco Energy Systems, U.S. Dept. of Energy DE-FC26e99FT40658; 2004. 26. Symoniak MF. Upgrade naphtha to fuels and feedstocks. Hydrocarbon Proc. 1980;110-114. 27. Riddler DE, Twigg MV. Steam reforming. In: Twigg MV, ed. Catalyst Handbook. 2nd ed. Wolfe Publishing; 1989;225-282. 28. Krishna KR, Rainis A, Marcantonio PJ, Mayer JF, Biscardi JA, Zones SI. Next generation Isodewaxing and hydrofinishing technology for production of high quality base oils. Presented at the NPRA Lubricants and Waxes Meeting, Houston; Nov. 14-15, 2002. 29. American Petroleum Institute. Appendix E: API base oil interchangeability guidelines for passenger car motor oils and diesel engine oils.; 2009. 30. Claude MC, Martens JA. Monomethyl-branching of long n-alkanes in the range from decane to tetracosane on Pt/H-ZSM-22 bifunctional catalyst. J. Catal. 2000;190:39-48. 31. Claude MC, Vanbusele G, Martens JA. Dimethyl branching of long n-alkanes in the range from decane to tetracosane on Pt/H-ZSM-22 bifunctional catalyst. J. Catal. 2001;203:213-231. 32. Green CL. U.S. base oil price report. Lube Report. 2010;10. no. 10. . 33. Van Bibber L, Thomas C, Chaney R. Alaska coal gasification feasibility studies-Healy coal-to-liquids plant. DOE/NETL 2007/1251. ; 2007. 34. Schaffers WC, Bell DA. Economics of Wyoming coal gasification. Presented at the Int. Pittsburgh Coal Conf., Pittsburgh; Sept. 21-24, 2009.
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Substitute Natural Gas and Fischer-Tropsch Synthesis Contents Overview Substitute Natural Gas SNG Processes Competition Between Coal Gasification and Natural Gas SNG as a Carrier of Coal Energy SNG Versus Synthetic Liquid Hydrocarbons Fischer-Tropsch History Fischer-Tropsch Chemistry FT reactor design Refining Fischer-Tropsch Fluids Diesel Production Naphtha and Gasoline Lubricating Base Oils Outlook for FT Fluid Refining FT Economics References
373 373 375 377 378 378 379 379 381 385 385 386 387 388 389 390
OVERVIEW This chapter covers the direct conversion of syngas to hydrocarbons, either substitute natural gas, or higher hydrocarbons via Fischer-Tropsch synthesis. In the prior chapter, an indirect route via methanol and dimethyl ether synthesis was considered. Fischer-Tropsch synthesis produces a synthetic crude oil, and the unique properties of this oil require special considerations during refining to produce petroleum products.
SUBSTITUTE NATURAL GAS Substitute natural gas (SNG), also known as synthetic natural gas, is methane. Methane is formed during gasification, and Chapter 3 briefly described gasification processes designed to produce methane as their primary product.1,2 The usual approach, however, is the catalytic hydrogenation of carbon monoxide: COðgÞ þ 3H2 ðgÞ/CH4 ðgÞ þ H2 OðgÞ
DH rxn ¼ 206:2 kJ=gmole
Coal Gasification and Its Applications. ISBN B978-0-8155-2049-8.10013-0, doi:10.1016/B978-0-8155-2049-8.10013-0
R-13.1
Ó 2011 Elsevier Inc. All rights reserved.
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The syngas will generally contain carbon dioxide, and this, too, can be hydrogenated to produce methane: CO2 ðgÞ þ 4H2 ðgÞ/CH4 þ H2 OðgÞ DH rxn ¼ 165:0 kJ=gmole
R-13.2
SNG catalysts, usually nickel, also have significant water gas shift activity: COðgÞ þ H2 OðgÞ4CO2 ðgÞ þ H2 ðgÞ
DH rxn ¼ 41:21 kJ=gmole
R-13.3
(a)
Equilbirum Conversion of CO,
Figure 13.1a shows the equilibrium conversion of carbon monoxide to methane via R-13.1 as a function of temperature and pressure. Conversion falls with increasing temperature. Reaction R-13.1 shows that four moles of gas react to form two moles of gas, so increasing reaction pressure pushes the reaction to the right. In all of the equilibria calculated for Figure 13.1a, the conversion of carbon monoxide was 99.2% or higher.
100 99 98 97 96 95 94 93 92 91 90 260
5 MPa 2 MPa 1 MPa
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(b)
Equilibrium Conversion of CO,
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6 5
5 MPa
4 3 2 1 0 260
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380
400
420
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460
480
500
Temperature, C
Figure 13.1 (a) Equilibrium conversion of carbon monoxide to methane via R-13.1. Most of the drop in conversion with increasing temperature is due to conversion of carbon monoxide to carbon dioxide via the water gas shift reaction (R-13.3). A stoichiometric 3:1 hydrogen to carbon monoxide feed ratio was used in these calculations. (b) Equilibrium conversion of carbon monoxide to carbon dioxide via the water gas shift reaction (R-13.3) during the conversion of carbon monoxide to methane via R-13.1.
