Gasification process technology

Gasification process technology

5 Gasification process technology C. HIGMAN, Higman Consulting GmbH, Germany Abstract: This chapter discusses gasification as a means of conversion o...

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5 Gasification process technology C. HIGMAN, Higman Consulting GmbH, Germany

Abstract: This chapter discusses gasification as a means of conversion of hydrocarbon residuals. After a discussion of typical gasification feedstocks in the hydrocarbons industry, there follows a section on the basic principles of gasification. This includes a review of the characteristics that define a particular gasifier. The section on commercial gasifiers looks at those specifically that have been used in hydrocarbon (as opposed to coal) service. A section on gas treatment precedes a description of a complete system for producing hydrogen and power. Finally, some future developments in gasification and related technologies are discussed. Key words: gasification, synthesis gas, petroleum coke, petroleum residues.

5.1

Introduction

Gasification is a technology that converts carbonaceous fuels, including typically both solid and liquid refinery residues, to synthesis gas (syngas) (H2 + CO) by partial oxidation. In the refinery environment this syngas is used to manufacture hydrogen, produce chemicals such as methanol and ammonia, and to generate power. Other applications include processing of residues from oil sands recovery. The same technology is also used to generate syngas from natural gas or refinery gases, mainly when lower H2/CO ratios are required than those that can easily be obtained from a steam-methane reformer. The techniques for gasifying solid hydrocarbon residues are very similar to those developed for coal gasification. This chapter will, however, concentrate on petroleum coke applications. For those interested in coal applications, reference is made to the literature (for example, Higman and van der Burgt, 2003).

5.2

Gasification in the refinery environment

5.2.1 Feedstocks in the refinery Before entering into a detailed description of partial oxidation and gas treatment technology, it may be useful to review how the partial oxidation process fits into the overall refinery picture. What is its location in the overall refinery flowsheet? What products can it produce and how do they relate to the market? What feedstocks can it handle and what are the environmental demands placed on it? Figure 5.1 shows an atmospheric distillation unit (ADU), vacuum distillation unit (VDU), a solvent deasphalter (SDA) or alternatively a coker and a gasification unit processing the asphalt or coke. 155 © Woodhead Publishing Limited, 2011

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5.1 Gasification in the refinery environment. ADU: Atmospheric distillation unit, LN-HDT: light naphtha hydrotreater, HN-HDT: heavy naphtha hydrotreater, VDU: vacuum distillation unit, HCU: hydrocracking unit, SDA: solvent deasphalter, HDS: hydrodesulfurization, FCC: fluid catalytic cracker, MD-HDS: middle distillate hydro-desulfurization.

This refinery configuration is aimed in the first instance at maximizing white products and eliminating any heavy fuel oil export. On the basis of the planned crude slate the total vacuum residue will be about 16% of crude intake. A small proportion of this will exported as street asphalt. The rest is fed to a SDA unit, which is capable of recovering up to 50% of the vacuum residue as deasphalted oil (DAO). After desulfurization this is suitable as fluid catalytic cracker (FCC) feed. The resulting asphalt, about 6.5% of intake, is then gasified. This type of processing produces an asphalt that is high in pollutants such as sulfur and heavy metals. It also has a high viscosity, which is where gasification comes in. By gasifying the residue instead of putting it into the fuel oil pool, it is possible to extract both major pollutants separately and in an economically usable form. The products of gasification and subsequent gas treatment are hydrogen, methanol and a clean, high-pressure fuel gas for power generation in gas turbines. Alternative refinery configurations may see a visbreaker or a coker in place of the SDA. While this may alter the anticipated yields of gasoil, the layout of the refinery is, in principle, the same. The design of the solids gasifier processing coke, however, is very different from (and more expensive than) unit processing liquid residues from a visbreaker or SDA. Table 5.1 shows some typical properties of these feedstocks.

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Table 5.1 Typical refinery residue specifications Feedstock type (elementary analysis) Carbon (wt%) Hydrogen (wt%) Sulfur (wt%) Nitrogen (wt%) Oxygen (wt%) Ash (wt%) Total (wt%) C/H (kg/kg) Vanadium (mg/kg) Nickel (mg/kg) Iron (mg/kg) Sodium (mg/kg) Calcium (mg/kg) Kinematic viscosity (100 °C) (cSt) Density (15 °C) (kg/m3) LHV (MJ/kg)

Visbreaker residue

Butane asphalt

Petroleum coke (moisture free)

85.27 10.08 4.00 0.30 0.20 0.15 100.00

84.37 9.67 5.01 0.52 0.35 0.08 100.00

89.23 3.59 5.23 1.35 0.10 0.50 100.00

8.3 270–700 120 20 30 20 10 000 1 100 39.04

8.7 300 75 20 30 20 60 000 1 070 38.24

31.93 1500 340

30.26

Sources: Posthuma et al., 1997; Mahagaokar and Hauser, 1994

As can be seen, the liquids can have very high viscosities, which presents various challenges, not least the logistics of transporting the material from the SDA to the gasifier. All of these feedstocks have high sulfur contents. The ash is vanadium- and nickel-rich.

5.2.2 Products for the refinery As can be seen from Fig. 5.1, there are many parts of the refinery that require hydrogen, and there is a trend for requiring ever-increasing quantities, primarily for desulfurization of transport fuels. It is therefore not surprising that much of the syngas from residuals gasification is converted to hydrogen. Prominent examples include the 240 t/d gasification-based hydrogen plant in Shell’s Pernis, The Netherlands, refinery or the 150 t/d hydrogen plant at the Convent Louisiana refinery (US DOE NETL, 2010b). A number of refineries in Germany produce methanol, including Total’s Leuna refinery (2090 t/d) and BP’s Gelsenkirchen refinery (850 t/d). The latter also produces ammonia (800 t/d) and up to 50 000 Nm3/h of hydrogen (Laege and Pontow, 2002). The Coffeyville Resources, KS, 1000 t/d ammonia plant gasifies petroleum coke received over the fence from the Coffeyville Refinery (Barkley, 2006). Examples of power plants based on gasification of refinery residues include four in Italy, the largest of which is adjacent to the Saras refinery and generates

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about 550 MWe from the refinery residue. An example in the United States is the 160 MWe power plant in the Valero refinery in Delaware City, which operates on petroleum coke. Other plants owned by power utilities gasify petroleum coke bought on the free market either alone or in admixture with coal. These include Tampa Electric’s 250 MWe Polk Power Station, FL, and the 250 MWe Wabash River Power Station, IN. Plants for the partial oxidation of a gaseous feed tend to be small plants for production of carbon monoxide or oxo-synthesis gas. Gas-to-liquids facilities such as the Shell 140 000 bbl/d Pearl project in Qatar currently in start-up (2010), however, can be extremely large. A number of plants including some of those mentioned above produce more than one product. This situation, whereby the syngas is fed to different final uses, is generally known in the industry as polygeneration.

5.2.3 Gasification in oil sands development A further application of gasification to hydrocarbons can be found in the oil sands industry. Most of the current oil sands operations in Canada produce large quantities of coke, which is landfilled in the remote regions where the oil sands are located. At the same time a significant amount of natural gas is used to produce hydrogen for bitumen upgrading. There is certainly the potential to use the coke as raw material for hydrogen manufacture. A number of such projects have been discussed, though none have yet been realized. On the other hand, in 2007 one of the largest liquids gasification facilities, located at Long Lake, Alberta, was started up. It processes an asphalt residue obtained from oil sands recovery to generate hydrogen for the upgrading of bitumen, and steam for its extraction, using the steam-assisted gravity drainage (SAGD) method (Rettger et al., 2006).

5.3

Basic principles

Gasification can be described as the ‘conversion of any carbonaceous feedstock into a gaseous product with a useful chemical heating value’. Early processes, such as those used for the production of town gas from coal in the nineteenth century, emphasized devolatilization and pyrolysis reactions creating a gas with significant hydrocarbon content for lighting purposes. With modern processes the emphasis has shifted to partial oxidation and water gas processes generating syngas, in which carbon monoxide and hydrogen are the main components.

5.3.1 Chemistry and thermodynamics The principal reactions that take place during the gasification of pure carbon are those involving carbon, carbon monoxide, carbon dioxide, hydrogen, water (or steam) and methane. The most important are partial oxidation:

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−111 MJ/kmol

[Reaction 5.1]

+131 MJ/kmol

[Reaction 5.2]

and the water gas reaction: ← CO + H C + H2O → 2

In gasification processes the reactions with free oxygen are all essentially complete. The carbon conversion is in general also largely complete. Most processes operate at sufficiently high temperatures that equilibrium is reached and the final gas composition is determined by the CO shift reaction: ← CO + H CO + H2O → 2 2

−41 MJ/kmol

[Reaction 5.3]

and the steam methane reforming reaction: ← CO + 3 H CH4 + H2O → 2

+206 MJ/kmol

[Reaction 5.4]

Gasification temperatures are generally so high that, thermodynamically as well as in practice, no hydrocarbons other than methane can be present in any appreciable quantity. Depending on the reactor arrangement it is possible, however, that some pyrolysis products survive and are contained in the synthesis gas. The gas composition changes with the pressure and temperature of the gasifier. As pressure rises, methane and CO2 content in the syngas increase. With increasing gasification temperature, the methane contents drop and the H2/CO ratio moves towards increasing CO. Typical commercial gasification processes today operate in the range 25–80 bar, depending on application. At these pressures, temperatures of above 1250 °C are required to produce a syngas with a low methane content. Note that while an increased methane content is beneficial for power generation in terms of heat supply to a combustion turbine, large quantities of methane generated by processes optimized for synthetic natural gas (SNG) production, for example, would be counter-productive in a hydrogen manufacture or chemical production scenario.

5.3.2 Reactor type In the practical realization of gasification processes a broad range of reactor types has been and continues to be used. Petroleum coke gasifiers are, in general, units that have originally been designed for coal. Most considerations are similar except for the nature of the ash, which is discussed later. The most important differentiating characteristics of different processes are discussed in the following sections. Reactor bed type For most purposes, reactor types for solids gasification can be grouped into one of the three categories shown in Table 5.2: moving bed gasifiers, fluid bed gasifiers and entrained flow gasifiers. The gasifiers in each of these three categories share certain characteristics that differentiate them from gasifiers in other categories.

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Table 5.2 Characteristics of different categories of gasification process.

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Category Ash conditions

Moving bed Dry ash

Moving bed Slagging

Fluid bed Dry ash

Fluid bed Agglomerating

Entrained flow Slagging

Typical processes

Lurgi

BGL

Winkler, High temperature winkler (HTW), circulating fluid bed (CFB), KBR Transport Gasifier

Kellogg Rust Westinghouse, U-Gas

Shell, GEE, E-Gas, Siemens, Koppers-Totzek

Feed characteristics Size (mm) Acceptability of fines

6–50 Limited

6–50 Better than dry ash Yes High

6–12 Good

6–12 Better

< 200 µm Unlimited

Possibly Low

Yes Any

Yes Any (dry feed) High (slurry feed)

Low (450–650)

Moderate (900–1050)

High (1250–1600)

Low Low Hydrocarbons in gas

Moderate Moderate Lower carbon conversion

Moderate (900–1050) Moderate Moderate Lower carbon conversion

Acceptability of caking coal Yes (with stirrer) Preferred coal rank Any Operating characteristics Outlet gas temperature (°C) Low (450–650) Oxidant demand Steam demand Other characteristics

Low High Hydrocarbons in gas

High Low Pure gas, high carbon conversion

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Moving bed gasifiers Moving bed gasifiers (also called fixed bed gasifiers) are characterized by a bed, in which the solid feed (coal or coke) moves slowly downward under gravity as it is gasified by a blast or oxidant, which generally, but not universally, moves in counter-current to the feed. In such a counter-current arrangement, the hot syngas from the gasification zone is used to preheat and pyrolyze the downward-flowing solids. This arrangement reduces the oxygen consumption but, when coal is being processed, pyrolysis products as well as moisture are present in the product syngas. The outlet temperature of the syngas is generally low, even if higher temperatures are reached in the heart of the bed. Moving bed processes operate on lump material. An excessive amount of fines can block the passage of the up-flowing syngas. Fluid bed gasifiers Fluid bed gasifiers offer good mixing between feed and oxidant, which promotes both heat and mass transfer. This ensures an even distribution of material in the bed and hence a certain amount of only partially reacted fuel is inevitably removed with the ash. This places a limitation on the carbon conversion of fluid bed processes, typically to a maximum of about 95%. The operation of fluid bed gasifiers is generally restricted to temperatures below the softening point of the ash, since agglomeration of soft ash particles will disturb the fluidization of the bed. Sizing of the particles in the feed is critical: material that is too fine will tend to become entrained in the syngas and leave the bed overhead. This is usually partially captured in a cyclone and returned to the bed. The lower temperature operation of fluid bed processes means that they are generally better placed to handle reactive feedstocks such as low-rank coals and biomass rather than coke. Nonetheless there are examples of fluid bed gasification of petroleum coke, such as in the ExxonMobil Flexicoker process.

Entrained flow gasifiers Entrained flow gasifiers operate with feed and oxidant in co-current flow and are the most commonly used form of gasifier for petroleum coke feeds. The residence time in these processes is short (a few seconds). The feed is ground to 200 µm or less to promote mass transfer and allow transport in the gas. Given the short residence time, high temperatures are required to ensure a good conversion. Therefore, all entrained flow gasifiers operate in the slagging range (i.e. above the melting temperature of the ash). The high-temperature operation creates a high oxygen demand for this type of process. The ash is produced in the form of an inert slag or frit, which can be used as a construction material. This is achieved with the penalty of additional effort in feed preparation as well as a high oxygen consumption.