Substitute Natural Gas and Fischer-Tropsch Synthesis
The reason for the decline in conversion of carbon monoxide to methane with increasing temperature is the competing water gas shift reaction, R-13.3. Figure 13.1b shows how the equilibrium carbon dioxide concentration increases as a result of the water gas shift reaction with increasing reaction temperature. The conversion of hydrogen was as high as 99.6% at the minimum temperature, 260 C, and the maximum pressure, 5 MPa, and as low as 87.9% at the maximum temperature, 500 C, and the minimum pressure, 1 MPa. Figure 13.2 shows the equilibrium conversion of carbon dioxide to methane via R-13.2. Methane yields are a little lower for carbon dioxide compared to carbon monoxide at the same reaction conditions. For carbon dioxide, a lower yield of byproducts are obtained. Less than 0.9% of the carbon dioxide is converted to carbon monoxide via the reverse water gas shift reaction, R-13.3.
SNG PROCESSES The Great Plains Synfuels plant,3 in Beulah, North Dakota, started operation in 1984. Lignite from an adjacent mine is crushed and fed to Lurgi gasifiers. Since these gasifiers cannot process fine coal, fines are sent to an adjacent pulverized coal combustion power plant. Syngas is cleaned in a Rectisol unit, also a Lurgi process. Carbon dioxide from the Rectisol unit is pipelined to Weyburn, Saskatchewan, where it is used for enhanced oil recovery. Cleaned syngas from the Rectisol unit is sent to a Lurgi methanation unit, where it is converted to SNG. The catalyst in the methanation unit is gradually poisoned by sulfur, so the catalyst lifetime is dictated by operation of the Rectisol unit, which removes the sulfur.
Equlibrium Conversion of CO2,
100
95
5 MPa
2 MPa
90 1 MPa
85 260
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400
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Figure 13.2 Equilibrium conversion of carbon dioxide to methane via R-13.2. The feed for these calculations was a stoichiometric 4:1 hydrogen to carbon dioxide ratio.
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H2, CO, CO2
reactor 1
reactor 2
reactor 3
recycle gas
steam cooler
water
steam water
boiler feed water preheater cooler methane
recycle compressor water
Figure 13.3 Haldor Topsoe TREMPÔ (Topsoe Recycle Energy-efficient Methanation Process) for substitute natural gas.
Figure 13.3 shows a methanation unit design by Haldor Topsoe, called TREMPÔ (Topsoe Recycle Energy-efficient Methanation Process).4 A nearly stoichiometric blend of hydrogen, carbon monoxide, and carbon dioxide is mixed with recycle gas and fed to the first reactor. The effluent from this reactor is cooled, and then split into a recycle stream, and the feed is then sent to the second reactor. The recycle stream is further cooled in two exchangers. This gas is then fed to the recycle compressor. The recycle gas dilutes the feed gas, and reduces the temperature rise across the first reactor. A similar recycle gas cooling scheme is used in the Lurgi methanation process,5 in which the gas enters the first reactor at 300 C and leaves the first reactor at 450 C. The Lurgi methanator operates at about 1.8 MPa. Unlike ammonia or methanol synthesis, there is no separation of unreacted feed gas and product. With SNG, this would require an expensive cryogenic distillation process. Instead, the TREMPÔ SNG process is designed to achieve high single pass conversion. This is done by a series of three reactors, each reactor operating at a lower temperature than the previous reactor. As can be seen in Figures 13.1a,b, and 13.2, lower operating temperatures favor the formation of methane, and reduce the formation of carbon dioxide byproduct. Table 13.1 shows typical gas compositions and heating values for SNG produced by the TREMPÔ process.4
Substitute Natural Gas and Fischer-Tropsch Synthesis
Table 13.1 Typical gas composition and heating value for substitute natural gas produce by the TREMPÔ process.4 Component Mole %
Methane Carbon dioxide Hydrogen Carbon monoxide Nitrogen þ argon Higher heating value, KJ/Nm3
94e98 0.2e2 0.5e2 <100 ppm 2e3 37,380e38,370
COMPETITION BETWEEN COAL GASIFICATION AND NATURAL GAS Coal gasification and natural gas have a long history of commercial competition. The first coal gasifiers produced town gas for household energy use. These gasifiers were shut down when natural gas and electric power became widely available. The first ammonia and methanol plants were designed to use syngas from coal gasifiers. Now, most ammonia and methanol is made from natural gas-derived syngas. In the late 1960s through the early 1980s, there was widespread concern that North America was about to exhaust its natural gas reserves. This led to the construction of the Great Plains Synfuels plant. Natural gas reserves were not near exhaustion, and in the mid 1980s the price of natural gas fell. The Great Plains Synfuels plant was uneconomic, and the US government intervened to keep it operating. In the early 2000s rapidly rising natural gas prices had a number of companies considering ammonia and methanol production from coal, but natural gas prices fell again. When this book was written, the outlook was for abundant natural gas supplies and little likelihood of large price increases. Two developments are responsible for this situation. First, consider that natural gas markets traditionally are regional, rather than global. Regional producers are connected to regional consumers via a pipeline network. The USA imported most of its crude oil in 2009, but the USA produced 92% of the natural gas that it consumed that year. The imported gas was almost entirely from Canada, and shipped to the USA via pipeline. A similar regional market exists in Europe, with producers in the North Sea, Russia, and northern Africa. In some regions, such as east Asia, the potential consumption of natural gas is greater than the regional supply. In other regions, such as the Persian Gulf, regional natural gas reserves are large compared to regional consumption. One of the two major developments that affected the natural gas outlook recently is liquefied natural gas (LNG), which is converting the natural gas market into a global market. In regions with an abundance of natural gas, such as the Persian Gulf, natural gas is liquefied and loaded onto transoceanic tankers. It is then shipped to regions where local demand exceeds supply, such as east Asia, where it is vaporized and put into the
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local pipeline network. No great technological breakthrough was necessary to develop this market. Instead, sustained price differentials between supplying regions and consuming regions made LNG profitable. The second major development is a combination of two technical advances. The first is directional drilling, which allows the driller to follow the length of natural gas-bearing reservoir rock, rather than merely drilling a vertical hole through it. The second technical development is hydraulic fracturing, which fractures reservoir rock around the well bore. These two developments allow drillers to create large drainage areas in each well. Now, natural gas production from low permeability rocks, especially shale, is economically feasible. Huge volumes of shale in North America have been recently re-classified as natural gas reserves. Shale gas production, now common in North America, is expected to spread to other continents. So where does coal gasification fit in? Well, for many applications, such as ammonia and methanol production, the abundance of natural gas put gasification plans on hold. Uses of natural gas are expanding, such as an increase in natural gas for electric power production, as opposed to coal combustion, to reduce greenhouse gas production, and the potential use of natural gas as a transportation fuel to replace scarce petroleum. Of course, natural gas will not last forever, and SNG is a valuable back-up supply of methane when natural gas supplies diminish.
SNG AS A CARRIER OF COAL ENERGY Currently, most coal is burned to produce electric power. Coal is hauled over land in trains, and a large fraction of a power plant’s delivered coal cost is transportation cost. Natural gas, on the other hand, may be transported long distances at relatively low cost in pipelines. Compared to coal, natural gas combustion produces much less carbon dioxide per unit of energy. Rather than shipping coal in a train, coal could be converted to SNG in a plant near the coal mine. By-product carbon dioxide could be locally sequestered. The SNG could then be shipped via pipeline to power plants and other consumers that might otherwise burn coal. This approach will lower transportation costs and greenhouse gas emissions.
SNG VERSUS SYNTHETIC LIQUID HYDROCARBONS Traditionally, liquid petroleum products were more expensive than natural gas on an equivalent energy basis. This is why burning oil to produce heat or electricity is generally limited to regions without natural gas supplies. Lately, the spread between oil and natural gas prices has widened considerably. If current projections about crude oil supply and demand come true, that gap will widen further.
Substitute Natural Gas and Fischer-Tropsch Synthesis
As shown in the previous chapter, liquid hydrocarbons can be made from syngas via methanol and dimethyl ether. This chapter discusses a second, better-known approach which is the Fischer-Tropsch synthesis. If a decision between converting syngas to SNG or liquid fuels needs to be made, one should remember that the liquids are worth much more.
FISCHER-TROPSCH HISTORY During the 1920s, Franz Fischer and Hans Tropsch, working at the Kaiser Wilhem Institute in Berlin, developed a process for the direct conversion of syngas to hydrocarbons. This is known today as the Fischer-Tropsch (FT) process. During World War II, Germany, cut off from crude oil suppliers, produced synthetic oil using the FT process and the Bergius process, which is a direct coal hydrogenation process (see Chapter 2). Today, the Fischer Tropsch Archive5 maintains an electronic collection of early FT process documents. Sasol, using Lurgi technology, started production of FT fluids in 1955 from locally mined coal in Sasolburg, South Africa.6 A second FT plant in Secunda, South Africa, started operations in 1980. Much of the FT process technology literature is based on Sasol’s South African experience. Shell started their SMDS (Shell middle distillate synthesis) plant in Bintulu, Malaysia in 1993. This plant uses natural gas as the feedstock, and the FT fluids are hydrocracked. The Oryx GTL (gas to liquids) plant, a joint venture between Qatar Petroleum and Sasol, started operations in 2006. The Pearl GTL plant, a joint venture between Qatar Petroleum and Shell, is expected to be the world’s largest gas to liquids plant when it reaches its full 140,000 bbl/day liquids production in 2011.