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The majority of successful solids gasification processes that have been developed since 1950 are entrained flow, slagging gasifiers operating at pressures of 20–80 bar and at temperatures of at least 1250 °C. Entrained-flow gasifiers have become the preferred gasifier for hard coals and petroleum coke. Feed preparation There are two principal feed systems for feeding a solid fuel into a pressurized gasifier. Some licensors such as GE Energy (GEE), East China University of Science and Technology (ECUST) and ConocoPhillips (E-Gas) use a coal or coke–water slurry whereas Shell, Siemens and Mitsubishi use dry-feed systems. In a dry-feed system the coke or coal is ground and dried together with a fluxing agent in a roller mill with a hot gas drying circuit, similar to that used in conventional pulverized coal power plants. The pulverized solids are fed through a lock hopper system into the pressurized feed vessel. The coke is then transported to the burners from the feed vessel by pneumatic conveying in the dense phase. The carrier gas is typically pure nitrogen from the air separation unit (ASU), but for some chemical applications where nitrogen is undesirable, CO2 can be used. Generally, a dry feed system contributes to a higher gasifier efficiency. However, the amount of carrier gas required for the pneumatic transport of the solids into the gasifier increases with pressure. The economic limit for dry feed systems is generally considered to be about 40 bar. For wet feed systems the slurry is made in a rod mill into which pre-crushed (∼50 mm) coke, fluxing agent and water are fed. The solids are ground in a wet milling process to a size of about 100 µm. Typically, the slurry is pumped to the reactor pressure by a membrane piston pump, which allows gasifier operation at up to 80 bar. This can be an advantage for some chemical applications. The water content of the slurry is usually in the range 35–40%. The need to evaporate the water from the slurry in the gasifier reduces the efficiency of slurry feed systems. Operating temperature Operating temperature, i.e. whether to go for a slagging or non-slagging operation, is another fundamental choice. While a decision may be connected with the bed type – all entrained flow processes operate in the slagging zone – the moving bed offers a choice between, for example, the Lurgi dry bottom gasifier and the BGL slagging gasifier. In all cases it is necessary to ensure that the temperature is either high enough that there is an adequate safety margin above the ash fluid temperature and its critical viscosity temperature (TCV, i.e. the temperature at which the slag flows easily, typically when the viscosity is about 25 Pa s) or that the operating temperature is sufficiently lower than the ash softening temperature that ash particle agglomeration does not interfere with the operation of the bed (whether

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fluid or moving). In between these two temperatures is effectively a ‘no-go zone’ in which sticky ash will create operating problems in any system. Oxidant The choice of oxygen or air is another issue that gives rise to discussion. Historically the high-temperature entrained flow processes use oxygen, and there are good reasons for this. Partly this has been dictated by the fact that in the period 1935–1985 most gasifiers were built for chemical applications where the presence of large quantities of nitrogen in the syngas was detrimental to the downstream synthesis process. There are, however, other technical reasons why oxygen is preferred over air in this configuration, even when the presence of nitrogen is not a fundamental problem. As can be seen from Fig. 5.2, while the cold gas efficiency does not vary very much over the range 85–99% oxygen, it falls off ever more rapidly the closer it approaches the level of 21% oxygen found in the atmosphere. Essentially, this represents the penalty of having to heat up the nitrogen to the gasification temperature, which was chosen as 1500 °C in the example. The oxygen demand is increased and the syngas quantities to be cooled and treated are approximately doubled. These disadvantages are more than enough to offset the capital and operating cost of an air separation plant. Reactor containment Containment of the reactor is another choice that requires consideration. Some means of protecting the pressure shell from the reaction temperature is required.

5.2 Cold gas efficiency as a function of air enrichment in gasifier oxidant.

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The alternatives are:

• • •

a refractory lining a water-cooled membrane wall between the reaction space and the pressure shell a water jacket integral with the pressure shell.

Although most processes have settled for one system or another, Siemens offers a choice between all three according to application. Refractory lining is the cheapest first-cost solution, but for slagging gasifiers it requires regular maintenance. The hot face may require replacement every two years depending on gasifier operating temperatures and flux quality, possibly with some minor maintenance in between. A full hot face change out can take up to four weeks, so that plants with a high-availability requirement tend to install a spare gasifier, which negates the cost advantage. The membrane wall solution relies on a layer of solidified slag between the watercooled wall and the hot gas space to protect the former. The liquid slag then flows down the wall of solid slag, as shown in Fig. 5.3. This solution requires considerably more in terms of investment than refractory lining, but once installed it is relatively maintenance-free. Membrane walls have given proven, successful, trouble-free operation of ten or more years, and they have a predicted life of over twenty years. Primary syngas cooling The next choice that we will consider is that of primary syngas cooling. The purpose of primary cooling in a slagging gasifier is to bridge the no-go zone

5.3 Gasifier containment systems.

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between free flowing liquid slag and dry solid ash. The method of cooling must be chosen so that no sticky ash adheres to heat exchanger or other surfaces while it is in the intermediate temperature range. Here one can select between water quench, gas quench and radiant cooling, together with or without a convective cooler. GEE, for instance, offers a choice between a full water quench and (for solids gasification) a radiant cooler. Shell offers a gas quench. Staged gasification is a feature which can increase efficiency. It is used by E-Gas and Mitsubishi. Part of the feed is injected to a first stage (Fig. 5.4), where all of the oxygen is consumed. This stage operates hot, under slagging conditions, and the slag is drawn off at the bottom. The remainder of the feed is added at the second stage where it reacts with the hot syngas from the first stage. In the second stage the additional feed is dried and devolatilized and part of the fixed carbon is gasified. The gas leaving the second stage is cooled to below the ash fusion temperature by the endothermic reactions. The ash is therefore dry at this point and the gas can be cooled in a convective heat exchanger. The remaining un-gasified char and the ash are recovered from the gas and recycled to the first stage. This ensures a high overall carbon conversion despite the low exit temperature from the reactor. It also has the effect that the ash from the second stage feed is also slagged and discharged from the first stage. Primary gas cleaning Primary gas cleaning, i.e. removal of particulates as well as chlorides and bulk ammonia from the gas, is usually considered to be part of the complete gasification process and as such is supplied by all of the gasification technology vendors. Particulate removal may be performed wet, in a scrubber (e.g. GEE and Siemens), or dry with a candle filter (E-Gas and Shell) for solids gasification. The wet systems produce ‘black water’ from the scrubber bottoms, which must be

5.4 Two-stage gasification.

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cooled and cleaned. The bulk of the chlorides and ammonia is removed in the scrubber, though the process condensate from the low-temperature gas cooling will still contain significant quantities of ammonia. While the water handling in the dry systems is much simpler, the candle filter is itself a source of maintenance cost. A variety of materials is used for the candles. Shell uses ceramic candles; E-Gas uses sintered metal. In both cases a subsequent water wash is required to remove the ammonium chloride. The wash water is, however, largely free of particulate matter. Specifics of petroleum coke gasification While most entrained flow coal gasification processes can be operated on a petroleum coke feedstock, there are a number of issues that need to be considered. The most important aspect is the low ash content and the nature of the ash, a large part of which is composed of vanadium and nickel. Flux material, which can be selected to ensure a relatively low ash melting point, is therefore added to the feedstock, typically about 5% of the latter. The slag produced from the flux helps to maintain the insulation layer of membrane wall gasifiers and dilutes the vanadium and nickel content in the slag (Mahagaokar and Hauser, 1994). A second issue is the very low oxygen content, which results in a lower reactivity. Furthermore, the content of volatile matter and hydrogen are also low. This can result in the need to operate at a slightly higher temperature or accept a lower per pass carbon conversion. For dry feed gasifiers, moderate amounts of steam are added to the coke–oxygen mixture (Mahagaokar and Hauser, 1994). The third characteristic of petroleum coke is its high sulfur content (typically 4–6%). This is less a matter for the gasifier than for subsequent gas treatment and sulfur recovery. Nonetheless it can be of importance for material selection in some gasifiers.