FISCHER-TROPSCH CHEMISTRY Fischer-Tropsch synthesis is a polymerization reaction. An alkane chain, absorbed on a catalyst site ()), reacts with carbon monoxide and hydrogen to add a methyl unit (CH2): H-ðCH2 Þn- þCO þ 2H2 /H-ðCH2 Þn CH2- þ H2 O
R-13.4
The heat of reaction per methyl group, estimated from the gas phase heat of formation for dodecane (C12H26), is 155.5 kJ/mole. This heat of reaction is about 20% of the heat of combustion of the feed gases. Iron and cobalt catalysts are typically used, although other metals, such as ruthenium and nickel, are also active. Iron has significant water gas shift activity (R-13.3), and typical FT synthesis conditions favor the right side of the water gas shift reaction. Combining R-13.3 with R-13.4 for an iron catalyst gives: H-ðCH2 Þn- þ2CO þ H2 /H-ðCH2 Þn CH2- þ CO2
R-13.5
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The heat of reaction per methyl group, again based on dodecane, is 196.7 kJ/mole. This is somewhat higher than R-13.4 due to the exothermic contribution of the water gas shift reaction. This heat of reaction is about 24% of the heat of combustion of the feed gases. An examination of R-13.4 and R-13.5 shows that, for a cobalt catalyst, the ideal feedstock has a 1:2 carbon monoxide to hydrogen ratio, while, for an iron catalyst, the ideal feedstock has a 2:1 carbon monoxide to hydrogen ratio. With a typical syngas from a coal gasifier, a water gas shift reactor is required for a cobalt catalyst, but not required for an iron catalyst. The water gas shift reactor, could, however, shift part of the heat removal from an iron FT synthesis reactor to the water gas shift reactor, and, as we will see later, this could improve operations. The absorbed alkane chain may desorb, halting growth. The usual desorption product is a normal alkane, but alcohols and olefins also form. This means that the final product has a range of molecular weights, ranging from methane to very high molecular weight alkanes. The choice of catalyst and operating conditions determine the distribution of molecular weights. Anderson7 adapted a model developed for other types of polymerization to FT synthesis. The result is known as the Anderson-Schultz-Flory (ASF) distribution. A single adjustable parameter, a, describes the probability that an alkane absorbed on the catalyst, H-(CH2)n-), will add a methyl unit to form H-(CH2)nþ1-). Conversely, (1a) is the probability that the absorbed alkane will desorb, halting growth. The final weight fraction of each alkane compound (using the approximation that only alkanes are formed) is: Eqn. 13.1 W ¼ nan-1 ð1-aÞ2 n
Variations on the ASF distribution equations are used to improve the fit to experimental measurements of molecular weight distributions. For example, olefins appear to reabsorb and re-initiate growth, so high molecular weight alkanes are a little more abundant than predicted by Eqn. 13.1. Except for a relatively small quantity of hydrocarbon solvents, petroleum products are not sold as pure compounds. Instead, petroleum products are sold as hydrocarbon mixtures that are primarily distinguished by their boiling point ranges. There can be considerable overlap between adjacent products. For example, some of the higher molecular weight compounds in gasoline can also be some of the lower molecular compounds in diesel. Still, it is useful to approximate the FT fluid distillation product yield using molecular weight ranges. The approximate distillation yield for an FT fluid versus the chain growth probability factor, a, is shown in Figure 13.4. The catalyst, reactor design, and operating conditions are all optimized to maximize the value of the hydrocarbon products. In a typical petroleum refinery, the refinery may be operated to maximize the yield of naphtha, which is further processed to make
Substitute Natural Gas and Fischer-Tropsch Synthesis
100
Fuel gas (C1 + C2)
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Distillation Yield, wt.
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Figure13.4 Product distillation yield for a Fischer-Tropsch liquid as a function of a, the chain growth probability factor. Carbon numbers are used to estimate distillation boiling ranges. Yield versus carbon number was estimated using Eqn. 13.1.
gasoline. FT naphthas, on the other hand, tend to have very low octane numbers, while FT diesels tend to have very high cetane numbers. Lamprecht8 reports a 62 cetane number for an FT diesel, versus a 40 minimum cetane specification9 for US diesel. Both the excellent FT diesel cetane number and the poor FT naphtha octane number are for the same reason, a high normal alkane content. Therefore, FT plants tend to maximize diesel production. The least valuable fraction is the fuel gas. In a typical oil refinery, this gas is burned to provide heat for refinery operations. The next least valuable fraction is LPG. This may be recovered and sold, but the price is substantially less than for gasoline, jet fuel, or naphtha. For example, in 2009, the average US price for propane was $1.236/ gallon, versus $1.888, $1.704, and $1.833 per pre-tax gallon for gasoline, jet fuel, and diesel, respectively.10 Many refineries have dimerization processes that convert LPG to naphtha. Demand for paraffin wax is small, so the usual approach is to hydrocrack FT wax to produce lower molecular weight hydrocarbons. To maximize production of diesel, the FT reactor is designed to operate with a high chain growth probability factor, a, and then the wax is hydrocracked.
FT REACTOR DESIGN Lu11 reviewed a number of FT experimental studies; and developed approximate correlations between the chain growth probability factor, a, and reactor temperature and pressure. The most prominent trend is that a tends to decline with increasing
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temperature. This means that the operating temperature is a compromise between good reaction rates, they increase with increasing temperature; and high molecular weight distributions, which increase with decreasing temperature. A further challenge is that FT synthesis is highly exothermic, so the reactor must be designed to remove a large quantity of heat to maintain the reactor temperature within an optimum operating temperature range. Figure 13.5 shows the Sasol Synthol circulating fluid bed reactor. Fresh and recycled gas are mixed with a hot catalyst stream and fed to a vessel containing cooling bundles. The catalyst/gas mixture is then fed to a disengagement vessel where product gas is taken off the top, through cyclones to remove fine catalyst solids. The catalyst flows through a standpipe on the bottom of the disengagement vessel. The first Sasol Advanced Synthol reactor12 started production in Sasolburg in 1989. This reactor, shown Figure 13.6, is a bubbling fluidized bed reactor with heat transfer tubes in the catalyst bed. Compared to the original Synthol reactor, the Advanced Synthol reactor is more compact for a similar production capacity. The Synthol reactors operate at a fairly high temperature, 330 to 350 C. Sie and Krishna13 analyzed fluidized bed FT reactors, such as the Synthol reactors. They noted that because of the small catalyst diameter, 100 mm, there are no intraparticle diffusional
catalyst + gas product gas cyclones
cooling bundles catalyst
Fresh gas + recycle
catalyst + gas
Figure 13.5 The Sasol Synthol circulating fluid bed reactor.
Substitute Natural Gas and Fischer-Tropsch Synthesis
product gas steam
water syngas
Figure 13.6 The Sasol Advanced Synthol reactor is a bubbling fluidized bed reactor with heat transfer tubes in the catalyst bed.
restrictions. The reactor temperature and chain growth factor, a, are chosen to prevent the condensation of liquid products in the catalyst bed. The Synthol reactor, shown in Figure 13.5, was one of two original reactor types used by Sasol. The other type is the ARGE reactor, shown in Figure 13.7. This reactor is built like a shell-and-tube heat exchanger, with catalyst packed in the tubes and coolant circulating through the shell side. Both gaseous and liquid hydrocarbons are formed, so syngas is fed to the top of the reactor and liquid wax trickles down through the packed
feed gas
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water product gas
wax
Figure 13.7 The Sasol ARGE reactor. Catalyst is packed in the tubes, and coolant circulates on the shell side.
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Table 13.2 Fischer-Tropsch product distribution for iron-based catalysts.14 Component, % Low temperature (220e250 C) High temperature (330e350 C)
CH4 C2-C4 olefins C2-C4 alkanes Naphtha Distillate Oils and waxes Oxygenates
4 4 4 18 19 48 3
7 24 6 36 12 9 6
catalyst tube. There are two product ports on the bottom of the reactor, one for liquid wax and the other for gas. The ARGE reactor is operated at about 220 to 250 C, a much lower temperature than the Synthol reactor. This results in a higher chain growth factor, a, leading to a heavier product, as shown14 in Table 13.2 for iron-based catalysts. Cobalt catalysts tend to give higher a values,15 which lead to higher molecular weight products. The Shell SMDS reactor in Bintulu, Malaysia uses a packed tube configuration similar to the ARGE reactor. More recently, slurry bubble reactors have been used for low temperature FT synthesis. The reaction media is a three phase churning mixture of solid catalyst, liquid wax product, and feed and product gas. Heat exchange tubes in the mixture remove the heat of reaction. Catalyst particle diameters14 are between 20 and 200 mm. The lower diameter limit is set by the difficulty of removing smaller solids from the product wax, and the upper diameter limit is set by the need to keep the catalyst particle suspended in the reaction media. There have been reports of operational difficulties with slurry bubble reactors. Figure 13.8 shows a microchannel reactor.16 This reactor is created by etching parallel grooves in flat metal plates. The plates are stacked to form flow channels bounded by
Figure 13.8 A pilot scale microchannel reactor for Fischer-Tropsch synthesis built by Velocys, Inc.