5.3.3 Commercial processes There is an extremely wide variety of gasifier systems that have been implemented, particularly in the field of small-scale biomass and waste gasification. The number of processes operating or under construction at sizes over 100 MWth for hydrocarbon feeds is, however, limited. It should also be noted that most entrained flow coal gasifiers are, in principle, suitable for operation with petroleum coke, but only those with actual commercial experience with petroleum coke or liquid refinery residues are discussed here. Those not discussed for this reason include the Siemens Fuel Gasification (SFG) process and ECUST’s opposed multiple burner (OMB) process. In the refinery environment, operation with air as an oxidant is unlikely to be attractive, since the dilution of syngas with approximately 50% nitrogen is not compatible with hydrogen production, which is likely to be at

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least part of a refinery-based gasifier’s product slate. Thus air-blown processes such as those of Mitsubishi or KBR are also not described here. A complete listing of gasification plants in operation and planned as of November 2010 has been compiled by the US Department of Energy National Energy Technology Laboratory (DOE NETL, 2010b). A total of 412 gasifiers are in operation, with a further 17 under construction and 76 in planning. The output of the gasifiers operating and under construction is over 80 GWth. Of this about 18 GWth is based on petroleum residues and 1 MWth on petroleum coke. This latter figure does not include those plants originally designed for coal and now operating largely on petroleum coke such as Tampa and Wabash River. GE Energy (formerly Texaco) The GEE solid-feed gasifier is a down-flow, entrained flow, slagging design. It uses a coal/coke–water slurry feed, which makes for a simpler, cheaper design than a dry feed operation, but at the cost of a slight efficiency penalty. The reactor containment uses a refractory lining. Depending on application, cooling may be by water quench (Fig. 5.5), or radiant cooling. The first Texaco coal gasifier was demonstrated in Oberhausen, Germany, in 1978. In 1984 commercial units for Eastman in Kingsport, TN, and for Ube in Japan were taken into service. The Eastman plant uses coal feed and

5.5 GE quench gasifier (courtesy of GE Energy).

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the Ube plant, which has since been expanded, operates on petroleum coke. In the same year the Cool Water 100 MWe demonstration Integrated Gasification Combined Cycle (IGCC) was started up. In 1996 a 250 MWe power plant for Tampa Electric went into operation in Polk County, FL. A 630 MWe power plant without carbon capture and storage (CCS) is currently under construction at Edwardsport, IN. GEE technology has been selected for the Hydrogen Energy 390 MWe petroleum coke gasification project at Bakersfield, CA (Hydrogen Energy, 2010). Shell and Prenflo technologies The Shell Coal Gasification Process (SCGP) uses an up-flow, entrained flow gasifier operating on a dry feed at slagging temperatures (Fig. 5.6). The up-flow arrangement allows separation of syngas and slag largely within the reactor itself. Typically there are four side-mounted burners located in the lower part of the reactor. The vessel containment uses a membrane wall, which has a demonstrated lifetime of over 15 years. Syngas cooling is via a cooled-gas recycle quench to bring any ash particles in the syngas through the ash solidification temperature

5.6 Shell SCGP gasifier (courtesy of Shell).

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range, followed by a convective steam-raising syngas cooler. A modified version replacing the syngas cooler with a partial water quench has been developed for CCS and hydrogen/chemical applications. The Shell process was first demonstrated in Hamburg in 1978, at a pilot plant built by a joint-venture between Shell and Koppers. The next demonstration plant was built in Deer Park, TX, and operated for part of its lifetime on petroleum coke (Mahagaokar and Hauser, 1994). The process has been operated by Nuon in its 250 MWe power plant in Buggenum, The Netherlands, continuously since 1994. A further 20 plants of similar size are in various stages of planning, construction and operation in China, mostly for chemical operations. A 300 MWe plant, designed by Koppers (now Uhde) under the name Prenflo, based on the results of the original Hamburg joint venture operates in Puertollano, Spain. The feedstock for the plant is a 40/60 petroleum coke/coal mix. E-Gas (ConocoPhillips) The E-Gas process uses a coal–water slurry feed into an up-flow, entrained flow gasifier (Fig. 5.7). A key identifying feature is the use of two-stage gasification. The second stage uses the heat in the syngas product from the first stage to devolatize and gasify the second-stage feed. This allows the syngas outlet to be cooler than the ash melting temperature, although the first stage is operating at slagging temperatures. This contributes to improved cold gas efficiency for the process. Further cooling is by means of a convective cooler. The reactor is refractory lined. The E-Gas technology now owned by ConocoPhillips was originally developed by Dow Chemical, which built a 550 t/d pilot plant in Plaquemine, LA, in 1983. This was followed by a 1600 t/d 165 MWe IGCC production unit on the same site, which operated on sub-bituminous coal from 1987 until 1995. These plants provided the basis for the Wabash River 250 MWe IGCC, which went on-stream in 1996 (Wabash River Energy Ltd, 2000). The Wabash plant has operated for many years on a 100% petroleum coke feed, and continues to do so. Gasifiers for liquid feeds There are four gasifiers for liquid feeds available: those offered by Shell (SGP), GE Energy, Lurgi (MPG) and Siemens. Texaco (now GE Energy) and Shell began development of processes for partial oxidation of liquid feeds in the 1940s and 1950s respectively, and these two processes have come to dominate the market. Lurgi had worked with solid feed gasifiers since the 1930s, but only entered the market for liquid feeds much later, in the 1990s, with a development originally focusing on handling the tars produced in its fixed bed dry bottom coal gasifier. The Siemens liquids process is also derived from its coal gasification process and

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5.7 ConocoPhillips E-Gas gasifier (courtesy of ConocoPhillips).

is in operation gasifying tars from a fixed-bed dry-bottom coal gasifier at Vresova, Czech Republic. Many key features of all four processes are similar. All use entrained-flow reactors. The burners are top-mounted in the down-flow, refractory-lined reactor vessels. Operating temperatures are similar (in the range 1250–1400 °C). When operating on liquid feed, all four processes produce a small amount of residual carbon, which is necessary to sequester the ash from the reactor. The important differences between the processes are in the details of burner design, in the method of syngas cooling, and in soot handling. For this reason the gasification of liquid hydrocarbons will only be described at a generic level. Figure 5.8 shows a typical liquid residue gasification unit with syngas cooler. The residue must be transferred to the POX unit at a temperature sufficient to ensure its pumpability. It is then fed to the reactor, typically with a plunger pump. Additional preheat is usually required to further reduce the viscosity to accommodate the requirements of the burner. Burner designs are proprietary, but all atomize the feed and mix it with the steam and oxygen that take part in the partial oxidation reaction. The reactor vessel is refractory-lined with a top-

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5.8 Generic liquid feed gasification with syngas cooler.

mounted burner. Typical operating temperatures are in the range 1250–1400 °C. Pressures lie between 30–80 bar and are selected according to the application. The hot gas leaving the reactor is used to generate high-pressure steam (up to 100 bar or more) in the fire tube syngas cooler. In the reactor the carbon conversion is typically about 99.5%. The unconverted carbon or soot is washed out of the gas in a two-stage water wash, which cools the gas further to ambient temperature and reduces the solid’s content of the scrubbed gas to less than 1 mg/Nm3. The run-down water, known variously as carbon slurry or black water, contains about 1% solids and may be processed in a number of ways. The carbon may be extracted and recycled with the feedstock to the POX reactor. With high ash feeds this approach may be limited because of the buildup of ash in the recycled material. Alternatively, the solids may be filtered out of the slurry. Different techniques are available for handling the vanadium-rich filter cake. In all cases the bulk of the water is recycled to the scrubber. Excess water is stripped to remove ammonia, HCN and H2S. The quench version of the process is shown in Fig. 5.9. The POX reactor and its feed system are as in the syngas cooler version. The hot gas is in this case quenched in a water bath below the reactor to about 220–230 °C. The steam-saturated quenched gas is scrubbed hot in a venturi scrubber to ensure adequate removal of solids. In this condition it is ideally prepared for shift conversion over a sour shift catalyst. If it is not intended to shift the gas for, say, hydrogen production, then the gas is further cooled generating medium- or low-pressure steam. Table 5.3 shows typical gas analyses of syngas generated by the processes described above.

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5.9 Generic liquid feed gasification with water quench.

Table 5.3 Typical raw gas analyses Process

GEE (quench)

CO2 (mol%) CO (mol%) H2 (mol%) CH4 ( mol%) N2 + Ar ( mol%) H2S + COS ( mol%)

20.7 41.7 37.1 0.1 0.4

Shell 2.2 65.2 26.0 <0.1 5.0 1.63

E-Gas 12.8 48.7 35.9 1.3 1.3

Liquids gasifier 2.30 52.27 43.80 0.30 0.25 1.08

Note: all analyses are given on dry basis. Values will vary according to feed quality, oxygen purity, operating pressure, etc. Data for GEE and E-Gas are on a sulfur-free basis

5.3.4 Process control Upon examination of the thermodynamics of the reactions in a gasification reactor, it becomes clear that the operating temperature is very sensitive to the oxygen-tocarbon (O/C) ratio of the reactants. Typically, 1% additional oxygen can raise the temperature by about 40 °C. Thus, maintaining the correct O/C ratio is key to good and safe operation. This is easier for hydrocarbons than for coal, where varying ash contents can play havoc even with a well-designed O/C control system. A major difficulty is that the conditions in the reactor do not allow reliable

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temperature measurement. The reactor temperature can only be determined indirectly by inference from the syngas analysis or the steam made in the wall of a membrane wall reactor. For liquids gasification this is usually done by reference to the methane content, which in some plants has been used as a trim variable on the ratio control. Often the ratio control is executed as a ‘lead-lag’ system to ensure that, during the transitions of a load change operation, the O/C ratio is always slightly lower than the set point and never higher.

5.4

Building blocks for complete systems

To gain a full picture of how gasification can be used in the context of a complete plant, it will be necessary to understand some of the gas treatment processes used. These include both chemical and physical absorption technologies for desulfurization and CO2 removal, CO shift for the conversion of carbon monoxide and steam to hydrogen and CO2, as well as COS hydrolysis and possibly the use of activated carbon for mercury removal. These topics are all handled in Section 5.4.2 below.

5.4.1 Air separation The air separation unit (ASU) is a standard, cryogenic unit as used by the industrial gas industry around the world. The oxygen quality for hydrogen or chemical applications is typically 99.5% O2; for power applications it is generally 95% O2. The decision whether to own and operate the oxygen plant or to buy oxygen over the fence is a project decision and there are plenty of successful examples for both business models.

5.4.2 Syngas treating In syngas, the sulfur is present as H2S and COS, rather than as SO2 and SO3 as would be the case after combustion. NOx is non-existent except in the form of NOx precursors such as ammonia and HCN. Where CO shift is applied, carbon dioxide may be present in concentrations of over 30%, rather than 12–15% as in a standard flue gas. Furthermore, the volume of syngas is only a fraction of that of flue gas and it is under a pressure of 25 bar or more. All this makes the removal of criteria pollutants to very low levels considerably easier than is the case with flue gas clean-up. Primary gas treatment, which includes particulate removal and a water wash to extract ammonia and chlorides from the gas, is generally considered to be an integral part of the gasification process. The measures taken vary slightly from process to process, and have been discussed in Section 5.3.2. Additional treatment always includes desulfurization, for which a number of different solutions are available. Other gas treatment requirements will be project-

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specific depending on application (hydrogen, power or chemicals) and regulatory requirements. For hydrogen and chemical applications, CO shift (converting CO and steam to hydrogen and CO2) and CO2 removal are required, the latter in general being integrated in one way or another with the removal of H2S in a consolidated acid gas removal (AGR) system. Mercury removal may need to be included as an additional treatment task. Acid gas (H2S and CO2) removal The desulfurization technologies applied to syngas on a commercial scale have a long history in the oil refining and natural gas industries. The primary technologies are based on the chemical or physical absorption of H2S in a suitable solvent. About 5% of the sulfur in the raw gas from a gasifier is present as COS so that, when deep desulfurization is required, attention must be given to its removal as well. Physical solvents have at least a partial capability to remove COS, whereas chemical solvents do not. In the latter case the COS is therefore hydrolyzed to H2S before the main desulfurization step. The H2S extracted is then processed to a useable/saleable product such as elemental sulfur (in a Claus plant) or sulfuric acid. CO2 removal is an integral part of chemicals production. For ammonia or hydrogen production the carbon capture rate must be 100%, since no carbon compounds can be contained in the product. For methanol or Fischer–Tropsch fuels the carbon capture rate is about 50%, and for SNG it lies between these values. Since carbon capture (or CO2 removal) is inherent to these processes, its cost is incorporated into the cost of the product. Thus, while the cost of CO2 compression, transport and storage is not yet paid for, the capture portion of a carbon capture and storage (CCS) application is already included. This has, of course, contributed to the positive economics of the Dakota–Weyburn project (Miller and Pouliot, 2008). This is not the case for power production. There is currently no obligation or incentive, either chemical, legal or economic to remove carbon from the power production process, whether combustion- or gasification-based. As long as this situation prevails, there is little likelihood of CCS being implemented on a large scale anywhere, no matter what technology is available. Studies have been made showing that, while gasification for conventional power has a capital cost disadvantage, when the cost of CO2 removal is added it is competitive with combustion systems (e.g. Booras, 2008). Certainly, the background in chemical applications places it in an advanced stage of the scale-up process. At the time of writing, a 524 MWe lignite-fired power plant with 65% CO2 capture is in planning in Kemper County, MS (Rush, 2010). Also in this case, the economics of carbon capture are supported by the sale of three million tonnes per year of CO2 to the oil industry. In the medium term, there is certainly further scope for cost reduction by recovering the CO2 under pressure, thus saving on the compression component in the overall CCS scheme as described in Section 5.7.1.