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adjacent plates. Alternating layers of flow channels are filled with catalyst. The channels between the catalyst channels contain a heat transfer fluid, flowing at right angles with respect to the reactant flow. Conceptually, this reactor is much like the packed tube ARGE reactor. The microchannel reactor, however, has a much higher heat transfer area/catalyst volume ratio and a much smaller flow channel cross section. Both channels allow for a much better control of the reaction temperature range.
REFINING FISCHER-TROPSCH FLUIDS FT fluid can be thought of as a synthetic crude oil. As such, it is an unusual crude oil that requires special refining considerations. A typical crude oil contains normal alkanes, branched and cyclic alkanes, aromatics, sulfur compounds, and nitrogen compounds. Unless the crude oil is produced using thermal techniques, crude oil contains essentially no olefins. Oxygen compounds are present in crude oil, but usually at low levels. FT fluid, on the other hand, consists primarily of normal alkanes, with lesser but significant levels of olefins and alcohols. FT fluid contains very few branched alkanes, and essentially no cyclic compounds, sulfur or nitrogen compounds. We17,18 modeled FT refineries using a pure component approach and the Aspen Plus process simulation software. Three scenarios were considered in this study; a refining scheme that maximized diesel production, another, very similar scenario that maximized jet fuel production, and a third refining scheme to produce lubricating base oils.
Diesel Production FT diesel fuel, due to its high alkane content, has very high cetane numbers. Blends of FT and conventional diesel are sold as premium products in Europe. Normal alkanes, however, have high melting points compared to branched or cyclic hydrocarbons. Normal hexadecane, for example, has a melting point of þ18 C, versus 23 C for its branched isomer, 3-methyl-pentadecane. An FT diesel produced simply by distilling an FT fluid would have poor low temperature properties. This may be one of the reasons why FT diesel is sold in Europe as a blend with conventional diesel, rather than as a pure product. Most FT diesel, however, is not simply an FT fluid distillation product, but a blend of this distillation product and products of hydrocracked FT wax. Sie19 studied the hydrocracking of normal alkanes, and presented evidence for a cyclopropane intermediate mechanism. As a result of this mechanism, propane is the smallest fragment produced by a hydrocracking reaction. This reaction converts a normal paraffin molecule into two smaller molecules. One of these is also a normal alkane, while the other is a branched alkane. This shows a close relationship between hydrocracking and isomerization. As an approximation, any of the carbonecarbon bonds within the molecule,
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except for the two closest to the ends, have an equal probability of fracturing. Sie’s data actually show that the bonds towards the center of the molecule are slightly more likely to fracture than bonds near the end of the molecule. Leckel and co-workers20-23 published several papers describing Sasol FT wax hydrocracking pilot plant experiments. The FT wax is hydrocracked and then distilled. The distillation bottoms, which are unconverted wax, are recycled to extinction. Hydrocracking severity (conversion) can be controlled using a number of process parameters, including catalyst acidity, reactor temperature, reactor pressure, and reactor space velocity. Leckel showed that low temperature properties of diesel improve with increased hydrocracker severity, which is expected since hydrocracking creates branched alkanes. Increased hydrocracking also resulted in lower diesel yields, since a larger fraction of the hydrocracked product was naphtha or lighter products. Thus, there is a close relationship between diesel yield and low temperature properties. In our diesel model, we distilled an FT fluid produced by a cobalt catalyst with a high chain growth factor, a. Both the wax and the diesel fractions were hydrocracked. The diesel fraction did not need to be hydrocracked to meet the boiling point specifications. Rather, this was done to increase the branched alkane/normal alkane ratio in the final diesel product. Sie’s rules were used to model the hydrocracker; and the extent of hydrocracking was arbitrarily adjusted so that our simulation results fit Leckel’s experimental results. Although we were able to estimate the diesel composition, we did not find a reliable means of predicting low temperature properties based on that composition. In the USA, diesel cloud points are specified9 according to the time of year and the local climate. The cloud point is the temperature at which wax crystals become visible as the sample is cooled. Below the cloud point, these wax crystals plug fuel filters. Krishna et al.24 offer an approximate method of predicting pour points based on composition. Pour points are typically a few degrees lower than cloud points. Leckel’s results show that FT diesel can meet low temperature specifications, but we are not currently able to reliably predict diesel yields without experimental data.