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Chemical solvent processes Chemical solvents are characterized by the acid gas component being chemically bonded to the solvent during the absorption step. A wide variety of amines have been used for acid gas removal. Today, methyl diethanolamine (MDEA) is probably the most widely used amine in syngas applications. The solvent circulation rate, which is the main determinant for both capital and operating cost, is approximately proportional to the quantity of acid gas removed. Figure 5.10 shows the flow sheet of a typical MDEA wash. The raw syngas is contacted in a wash column with lean MDEA solution, which absorbs the H2S and some of the CO2. MDEA is to some extent selective, in that the bonding of the amine with H2S takes places faster than with CO2, and advantage can be taken of this in the design if appropriate. The rich solution is preheated by heat exchange with the lean solution and enters the regenerator. Reboiling breaks the chemical bond, and the acid gas components discharged at the top of the regenerator are cooled to condense out the water, which is recycled. Physical solvent processes In a physical wash the acid gas is dissolved physically in the solvent. The solution loading is therefore largely dependent on Henry’s law and hence on the partial pressure of the acid gas component. The solvent circulation rate at any particular operating pressure is approximately proportional to the volume of raw gas to be processed.

5.10 Typical MDEA flowsheet.

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Selexol The Selexol process was originally developed by Allied Chemical Corporation and is now owned by UOP. It uses dimethyl ethers of polyethylene glycol (DMPEG). The typical operating temperature range is 5–40 °C. The ability to operate in this temperature range offers substantially reduced costs by eliminating the necessity for refrigeration. For a chemical application such as ammonia, the residual sulfur in the treated gas may be 1 ppmv H2S and 1 ppmv COS after the CO2 wash (Sharp, 2002). This is, however, not an issue in power applications where the sulfur slip is less critical. Selexol has a number of reference plants including the original Cool Water IGCC demonstration unit, the 550 MWe Sarlux IGCC facility in Italy and, most recently, the 630 MWe Edwardsport plant, which is under construction. The Selexol flowsheet in Fig. 5.11 exhibits the typical characteristics of most physical absorption systems. The intermediate flash allows co-absorbed syngas components (H2 and CO) to be recovered and recompressed back into the main stream. Provision is made for H2S concentration in the acid gas to make it suitable for processing in a Claus plant. Separate CO2 recovery using staged flashing techniques is applied so that part of the CO2 can be recovered at a moderately elevated pressure (MP CO2).

5.11 Typical two-stage Selexol flowsheet (courtesy of UOP LLC, a Honeywell Company).

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Rectisol The Rectisol process, which uses cold methanol as a solvent, was originally developed to provide a treatment for gas from the Lurgi moving-bed gasifier which, in addition to H2S and CO2, contains hydrocarbons, ammonia, hydrogen cyanide and other impurities. In the typical operating range of −30 to −60 °C the Henry’s law absorption coefficients of methanol are extremely high, and the process can achieve gas purities unmatched by other processes. This has made it a standard solution in chemical applications such as manufacture of ammonia and methanol, and methanation, where the synthesis catalysts require sulfur removal to less than 0.1 ppmv. This performance has, however, a price in that the refrigeration duty required for operation at these temperatures involves considerable capital and operating expense. The flowsheet is basically similar to that of Selexol, but the need to economize on refrigeration introduces more heat exchange equipment. The process is used in over 100 plants, mostly in chemical applications. It is the system used at the North Dakota SNG plant mentioned above, which captures CO2 for enhanced oil recovery at Weyburn in Canada. Rectisol is also used in Shell’s IGCC and hydrogen facility at the Pernis refinery, where about one million tonnes of CO2 is captured per year. About one-third of this is piped through an 85 km pipeline for use in greenhouses (VDI Nachrichten, 2009). CO shift The so-called CO shift conversion process has an important place in any hydrogen manufacturing or pre-combustion carbon capture scheme. It converts carbon monoxide in the syngas with steam to form carbon dioxide and hydrogen according to the reaction: ← CO + H CO + H2O → 2 2

−41 MJ/kmol.

[Reaction 5.5]

In chemical applications it is used to adjust the H2/CO ratio of the syngas to suit the requirements of the synthesis. In power applications it may be used as part of a carbon capture strategy. As can be seen, one mole of hydrogen is produced from every mole of CO. The reaction is largely independent of pressure. The equilibrium for hydrogen production is favored by low temperature. CO shift conversion in the gasification environment exhibits a number of differences from the steam–methane reformer (SMR) situation, more familiar from the natural gas-based hydrogen plant. The most important difference is the amount of CO to be converted, which can vary from about 45% downstream of an asphalt gasifier to over 60% at the exit of a Shell gasifier processing coal or petroleum coke. Managing the corresponding exotherm places very different demands from the 16% or so at the outlet of an SMR. A second difference is in many cases the presence of sulfur in the raw gas.

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The CO shift reaction will operate with a variety of catalysts between 180 and 500 °C. The types of catalyst are distinguished by their temperature range of operation and the quality (sulfur content) of the syngas to be treated. Raw gas shift In cases where the gasifier has been designed with a water quench, this provides the steam loading in the raw syngas required for the shift reaction. The gas still contains sulfur, so a cobalt–molybdenum catalyst, variously described as a sour shift or dirty shift catalyst, is used. The catalyst requires typically 500–700 ppmv sulfur in the feed gas to maintain it in the active sulfided state. The shift reaction is exothermic so that the gas temperature rises as it passes through the catalyst bed and CO is shifted to CO2. If a high degree of shift is required, it will be necessary to have multiple beds with intermediate cooling. Figure 5.12 shows a typical three-stage shift as might be found in an ammonia plant with a residual CO concentration of about 1.6 mol% or 0.8 mol% after the second and third beds respectively. An important side-effect of the raw gas shift catalyst is its ability to handle a number of other impurities characteristic of gasification. COS and other organic sulfur compounds are largely converted to H2S, which eases the task of the downstream AGR. HCN and any unsaturated hydrocarbons are hydrogenated. Clean gas shift In some plants where a syngas cooler is used, it may be more convenient to desulfurize the gas upstream of the shift, providing the opportunity to remove the CO2 at a location where selective operation of the AGR is not necessary. In this case a ‘conventional’ (high-temperature), iron oxide-based catalyst promoted typically with chromium and, more recently, with copper is used. The operating range of these catalysts is between 420 and 500 °C. HT shift catalyst is tolerant of only a small amount of sulfur, but is not as sensitive as the (low-temperature) copper catalysts used in other applications.

5.12 Three-stage CO shift system.

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COS hydrolysis In all syngases produced by gasification, only about 95% of the sulfur is present as H2S, the remainder being COS. While some washes such as Rectisol can remove the COS along with the H2S, others, particularly amine washes, require the COS to be converted selectively to H2S if the sulfur is to be substantially removed. Where the COS is not converted on a sour shift catalyst, this is best achieved by catalytic COS hydrolysis, according to the reaction: ← H S + CO COS + H2O → 2 2

−30 MJ/kmol

[Reaction 5.6]

Commercially, this reaction takes place over a catalyst at a temperature in the range 150–200 °C. Various catalysts are available, including pure activated alumina, titanium oxide, or a promoted chromium oxide–alumina. Lower temperatures favor the hydrolysis equilibrium. Depending on process conditions the residual COS can be reduced to the range of 5–30 ml/Nm3. Mercury removal The necessity for mercury removal will be determined by the quality of the feedstock. It is more likely to be needed with coal that with petroleum coke. The technique used is to pass the syngas over a bed of sulfur-impregnated activated carbon at ambient temperature, typically just upstream of the desulfurization unit. In natural gas service 99.99% mercury removal can be achieved this way. The experience at Eastman Chemicals’ Kingsport, TN, plant shows a mercury removal efficiency of 90–95% (US DOE, 2002).