Naphtha and Gasoline Naphtha processing proved to be challenging, due to its low octane numbers for FT naphtha. Brent et al.25 investigated FT naphtha upgrading techniques. They approached two naphtha reforming process providers, and were told by both that FT naphtha was not a good reforming feedstock. Naphtha reforming increases octane numbers by increasing aromatic content, primarily through dehydrogenation of cycloalkanes. Instead, they chose to sell FT naphtha as an olefin production feedstock. FT naphtha, compared to other naphthas, is an excellent olefin production feedstock because of its high alkane content. There are two problems with this approach. The first is that this would put FT
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naphtha in direct competition with natural gas liquids, which sell at substantially lower prices than gasoline. The second is that olefin plants designed to use naphtha feed are primarily located in countries that do not have significant natural gas production. In our model, we sought to convert FT naphtha to gasoline. To do this, we used two processes. The first is the TIP (total isomerization process) process,26 in which the octane number is increased by converting normal alkanes into branched alkanes. The isomerization reactor effluent is fed to a mole sieve bed, where normal alkanes are selectively adsorbed and recycled to the isomerization reactor. We used Sie’s hydrocracking rules to model the isomerization reactor, although catalysts with better isomerization/hydrocracking selectivity are available. The second process is an alkylation process that converts C4 compounds to isooctane. The problem with this refining approach was that octane improvement was accompanied by molecular weight reduction due to hydrocracking in the isomerization reactor, and we could not simultaneously meet both octane specifications and boiling point specifications. A more selective isomerization catalyst would improve the performance of this refining approach. Just as steam reforming of natural gas is a common method of producing syngas, LPG and naphtha can also be steam reformed to produce syngas.27 If a suitable outlet for FT naphtha or LPG is not found, then these streams can be recycled, thereby increasing the fractional yield of diesel or jet fuel. Our jet fuel model was very much like our diesel model, with the boiling point range adjusted to meet jet fuel standards. When jet fuel production is maximized, the diesel yield is insignificant.
Lubricating Base Oils We also examined a refinery model to maximize the production of lubricating base oils. This is the largest volume non-fuel petroleum market. A typical lubricating oil contains about 90% base oil and about 10% of an additive package. Branched alkanes, primarily in the C20 to C40 range, offer the best combination of volatility, viscosity, and low temperature performance for base oils. Traditionally, base oils with high branched alkane content are made using a paraffinic crude oil. A vacuum distillate cut is processed to increase the alkane content, either by solvent extraction of aromatics, or by converting aromatics to alkanes in a catalytic hydrogenation reactor. The oil is then dewaxed (high melting point normal alkanes are removed) by mixing the oil with a blend of methyl ethyl ketone and toluene, chilling the oil to crystallize wax, and then filtering the wax/ oil/solvent slurry. The solvent dewaxing process is the most expensive step in the traditional base oil production process. The traditional base oil process produces a complex and variable mixture of hydrocarbons. With increasingly stringent lubricating oil standards, producers found it difficult to meet standards with these traditional base oils. This led to the production of
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synthetic oils, such as polyalpha olefins, in which the hydrocarbon molecular structure can be more closely controlled to optimize oil properties. The cost of synthetic base oils is much higher than that of traditional base oils. In the Chevron Isodewaxing process28 normal alkanes are isomerized to branched alkanes. The pore size of the catalyst allows methyl branches to form, but the pores are too small to readily allow the formation of the cyclopropane intermediate compound that leads to hydrocracking. As a result, this catalyst has excellent isomerization/ hydrocracking selectivity. The Isodewaxing process was initially marketed as an inexpensive alternative to solvent dewaxing, but the isomerized waxes have physical properties that are superior to traditional base oils. When the feedstock is a paraffin wax, the base oil product is nearly as good as polyalpha olefins. This process offers a low cost route to high quality lubricants. Partly in response to this process, the American Petroleum Institute developed a classification system29 for base oils. Traditional base oils are either Class I or Class II, depending on their sulfur content. The higher quality base oils produced by wax isomerization are Class III. Poly alpha olefins are Class IV, and other types of synthetic base oils are Class V. In our lubricating base oil refinery model, the FT wax is isomerized using the Isodewaxing process. Detailed technical information on this process in the open literature is scarce, so we modeled the isomerization reactor based on pure component yield data from Claude et al.30-31 for a different catalyst, Pt/H-ZSM-22. This catalyst has a pore geometry that is similar to the Chevron catalyst. The yield data were for normal alkane feeds between C10 and C24. Yield data were extrapolated to higher molecular weights for this simulation. Three different viscosity grade oils were made. The base oil refinery model is attractive, in part, because base oils are worth considerably more than fuels. In March, 2010, US Gulf Coast Group III base oils sold for about $3.50/gallon,32 about twice the price of diesel.