5.5

Hydrogen and power plant as an example for a complete system

5.5.1 Basic considerations In a refinery environment, hydrogen is always likely to be one potential gasification product stream whatever the overall strategy for bottom-of-the-barrel processing. The volumes of residue in a refinery, however, are generally larger than those required to meet any internal hydrogen demand, so that a second product may be required. The selection of an additional product will depend on local markets so, for the following example, power generation by firing syngas in a gas turbine/combined cycle, which may be considered at any location, has been selected. Nonetheless, one must be conscious of the fact that of the possible products from a gasification plant, power is likely to be the least remunerative. When looking at the technology selection for hydrogen and power production from an asphalt feed, the first consideration needs to be given to the proportion of syngas to be used for hydrogen. If this proportion is less than 15–20% of the total

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syngas production then it may well be possible to draw off crude hydrogen from the syngas through a permeable polymer membrane and purify it in a pressure swing adsorption (PSA) unit. This would certainly be the lowest-cost approach. Much above this proportion of the final product as hydrogen will, however, require a CO shift unit to convert carbon monoxide into additional hydrogen. This will be the case in this example in which two-thirds of the gas is processed to hydrogen and one-third used as gas turbine fuel. A second fundamental issue is deciding on the hydrogen purity requirement. Many hydrocracking processes will accept a 98% H2 purity, even if at a small economic penalty. This grade of hydrogen can be produced by the traditional shift, CO2 removal and methanation route. A number of refinery processes do, however, require purities of 99% H2 or higher, in which case the final purification step will have to be PSA. The higher purity is, however, achieved at the cost of a lower hydrogen yield (about 85–90% instead of 98%) and the production of a relatively large quantity of low-pressure, low-BTU fuel gas which may or may not be accommodated in the refinery fuel gas balance. It is important in this context to review the hydrogen purity specification carefully with the hydrocracker requirements, since all too often a purity of >99.5% is specified on the basis of the economics of the conversion unit alone, or on the assumption that hydrogen will be generated from a steam reformer which can accommodate the PSA tail gas internally, without reference to the economics of the overall configuration. Based on these applications there will be no need to go to pressures over about 35 bar, though for large plants it may be advantageous to go for higher pressures, so as to reduce equipment sizing. The overall configuration, which includes CO shift and a PSA, is shown in Fig. 5.13.

5.13 Block flow diagram of hydrogen and power plant.

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5.5.2 Process description A conventional cryogenic air separation unit (ASU) generates the oxygen required for the gasifier. In addition, it supplies diluent nitrogen for the gas turbine in the combined-cycle unit. The air requirement for the ASU is provided by a dedicated air compressor, even though technically part of the air could be supplied as extraction air from the gas turbine. This configuration has the advantage of avoiding problems arising out of excessive process integration. The feedstock is gasified in the partial oxidation reactor with oxygen to produce raw syngas at a temperature of about 1300 °C and a pressure of typically about 60 bar. The gas produced is a mixture of hydrogen and carbon monoxide which, however, also contains CO2, H2S and COS together with some free carbon and the ash from the feedstock. The hot gas is cooled in the syngas cooler by generating saturated high-pressure steam, which can be superheated in the heat recovery steam generator (HRSG) of the combined-cycle unit. Particulates are removed from the gas using a water wash prior to desulfurization. Desulfurization is effected using the Rectisol process, in which cold methanol is used as the solvent. The one-third of the gas to be used as gas turbine fuel can be extracted at this stage and directed to the combined-cycle power plant. The acid gas is then processed in an oxygen-blown Claus unit to make elemental sulfur. The remaining two-thirds of the gas is processed in a two-stage CO shift unit in which most of the CO is converted to H2 and CO2. CO2 removal from the gas is then effected in a second-stage Rectisol unit. The inclusion of this processing step may be surprising to those familiar with hydrogen production using steam–methane reforming (SMR) technology since, in that scenario, PSA acts to integrate CO2 removal with H2 purification. The C/H ratio of a typical gasification feedstock is much higher than that of methane, so that the proportion of CO2 to hydrogen in the shifted gas is correspondingly higher than in the SMR case. Typically it can be of the order of 35–40%. Leaving the CO2 removal task to the PSA unit would have two effects. Firstly, sizing the adsorbers for this amount of CO2 would reduce the hydrogen yield to unacceptably low levels. Secondly, the large amount of CO2 in the PSA tail gas would reduce its heating value to the extent that it would only have borderline value as a fuel. Final hydrogen purification takes place in a PSA unit. The tail gas is used as fuel in the refinery or in the power plant, e.g. for duct firing in the HRSG. The Combined-Cycle Unit (CCU) block described in Figure 5.13 covers the typical scope of a natural gas combined cycle complete with balance of plant. The principal differences lie in the use of syngas as a fuel and the addition of steam raised in the gas generation units to supplement the steam cycle. Gas turbines burning syngas use diffusion flame combustors. These combustors cannot achieve the same NOx emission reduction as the pre-mix burners used with

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natural gas. Syngas dilution is therefore required to reduce the NOx emissions from the gas turbine. The dilution medium may be nitrogen or steam, or a combination of the two. Steam is generally added by saturation using low-level heat to provide the necessary hot water. In some cases it is added by direct injection. In the example flowsheet, nitrogen has been applied.

5.6

Advantages and limitations

An important advantage of gasification is its ability to handle a flexible range of ‘disadvantaged’ and therefore low-price fuels and produce valuable commodities such as hydrogen, ammonia, and methanol from them. Typical of such fuels in the hydrocarbons industry are petroleum coke and asphalt. Gasification facilities are, however, capital intensive, so that specific project economics need to be examined carefully in each case.

5.7

Future trends

5.7.1 Advanced gasification systems Gasification is generally a mature process, so that current development emphasis is on incremental improvements in all the component blocks, rather than a radical rethink of overall systems. The focus for solids gasifiers is on cost reduction in coal-based gasification applications, but most of such improvements will be transferrable to petroleum coke gasification. Development in liquids gasifiers has concentrated on the ability to handle increasingly heavier feedstocks and to scale-up to larger sizes. All the major gasification technologies continue to undergo incremental changes. One step-change is being pursued by Pratt & Witney Rocketdyne with their Compact Gasifier. The use of multiple fuel injectors and a plug flow reactor is expected to reduce the gasifier size by an order of magnitude. An 18 t/d test unit was taken on stream during December 2009 (Darby, 2010). The lock hopper system for dry feed gasifiers, which relies on gravity flow to move the pulverized feedstock from the uppermost, atmospheric bunker through the lock hopper into the feed vessel, requires a support structure every bit as tall as the gasifier itself, so there is great interest in the development of a ‘solids pump’, which could reduce the cost of the feed system. One such has been developed by Stamet, now owned by GE Energy. GE is currently building a test facility with the University of Wyoming (High Plains Gasification Advanced Technology Center) to test and demonstrate the use of the Stamet Posimetric pump in connection with a GE gasifier (Zuiker, 2010). Once mature, this technology will allow GE to use dry feed for low-rank coals. A similar development is included in the Rocketdyne development package mentioned above.

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Research into gasifiers which produce methane directly for SNG production has varied with changing perceptions of the natural gas market. In the 1980s there were efforts in Germany (Rheinbraun) and the USA (Exxon) in this direction. Currently this line of investigation is being pursued by GreatPoint Energy, with its low-temperature (600–700 °C) catalytic Bluegas™ development. A test plant was commissioned in 2009 (GreatPoint Energy, 2010).

5.7.2 Auxiliary technologies An important development effort in air separation technology is based on the use of the transport of oxygen ions across a dense ceramic membrane at temperatures in the 800–900 °C range. This has the potential for substantial reduction of both capital and operating costs. Initial testing of 0.5 t/d O2 modules started in 2006. The first 100 t/d O2 pilot plant is currently under preparation (Steele et al., 2010). In the gas treatment area, the thermodynamic benefit of operating at elevated temperatures has led many to investigate the possibilities of ‘hot (or warm) gas clean-up’. Achievements in this direction over the last twenty years or so have been frustratingly slow and many attempts have been abandoned. RTI, however, appears to have made a breakthrough with its high-temperature syngas desulfurization process (HTDP). RTI uses a continuous fluid bed transport reactor system to adsorb both H2S and COS onto zinc oxide at temperatures in the range of 260–540 °C. A 0.3 MWe pilot plant has operated successfully for over 3000 hours in a commercial setting, and a 50 MWe demonstration facility is now in planning. In the CCS configuration the full thermodynamic benefits of warm gas clean-up can only be realized if the CO2 and other contaminants can also be removed at high temperature. RTI have announced a program to address this issue also (Gupta et al., 2010). An alternative development, which will leave CO2 at the relatively high gasification pressure, is the use of hydrogen membranes. One such is that developed by Eltron Research & Development, which is planning a 4–10 t/d precommercial module to be installed in a commercial setting (Mundschau, 2008, US DOE NETL, 2010a). Such technologies will substantially reduce the cost of CO2 compression for CCS.

5.8

Sources of further information and advice

Gasification Technologies Council website, http://www.gasification.org, which has an extensive library and database of gasification plants. Higman, C. GasLit.mdb database, which can be downloaded from http://www.gasification. higman.de. Contains over 1000 literature references to articles and papers on gasification. Higman, C. and van der Burgt, M., Gasification, Gulf Professional Publishers, Amsterdam, 2003.

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Kohl, A. and Nielsen, R., Gas Purification. Gulf Professional Publishers, Amsterdam, 5th Edition, 1997. Technical University of Freiberg (Saxony, Germany). Short courses in gasification technology are offered. Details are available at http://tu-freiberg.de/fakult4/iec/ schulung.en.html. US Department Of Energy, National Energy Technology Laboratory website, http://www. netl.doe.gov/technologies/coalpower, which contains a wide range of reports, studies and other documents on gasification and its application to CO2 capture and storage. A recent addition to the DOE website is their Gasifipedia section, http://www.netl.doe. gov/technologies/coalpower/gasification/gasifipedia/TOC.html.

5.9

References

Barkley N E; ‘Petroleum coke gasification based ammonia plant’, AIChE Ammonia Safety Conference, Vancouver, 2006. Booras G, ‘Economic assessment of advanced coal-based power plants with CO2 capture’. Paper presented at MIT Carbon Sequestration Forum IX: Advancing CO2 Capture, Cambridge, MA, 16 September 2008. Darby A, ‘Compact Gasification: development and test status’. Gasification Technologies Conference, Washington DC, October 2010 GreatPoint Energy, http://www.greatpointenergy.com/mayflowercleanenergycenter.php, November 2010. Gupta R, Jamal A, Turk B, and Lesemann M, ‘Scale up and commercialization of warm syngas cleanup technology with carbon capture and storage’. Gasification Technologies Conference, Washington, DC, October 2010. Higman C and van der Burgt M, Gasification, Gulf Professional Publishers, Amsterdam, 2003. Hydrogen Energy Plc, http://www.hydrogenenergycalifornia.com, 17 November 2010. Koopman E, Regenbogen R W and Zuideveld P L, ‘Experience with the Shell Coal Gasification Process’. VGB Conference Buggenum IGCC Demonstration Plant, November 1993. Laege J and Pontow B, 30 ‘Years of successful operation’. European Gasification Conference, Noordwijk, The Netherlands, April 2002. Mahagaokar U, and Hauser N, ‘Gasification of petroleum coke in the Shell Coal Gasification Process’. IMP Gasification Colloquium, Mexico City, February 1994. Miller C and Pouliot S, ‘Dakota Gasification Company: an international energy venture’. Pittsburgh Gasification Conference, Pittsburgh, 2008. Mundschau M V, ‘Hydrogen separation using dense composite membranes, part 1: fundamentals’. In Inorganic Membranes for Energy and Environmental Applications. Ed. Bose, A.C. Springer, New York, 2008. Posthuma S A, Vlaswinkel P L, and Zuideveld, P.L. ‘Shell gasifiers in operation’. IChemE Conference Gasification Technology in Practice, Milan, 1997. Rettger P, Arnold J, Brandenburg B and Felch C, ‘The Long Lake integrated upgrading project: status report and discussion of soot processing’. Paper presented at Gasification Technologies Conference, Washington DC, October 2006. Rush R, ‘Overview of the Kemper County IGCC project using transport integrated gasification (“TRIG”)’. Gasification Technologies Conference, Washington DC, October 2010.

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Sharp C R, Kubek D J, Kuper D E, Clark M E, Didio M et al., ‘Recent Selexol and Membrane/PSA Operating Experiences with Gasification for Power and Hydrogen’. Gasification Technologies Conference, San Francisco, October 2002. Steele R, Armstrong P and Bose A, ‘Ion transport membrane (ITM) technology for lower cost oxygen production’. Gasification Technologies Conference, Washington, DC, November 2010. US Department of Energy (DOE), The Cost of Mercury Removal in an IGCC Plant – Final Report. September 2002. US Department of Energy National Energy Technology Laboratory (DOE NETL), Industrial Carbon Capture Project Selections. http://www.fossil.energy.gov/recovery/ projects/iccs_projects_0907101.pdf, September 2010a. US Department of Energy National Energy Technology Laboratory (DOE NETL) Gasification database, http://www.netl.doe.gov/technologies/coalpower/gasification/ worlddatabase/index.html, November, 2010b. VDI Nachrichten, Der Dünger kommt mit Druck aus der Pipeline, VDI Nachrichten, 6 February 2009, 6 3 (2009). Wabash River Energy Ltd. Wabash River coal gasification re-powering project: final technical report, US Department of Energy, August 2000. Zuiker J, ‘GE Gasification Technology Update’. Gasification Technologies Conference, Washington DC, November 2010.

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