Outlook for FT Fluid Refining In our refinery models, we sought, to use conventional petroleum refining processes. The performance of the fuel refinery cases could be improved with the application of an isomerization catalyst with high isomerization/hydrocracking selectivity, such as the Chevron Isodewaxing catalyst used for base oil production. A combination of this catalyst and hydrocracking would give the refiner independent controls over molecular weight and low temperature properties, and this could greatly enhance diesel or jet fuel yields. It is not clear whether gasoline octane specifications can be met, but a higher octane is achievable if better isomerization selectivity is attained in the naphtha isomerization reactor. As a stand-alone process for liquid hydrocarbon production from syngas, FischerTropsch technology competes with methanol-based processes, such as the ExxonMobil
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MTG process discussed in the previous chapter. The MTG process is attractive because it produces specification grade gasoline in high yield with minimal processing. Fischer-Tropsch technology may be better suited as a companion technology for heavy oil refining. The trend in petroleum refining is away from light crude oils and towards heavy crudes. This produces much more aromatic products. High aromatic content in gasoline boosts octane numbers, but aromatic contents are limited due to toxicity concerns about aromatics, especially benzene. Aromatics tend to lower diesel cetane numbers. FT products are nearly free of aromatics, when blended with products from heavy crudes, the blends more closely resemble the mixture of alkanes and aromatics that are normally associated with refined products made from light crudes.
FT ECONOMICS Van Bibber et al.33 prepared a process and economic study of a Fischer-Tropsch plant in Healy, Alaska, based on Usebelli coal, a locally mined sub-bituminous coal. Their plant would produce 14,640 bbl/day of FT fluids, of which 58 vol.% would be diesel and 42 vol.% would be naphtha. Diesel prices were assumed to be the same as conventional diesel prices, rather than charge a premium for the higher cetane FT diesel. They also assumed that they could sell the FT naphtha for a $0.10 discount relative to gasoline, rather than dealing with naphtha upgrading issues. We34 modified this report for a Wyoming Powder River Basin (PRB) location. This is appropriate because PRB coal is very similar to Usebelli coal. To account for the higher elevation of the PRB site, another compressor was added to the air separation unit. Capital costs were updated to 2008. Locally, there is a demand for carbon dioxide for enhanced oil recovery, so we assumed that we could sell carbon dioxide for $1.00 per million standard cubic feet. This assumption had little impact on plant economics. With carbon capture and sequestration, the plant would cost about $1.9 billion. The FT fluid cost would have to be $89/bbl for a 10% internal rate of return. This is close to the pre-tax price of gasoline and diesel when this was written. Figure 13.9 is a tornado
Figure 13.9 Effect of a 25% change in assumed economic parameters on the percent internal rate of return for a Fischer-Tropsch coal-to-liquids plant with carbon capture and sequestration in the Powder River Basin of Wyoming. For the base case, $89/bbl is the assumed product price.
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plot showing how a 25% change in assumed economic parameters affects the internal rate of return. Not surprisingly, the plant economics are sensitive to product price. The process economics are dominated by capital construction costs, so changes in capital costs and plant availability have a strong impact on profitability. Coal costs, on the other hand, have a minor effect. The average mine-mouth price of Wyoming coal in 2008 was $11.39/ton, so a 25% change in this price does not have a large impact on operating costs. The price of carbon dioxide has little impact on the overall economics.
REFERENCES 1. Nahas NC. Exxon catalytic coal gasification process, fundamentals to flowsheets. Fuel. 1983;62: 239-241. 2. Great Point Energy. Our technology.
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23. Leckel D. Low-pressure hydrocracking of coal-derived Fischer-Tropsch waxes to diesel. Energy & Fuels. 2007;21:1425-1431. 24. Krishna R, Joshi GC, Purohit RC, Agarwal KM, Verma PS, Bhattacharjee S. Correlation of pour point of gas oil and vacuum gas oil fractions with compositional parameters. Energy & Fuels. 1989;3: 15-20. 25. Brent FD, Shah L, Berry E, Schrader CH, Anderson J, Erwin J, et al. Early entrance coproduction plant, task phase II, task 2.5: Fischer-Tropsch product upgrading. Texaco Energy Systems, U.S. Dept. of Energy DE-FC26e99FT40658; 2004. 26. Symoniak MF. Upgrade naphtha to fuels and feedstocks. Hydrocarbon Proc. 1980;110-114. 27. Riddler DE, Twigg MV. Steam reforming. In: Twigg MV, ed. Catalyst Handbook. 2nd ed. Wolfe Publishing; 1989;225-282. 28. Krishna KR, Rainis A, Marcantonio PJ, Mayer JF, Biscardi JA, Zones SI. Next generation Isodewaxing and hydrofinishing technology for production of high quality base oils. Presented at the NPRA Lubricants and Waxes Meeting, Houston; Nov. 14-15, 2002. 29. American Petroleum Institute. Appendix E: API base oil interchangeability guidelines for passenger car motor oils and diesel engine oils.
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